EXPERIMENTAL RESULTS AND COMPUTER SIMULATIONS FOR POST-

COMBUSTION CARBON DIOXIDE REMOVAL USING LIMESTONE

THESIS

Presented in Partial Fulfillment of the Requirements for the Degree Master of Science in the Graduate School of The Ohio State University

By

William Kane Wang, B.S. ChBE

Graduate Program in Chemical and Biomolecular Engineering

The Ohio State University

2009

Thesis Committee:

Liang-Shih Fan, Advisor

Jacques Zakin

ABSTRACT

Amid growing concerns of global climate change, governmental entities and industry throughout the world are performing research on the capture and sequestration of carbon dioxide (CO2). Coal accounts for nearly half of the United States electricity generation and 40% of the world’s electricity generation. This directly translates into 33% of the

United States CO2 emissions and 40% of the world’s CO2 emissions produced from coal- combustion power plants for electricity generation. Not surprisingly, a significant emphasis has been placed on capturing CO2 produced from coal-combustion power plants. To date, though, no large-scale power plant employs the use of carbon capture technologies.

Under the leadership of Professor Liang-Shih Fan and his research group at The Ohio

State University, a process that employs a solid sorbent to reactively remove CO2 from coal combustion flue gas has been developed. The Carbonation-Calcination Reaction

(CCR) Process relies on a carbonation reaction with a metal oxide to remove the CO2 and a subsequent calcination reaction to produce a pure stream of CO2 while regenerating the metal oxide. An additional benefit to the CCR process is the ability of the sorbent to simultaneously remove sulfur dioxide (SO2) present in the flue gas stream.

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Upon successful bench-scale experiments, a 20 pound per hour coal combustion facility was erected to demonstrate the CCR Process. With the ability to capture greater than 90%

CO2 and greater than 99% SO2 on a once-through basis with a commercially available calcium-based sorbent, the facility was re-renovated for multicylic investigations.

Consistent CO2 removals over multiple cycles were obtained using an intermediate hydration reaction for reactivation of the calcium oxide sorbent. The hydration reaction reverses and eliminates any effect of sintering that occurs during the calcination reaction.

Due to its inherently small particle size, along with an increase in surface area and pore volume upon spontaneous dehydration, calcium hydroxide is able to maintain constant

CO2 removals.

Computer simulations integrating the CCR Process into a power plant shows a high level of compatibility due to its high-temperature operation. Energy penalties between 15% and 20% were obtained, with compression of CO2 while producing a pure, dry stream of

CO2 ready for sequestration.

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Dedicated to my family-Mother, Father, and three sisters

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ACKNOWLEDGEMENTS

I would like to gratefully acknowledge my advisor, Professor Liang-Shih Fan, for providing the opportunity to obtain my graduate degree and perform research in his group. I have had the ability to learn a considerable amount in several fields that extend beyond research and academics. I would also like to acknowledge Professor Jacques

Zakin for agreeing to be on my defense committee.

The research performed could not have been achieved without several personnel from industry. I would like to thank Dr. Robert Statnick for providing input, sharing his previous experiences, providing continuous support, even in times of great stress, and his patience. I would also like to thank Dr. Mahesh Iyer for his initial role as mentor, support, and conversations, both during his time at Ohio State and beyond. I would like to thank

Bob Brown and the Ohio Coal Development Office for their continued financial support of the research project. The members of the Industrial Review Committee, who have provided important suggestions in areas of experimental design, operational procedures, industrial expectations, safety precautions, and troubleshooting, have my deepest appreciation. They include Jeff Gerken from American Electric Power (AEP), Bartev

Sakadjian from Babcock & Wilcox, Lew Benson, Dave McKinney, and James Derby of

Carmeuse Lime & Stone, Dan Connell, Steve Winberg, and Dick Winschel of CONSOL

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Energy, John Bloemer and Frank Carchedi of Duke Energy, and Mark Golightley of First

Energy.

I would also like to thank the entire Fan Research Group, which has undergone significant personnel changes during my five years as a graduate student. The members of the Carbonation Calcination Reaction (CCR) group consisting of Danny Wong,

Shwetha Ramkumar, Dr. Songgeng Li, Siddharth Gumuluru, and Sun Zhenchao have significantly advanced the progress of the research project. With my initial research focusing on chemical looping combustion, I would like to thank Dr. Luis Vargas, who provided not only invaluable advice but also comic relief, Fanxing Li, Deepak Sridhar,

Ray Kim, Andrew Tong, and Fu-Chen Yu. Transitioning into fluidization research, I would like to thank Dr. Bing Du, Dr. Zhe Cui, Dr. Raymond Lau, who taught me how to perform experiments while having fun, Dr. Warsito, Dr. Qussai Marashdeh, Orin

Hemminger, Fei Wang, and Zhao Yu. Last but not least, I would like to thank Dr. Alissa

Park for her guidance, compassion, and humor. I would also like to extend thanks to all the undergraduate help who didn’t mind the manual labor and not-so-clean conditions.

They are Yao Wang, Alex Brown, Zack Patterson, Zack Yoscovits, and Joe Braucher.

Without the administrative staff in both the Chemical and Biomolecular Engineering

Department and The Ohio State University, the research project would have been far too difficult to handle. I would like to thank Mike Davis, Geoff Hulse, and Dave Jones for maintaining the computer labs and network and David Cade and Carl Scott for maintaining not only Koffolt Laboratory, but also the laboratory at West Campus. Paul

Green and Leigh Evrard deserve thanks for their expertise in machining, fabrication, and installation of equipment. Without them, parts of the reactor would still be supported by

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bricks and wood. A special thanks must be given to Susan Tesfai and Lynn Flanagan, the two administrative workhorses of the research project. Their ability to keep the finances of the research project under control, incredible work hours, and ability to successfully obtain authorization for all the “odd” purchases that were made truly made the project run efficiently. The Ohio State Research Foundation also deserves a special thanks for authorizing all purchases in a timely fashion.

I would like to thank my entire family-my parents, Kang-bo Wang and Su-Huei Wang, sisters, Jenny, Judy, and Peggy for their unending support and love. I would also like to give a special recognition to my two wonderful cats, Jack Junior (JJ) and Remi, who have provided me with hours of entertainment and joy.

I would also like to thank the Capital Area Humane Society, as well as the Columbus

Ultimate Disc Community, for providing an enjoyable atmosphere for relaxation. Of special note is Erin Schran for providing Darby, Amy Sminchak for providing Flash,

Scott Miller for providing Clevis, Bob for providing Alfa and Bioko, and the entire

Ultimate dog community and their owners.

I would also like to extend my deepest thanks to Megan Covitz, Bredt Covitz, and their two wonderful cats, Jack and Eva, for providing me with countless meals, their incredible hospitality, and friendship. To Kelly Zilli, her three beautiful greyhounds, Vinnie, Guido, and Verdi, and her two adorable cats, Ezekiel and Elliott, I am forever indebted. The numerous dog walks, runs, randomness, and friendship have truly made my time in graduate school easier and more enjoyable.

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VITA

September 18, 1981…………………………………………Born – St. Louis, Missouri

June, 2004…………………………………………………...B.S. Chemical and Biomolecular Engineering, University of Illinois at Urbana-Champaign

2004-2009…………………………………………………..Graduate Research Associate The Ohio State University

PUBLICATIONS

Gupta, Puneet; Wang, William; Fan, Liang-Shih. Synthesis of High-Surface-Area SiC Through a Modified Sol-Gel Route: Control of the Pore Structure. Ind. Eng. Chem. Res. 2004, 43, 4732- 4739.

Wang, William; Li, Songgeng; Ramkumar, Shwetha; Wong, Danny; Iyer, Mahesh; Fan, L-S; Statnick, Robert M. Demonstration of Multi-Pollutant Capture Including CO2 and SO2 from Coal Combustion. Proceedings-7th Carbon Capture & Sequestration Conference, 2008.

Wang, William; Li, Songgeng; Ramkumar, Shwetha; Wong, Danny; Iyer, Mahesh; Fan, L-S; Statnick, Robert M. Results from 20 pound per hour CO2 Capture Facility. Proceedings- Pittsburgh Coal Conference, 2008.

Li, Songgeng; Wang, William; Ramkumar; Shwetha; Wong, Danny; Iyer, Mahesh; Fan, L-S; Statnick, Robert M. Demonstration of Carbon Dioxide Capture from Coal-Fired Flue Gas Using Calcium-based Sorbents. Proceedings-AICHE Annual Meeting, 2008.

Wang, William; Ramkumar, Shwetha; Li, Songgeng; Wong, Danny; Iyer, Mahesh; Sakadjian, Bartev; Statnick, Robert; Fan, L.-S. Sub-Pilot Demonstration of the Carbonation-Calcination Reaction (CCR) Process: High-Temperature CO2 and Sulfur Capture from Coal Fired Power Plants. Ind. Eng. Chem. Res. In Press.

FIELDS OF STUDY

Major Field: Chemical Engineering

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TABLE OF CONTENTS

Abstract ...... ii

Dedication ...... iv

Acknowledgements ...... v

Vita ...... viii

List of Tables ...... xi

List of Figures ...... xii

Chapter 1 Introduction ...... 1

Chapter 2 Review of Solid Sorbents for CO2 Removal ...... 7 2.1 Introduction ...... 7 2.2 Lithium Orthosilicate – Li4SiO4 ...... 7 2.3 Sodium Carbonate - Na2CO3 ...... 8 2.4 Calcium Oxide – CaO ...... 9 2.4.1 Instituto Nacional del Carbón - Juan Carlos Abanades ...... 9 2.4.2 CANMET and University of British Columbia ...... 11

Chapter 3 Calcination of Limestone ...... 14 3.1 Introduction ...... 14 3.2 Mechanism ...... 14 3.3 Factors Affecting Calcium Carbonate Calcination ...... 15 3.3.1 Particle Size Distribution ...... 15 3.3.2 Temperature Profile ...... 16 3.3.3 Calcination Atmosphere ...... 16

Chapter 4 Simultaneous CO2 and SO2 Removal Over Multiple Cycles ...... 19 4.1 Introduction ...... 19 4.2 Thermodynamics and Reactions ...... 20 4.2.1 Hydration of Calcium Hydroxide Reversible Reaction ...... 21 4.2.2 Carbonation of Calcium Oxide Reversible Reaction ...... 23

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4.2.3 Sulfation of Calcium Oxide ...... 25 4.3 Experimental ...... 27 4.3.1 Chemicals, Sorbents, and Gases ...... 27 4.3.2 Facility Design and Set-up ...... 30 4.4 Results and Discussion ...... 36 4.4.1 Initial Testing ...... 37 4.4.2 Calcium Oxide Carbonation-Calcination Cycles ...... 38 4.4.3 Calcium Hydroxide Carbonation Calcination Cycles ...... 41 4.5 Conclusions ...... 44 4.6 Recommendations ...... 45

Chapter 5 Commercial Considerations for the CCR Process ...... 47 5.1 Introduction ...... 47 5.2 Overall CCR Process Description ...... 49 5.2.1 Existing and Future Regulations ...... 51 5.2.2 Carbon Dioxide ...... 51 5.2.3 Sulfur Oxides ...... 52 5.2.4 Particulate Matter ...... 53 5.3 Calciner Considerations ...... 54 5.3.1 Energy Consumption ...... 55 5.3.2 Heat Integration ...... 56 5.3.3 Calciner Operating Costs ...... 63 5.3.4 Calcination Product-Calciner Design ...... 65 5.3.5 Calciner Operating Temperature ...... 67 5.4 Carbonator/Sulfator-CO2/SO2/H2O Management ...... 70 5.4.1 One Reaction Vessel ...... 71 5.4.2 Two Reaction Vessels ...... 74 5.4.3 Three Reaction Vessels ...... 76 5.5 Purge/Recycle and PCDs ...... 79 5.5.1 Purge/Recycle Physical Location ...... 79 5.5.2 Purge/Recycle Theoretical Location ...... 80 5.5.3 Particulate Capture Devices ...... 82 5.5.4 Final Considerations ...... 83 5.6 Aspen Simulations ...... 84 5.6.1 Overall Modelling Parameters ...... 84 5.6.2 Assumptions and Operating Conditions ...... 89 5.6.3 Coal-fired Power Plant without Carbon Capture ...... 90 5.6.4 CCR Process Integration with an Indirect-fired Calciner ...... 92 5.6.4 CCR Process Integration with a Direct-fired Calciner ...... 101 5.7 Conclusions ...... 109 5.8 Recommendations ...... 109

Chapter 6 Conclusions and Recommendations ...... 111

References ...... 114

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LIST OF TABLES Table Page

4.1 Proximate Analysis of as-received stoker-grade coal………………………….. 28

4.2 Ultimate Analysis of as-received stoker-grade coal…………………………… 28

4.3 Typical composition of Natural Gas…………………………………………… 28

4.4 Graymont Calcium Hydroxide and Calcium Oxide composition……………… 29

4.5 Theoretical flue gas compositions derived from coal combustion…………….. 30

4.6 Typical flue gas concentration generated from sub-pilot scale facility………... 30

5.1 Energy requirements, calciner capacity, and fuel used for various commercial calciners…………………………………………………………... 56

5.2 Parameters used to evaluate energy required for heating solids……………….. 58

5.3 Average cost and physical properties of chemicals used in calciner……………63

5.4 Properties of Pittsburgh #8 Coal………………………………………………...85

5.5 Chemical species identified for simulating CCR Process using Aspen…………86

5.6 Composition of Natural Gas……………………………………………………..86

5.7 List of Aspen Plus unit operations and reference for non-modeled units……….88

5.8 Flue gas components from Aspen simulation and their concentration………….90

5.9 Summary of Aspen simulation results………………………………………….109

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LIST OF FIGURES Figure Page 3.1 Thermal and mass transport phenomena governing calcium carbonate dissociation…………………….…………………………….14

4.1 Equilibrium pressure versus temperature diagram for calcium hydroxide………21

4.2 Equilibrium Pressure of CO2 versus Temperature for Calcium Oxide…………..25

4.3 Equilibrium Partial Pressure of SO2 versus Temperature for Calcium Sulfate….27

4.4 Process flow diagram of sub-pilot scale facility…………………………………35

4.5 Image of 20 pound per hour coal combustion sub-pilot scale facility…………...36

4.6 Initial set-up of calciner feed integration………………………………………...38

4.7 Relationship between Ca(OH)2-Cycle 1 and CaO-Cycle 2 versus CO2………....40

4.8 Average CO2 removals versus Cycle Number for Calcium Hydroxide…………43

4.9 Typical temperature profile exhibited in sub-pilot scale facility……………….. 44

5.1 General Process Flow Diagram for CCR Process……………………..………...50

5.2 Energy required for Calciner, which includes heating solids and calcining……..60

5.3 Total energy required for Calciner under realistic conditions…………………...61

5.4 Operating cost of Calciner as a function of Purge……………………………….64

5.5 Process Flow Diagram of Calciner using steam as gas diluent………………….69

5.6 Image of Riser section of 20 pph facility………………………………………...73

5.7 PFD of the CCR Process with a Carbonator/Sulfator and Dehydrator…………..75

5.8 SO2 removal on a once through CaSO4 loop…………………………………….77

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5.9 Recycle loop for CaO/CaSO4 for upstream SO2 removal………………………..78

5.10 Reference coal-fired power plant with no CO2 control and 507.58 MWe net…...91

5.11 Integration Option 1-Indirect fired calciner in one reaction vessel……………...95

5.12 Integration Option 2-Indirect fired Calciner with 2 vessels……………………...97

5.13 Integration Option 3-Indirect fired calciner with upstream SO2 removal………100

5.14 Integration Option 4-One reaction vessel using Direct fired Calciner…………104

5.15 Option 5-Two reaction vessel with Direct-fired Calciner and Dehydrator…….106

5.16 Integration Option 6-Direct fired Calciner with upstream SO2 removal……….108

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CHAPTER 1

INTRODUCTION

Although much rhetoric surrounds the subject of global warming, and more generally global climate change, increasing amounts of evidence continue to support the reality that global climate change is occurring. The Intergovernmental Panel on Climate Change

(IPCC), a scientific panel created by the World Meteorological Organization (WMO) and the United Nations Environmental Programme, has concluded that “Warming of the climate system is unequivocal, as is now evident from observations of increases in global average air and ocean temperatures, widespread melting of snow and ice and rising global average seal level” (IPCC, 2007). In slightly over a century, both marine air temperatures and sea surface air temperatures have increased between 0.4 °C and 0.8 °C (Sheppard and Socolow, 2007).

Global warming, which is an average increase in the temperature of the atmosphere near the Earth’s surface and in the troposphere, is mainly attributed to the increasing concentration of greenhouse gases, which are gases that limit the amount of heat escaping into space, emitted by humans (U.S. EPA, 2008). Although no standard definition exists for global climate change, it is always broader in scope and typically refers to any measurable alterations to the climate, not just a temperature increase, caused by either natural or human activity (IPCC, 2007).

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The observed and predicted environmental impacts associated with global climate change are deleterious. The dying of species, extreme weather conditions, and alterations in the production of food crops have all been linked to global climate change. For example, approximately 26,000 of the documented 29,000 (~90%) shifts in biological systems are attributed to temperature changes (Rosenzweig et al., 2008). Increasing global temperatures shifts the timing between food crop production and breeding seasons. The inconsistency between the two can ultimately lead to extinction of a species by reducing the number of offspring produced and lowering survival rates. Birds are highly sensitive to the timing inconsistency with evidence suggesting the greater the timing mismatch, the greater the decline in species population (Both et al., 2006; Møller et al., 2008).

The most compelling example of ecological impact attributed to global climate change is occurring in the Antarctic where the average temperature during winter has risen 6 °C in the past 60 years, while the oceanic temperature has risen almost 0.7 °C. The population of the Adélie penguin, who thrive on large expanses of ice that have largely melted, has decreased 70% since the 1970’s. This has been attributed to higher temperatures, lower availability of food sources, and increased difficulty in breeding, along with a discrepancy in timing between abundant food supply and breeding season (Stokstad,

2007; Staudt et al., 2008). The previous examples show the negative effects of increasing temperatures, which is largely attributed to human activity and emissions.

Currently, several compounds are known to contribute to global climate change. The greenhouse gas compounds include water vapor, carbon dioxide (CO2), methane (CH4), nitrous oxide (NO), and chlorofluorocarbons (CFCs). Humans do not directly influence the concentration of water vapor in the atmosphere, which is controlled by the water 2

cycle, so is not discussed in association with global warming (de Nevers, 2000). Carbon dioxide is the most significant contributor to global climate change, as measured by the radiative forcing index (Forster et al, 2007; Hofmann, 2009). The significant contribution of CO2 to global climate change can be solely attributed to its relative abundance in the atmosphere when compared to the other greenhouse gases.

Antarctic ice core records, which are used to determine historic CO2 concentrations, show a fairly constant concentration of CO2 in the atmosphere for centuries prior to the

Industrial Revolution, which occurred in the late 1700’s (Schwander and Stauffer, 1984;

Neftel et al., 1985). Beginning in 1958, continuous atmospheric CO2 concentrations have been measured at various locations around the world, most notably the Mauna Loa

Observatory (Scripps, 2009). Since the Industrial Revolution, atmospheric CO2 concentrations have steadily increased from 280 parts per million (ppm) to its current value of 385 ppm, representing an increase of 35% (Neftel et al., 1985; Tans, 2009).

More recently, annual global CO2 emissions have been growing more than 3% per year, which is mainly attributed to the combustion of fossil fuels (Raupach et al., 2007).

In 2006 the United States produced 2,300 million tons of CO2 from coal-combustion for electricity generation, while a total of 6,500 million tons of CO2 was produced (EIA,

2007a). Also in 2006, electricity generated from coal combustion produced roughly 2 billion Megawatt-hours, while a total of roughly 4 billion Megawatt hours of electricity was generated in the United States (EIA, 2007b). This translates into coal-combustion for electricity production accounting for 35% of the CO2 emissions and 50% of the electricity generation in the United States in 2006. Worldwide, coal-combustion is responsible for 42% of the CO2 emissions, while providing 41% of the electricity

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generated in 2006 (IEA, 2008). Due to the significant CO2 emissions output by power plants, a concerted, worldwide effort is ongoing to develop economical CO2 capture systems for coal-fired power plants.

Unfortunately, the concentration of CO2 in a coal flue gas stream is fairly dilute, comprising only 10% to 15% (volume). Although the dilute CO2 concentration makes separation difficult, several technologies are being developed to capture CO2 from coal- fired power plants. Currently, though, very few commercial-scale technologies exist for effectively capturing carbon dioxide.

Carbon dioxide capture is achieved through absorption, adsorption, and research is occurring in several other areas including membranes, oxy-combustion, and reactive separation.

Solvents capable of absorbing acid gases such as carbon dioxide have been available for over 50 years and continue to improve. The most common chemical solvents used to remove carbon dioxide are amine-based, such as Monoethanolamine (MEA),

Diethanolamine (DEA), and Methyl Diethanolamine (MDEA). The CO2-laden stream and chemical solvent are contacted countercurrently at temperatures slightly above ambient (40°-60° C). To regenerate the solvent, the CO2-solvent stream is heated to100°

C – 120° C and stripped with steam to break the CO2-solvent bond. Physical solvents include Selexol, which is a dimethylether of polyethylene glycol, and Rectisol, which employs chilled methanol to remove carbon dioxide from streams. The CO2-solvent bond occurs at high pressures and a reduction in pressure releases the CO2-solvent bond to regenerate the solvent and produce the pure CO2 stream (Wheeldon, 2005; Klemeš et al.,

2006; Freeman, 2007).

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Adsorption onto the surface of a high-surface area solid is under development. Several types of adsorbents have the ability to capture CO2. Currently, molecular sieves can retain 246 g CO2/kg adsorbent (Yang et al., 2008). A bed of high-surface area adsorbent such as alumina, zeolites, or activated carbon could also be used to adsorb the CO2 onto the sorbent. Desorption, which regenerates the adsorbent, occurs through Pressure Swing

Adsorption, Thermal Swing Adsorption, or the use of a purge gas. However, adsorption is a less developed technology and not truly considered for power plant CO2 capture due to its poor capacity and selectivity, which makes the technology economically prohibitive

(Klemeš et al., 2006).

Membranes that possess a high selectivity towards CO2 permeability are also being developed. Currently, membrane research has not produced a membrane that is durable while also being selective enough to produce a high-purity CO2 stream. Membrane selectivity of carbon dioxide to nitrogen, CO2/N2, should be over 200, yet current membrane selectivity is around 60 (IPCC, 2005; Favre, 2007; Yang et al., 2008).

Research into gas absorption membranes has shown promise, but it is still in the early stages of development.

Also in the early stages of development for CO2 removal from flue gas is the employment of an electrochemical pump, such as a molten carbonate or alkaline fuel cell. However, several problems plague its use for power plant CO2 capture (Granite et al., 2005).

Reactive separation utilizes a solid sorbent, typically a metal oxide, to react with CO2 to form a stable carbonate compound. Several compounds have the capability of capturing

CO2. Some examples include calcium oxide (CaO), sodium carbonate (Na2CO3), and lithium orthosilicate (Li4SiO4). The solid carbonate compound that is formed during CO2

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removal is then calcined in a separate reactor, where a pure stream of CO2 can be produced while simultaneously regenerating the solid sorbent. The overall process is carbon neutral, with the generated CO2 being sequestered. Chapter 2 will focus on the current status of reactive separation using solid sorbents for CO2 capture applications.

The overall focus of this research project is to develop a process that utilizes naturally occurring limestone, which is predominantly calcium carbonate (CaCO3), to capture CO2.

The reactions that occur, although simple in theory, are quite complex in mechanism and practice. Chapter 3 is a literature review of calcium carbonate calcination, which is the most energy-intensive step that decomposes calcium carbonate into calcium oxide and carbon dioxide. Chapter 4 focuses on the experimental results obtained in a sub-pilot scale unit that integrates the entire process of CO2 capture from a coal-combustion flue gas stream, known as the Carbonation-Calcination Reaction (CCR) Process. Chapter 5 focuses on the commercial aspects of the CCR Process, which includes a full set of computer simulations using Aspen Plus software for process integration into a 500 MWe power plant using results obtained from the sub-pilot scale. Chapter 6 provides conclusions and recommendations.

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CHAPTER 2

REVIEW OF SOLID SORBENTS FOR CO2 REMOVAL

2.1 Introduction

The attractiveness of using a metal oxide for CO2 removal is its reversibility potential and high capacity towards CO2 sorption. For example, when CO2 reacts with CaO, one mol of

CaO has the ability to react with one mol of CO2. In terms of weight, 56 grams of CaO has the capacity to react and retain 44 grams of CO2, which is approximately 78.5% by weight of CaO. In other words, 1 kg of CaO can retain 785 grams of CO2. Compared with the maximum capacity of synthetic molecular sieves of 246 grams per kilogram adsorbent, calcium oxide has over three times the carrying capacity of the molecular sieve. Additionally, the metal oxides are typically obtained from naturally occurring metal carbonates. The solid sorbents currently being studied include lithium orthosilicate, sodium bicarbonate, and calcium oxide.

2.2 Lithium Orthosilicate – Li4SiO4

Lithium orthosilicate is synthesized from lithium carbonate, a naturally-occurring carbonate found in the mineral zabuyelite, and silicon dioxide (SiO2). The calcination, or decomposition, temperature is 710 °C (Kato et al., 2002). The carbonation reaction, shown in Equation 2-1, occurs between 450 °C and 700 °C.

Li4SiO4 + CO2 Li2SiO3 + Li2CO3 [ Eq 2-1 ]

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The calcination occurs at temperatures greater than 710 °C. With a molecular weight of

119.85 and a 1:1 stoichiometric Li4SiO4:CO2 mol ratio, the Li4SiO4 is capable of retaining 36.7% of its weight, or 367 g CO2/kg Li4SiO4.

When used as a powder, the Li4SiO4 retains 25% of its weight in CO2 at 500 °C and a

20% CO2 atmosphere (Kato et al., 2005). In pure CO2 at 700 °C, Li4SiO4 was able to remove virtually all CO2; however, the pure CO2 stream is not an accurate representation of actual conditions encountered in a power plant. Over 50 cycles in a

Thermovgravmetric Analyzer (TGA) using pelletized Li4SiO4 exhibited little degradation

(Kato et al., 2005).

More realistic conditions using a slurry bubble column reactor at 700 °C for carbonation and 850 °C for calcination revealed that carbonation required approximately 30 minutes for carbonation and 30 minutes for calcination (Teresaka et al., 2006). Overall, one cycle requires approximately one hour, which may be prohibitive in a power plant where the flue gas flow rate are on the order of 1 million cubic feet per minute.

2.3 Sodium Carbonate - Na2CO3

Currently, the solid sorbent tested at the largest scale for CO2 removal is sodium carbonate. Carbonation occurs in the temperature range of 60 °C to 80 °C and is calcined at temperatures greater than 120 °C (Freeman, 2007). The reversible reaction is shown in

Equation 2-2.

Na2CO3 + H2O + CO2 2 NaHCO3 [ Eq 2-2 ]

Preliminary results have shown greater than 90% CO2 can be removed in 10 to 20 seconds (Freeman, 2007). More recently, the sorbent has been tested on a fossil-fuel fired combustor. The reactors were located downstream of a Flue Gas Desulfurization Unit,

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where sulfur dioxide (SO2) concentrations never rose above 20 ppm. Over 230 hours of testing, with flue gas streams derived from both coal and natural gas, has demonstrated the sorbent’s ability to remove 90% of the CO2 for several hours (Nelson et al., 2009).

Although successfully demonstrated, issues arise with sorbent deactivation with temperature fluctuations and the unknown effect of high sulfur concentrations.

Preliminary results indicate that sodium carbonate is a feasible solid sorbent for CO2 removal.

2.4 Calcium Oxide – CaO

The most widely tested solid sorbent for CO2 removal is calcium oxide. Natural deposits of limestone, whose main component is calcium carbonate, are abundant, composing slightly over 2% of the earth’s crust (Wedepohl, 1995; BCS, 2002). With an abundant supply, low prices, extensive history for sulfur removal, and high capture capacity, the calcium oxide/calcium carbonate looping cycle has potential for economically removing

CO2 from a flue gas stream.

2.4.1 Instituto Nacional del Carbón - Juan Carlos Abanades

Since 2002, Abanades’ research has focused on using natural limestone in a cyclical fashion for CO2 removal. An important observation and quantification that was made relates to the negative impact cyclic loops have on the limestone sorbent (Abanades,

2002). Over multiple cycles, natural calcium oxide will continually lose reactivity towards CO2, regardless of reaction conditions and residence times (Abanades, 2002).

Under fairly mild conditions of 650 °C carbonation and 850 °C calcination, a 50% decrease in conversion occurs between the first cycle and the last cycle (Abanades,

2004).

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In a series of TGA tests, several variables were examined, including calcination temperature, CO2 partial pressure for carbonation, limestone source, and number of cycles. Over a large number of cycles, 0.1 millimeters and 1 millimeter in diameter particles did not affect conversion. No significant differences were also observed between the two limestones tested. If calcination occurred between 850 °C and 950°C, the calcination temperature did not affect the carbonation conversion over the cycles.

However, a marked decrease in conversion occurs if calcining occurs above 950 °C.

Finally, at higher CO2 partial pressures, the sorbent conversion performance is better over multiple cycles (Grasa and Abanades, 2006).

An economic analysis of a CaO/CaCO3 cyclic loop integrated into a Pressurized

Fluidized Bed Combustor was also conducted. The Carbonator removes approximately

80% of the CO2 in the flue gas stream at a Calcium:Carbon mol ratio of 4. The Calciner is direct-fired using petroleum coke or coal and oxygen produced from an Air Separation

Unit. The overall SO2 removal is 90% and has a relatively high purge stream at 7.5%

Overall, the cost per metric ton of CO2 captured was approximately $24, which is significantly lower than the estimates placed on amine scrubbing, which varies between

$39 and $96, with all costs provided in 2005 Canadian dollars (MacKenzie et al., 2007).

Under the given conditions, the economics of utilizing calcium oxide as a solid sorbent for CO2 removal show the process is economically feasible.

Two final analyses examined the calciner energy requirements and the conditions available for integrating the calcium oxide/calcium carbonate loop into a coal-fired power plant. The results of the analyses concluded that proper thought must be given to the solids circulation, sorbent reactivity and decay, CO2 removal, and amount of fresh CaCO3

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addition to maintain a constant CO2 removal. Also, the possibility exists to install the calcium oxide/calcium carbonate in a coal-fired power plant and achieve high removal of

CO2; however, several conditions also exist where high removal of CO2 is not possible.

The conditions used for experiments and simulations must be carefully considered to determine its true CO2 removal feasibility in a coal-fired power plant (Rodriguez et al.,

2008; Alonso et al., 2009).

2.4.2 CANMET and University of British Columbia

Research into using a calcium oxide/calcium carbonate cycle has been occurring jointly in Canada between the University of British Columbia, led by John R. Grace, and

CANMET, which is Canada’s leader in clean energy research and technology

(CanmetENERGY, 2009). Three high-calcium limestones with identical particle size distributions were tested in a fluidized bed reactor under a CO2/SO2 atmosphere (Ryu et al., 2006). As confirmed previously by Abanades, CO2 removals decreased over multiple cycles (Abanades, 2002; Ryu et al., 2006). The effect of SO2, not previously tested, further increased the drop in CO2 removals over multiple cycles (Ryu et al., 2006). The level of decrease was limestone-dependent. The sulfation pattern of the limestone, by scanning with x-rays, shows its influence on CO2 removals over multiple cycles (Ryu et al., 2006). Since the CaO/SO2 cycle is not reversible like the CaO/CO2 cycle, the capture capacity of the CaO will automatically decrease in the presence of SO2. After a number of cycles, the CaO will be completely converted to CaSO4 if fresh solids are not added to replace the build-up of CaSO4 (Ryu et al., 2006).

By increasing the number of pollutants to remove, the integration options increase in complexity. Since sulfur is a component of coal, two options exist. One is to remove the

11

sulfur prior to CO2 removal; the other option is to remove the sulfur simultaneously with

CO2. Within each option are additional integration options. Four options were investigated (Sun et al., 2007). First SO2 is removed prior to the carbonation/calcination loop. Second, CaO is cycled in the carbonation/calcination loop until the reactivity of the

CaO towards CO2 is negligible. The CaO is then reacted with SO2. The third and fourth options are identical to the first two, except the reactors examined are Pressurized

Fluidized Bed Reactors instead of Atmospheric Fluidized Bed Reactors (Sun et al.,

2007). Results from a Pressurized TGA using 1.8 MPa at 850 °C and 14% CO2 for carbonation and 0.1 MPa in pure nitrogen at 850 °C indicate that carbonation/calcination cycles should occur prior to sulfation to protect the CaO reactivity. Even low concentrations of SO2 decreased the calcium utilization of CaO (Sun et al., 2007).

To prevent CaO decay, two different treatment options were investigated (Sun et al.,

2008). One involved the use of water/steam in multiple locations of the loop. From prior research, calcium hydroxide has shown superior reactivity towards SO2. With the limestone used in the experiments, calcium hydroxide performed only slightly better than calcium oxide over multiple cycles. When used as a sweep gas for CO2 in the calciner, the steam neither enhanced nor degraded the product CaO. Steam could be used as a sweep gas for the CO2, if necessary, to lower calcination temperatrues. With no SO2 present, the addition of steam during carbonation did not enhance carbonation. However, with SO2 present, the sorbent in the presence of steam did perform slightly better over multiple cycles. The addition of modifiers were also tested. Although several modifiers were added, most showed very little improvement to cyclic CO2 capture enhancement.

Alumina, Al2O3, has shown some promise. All experiments were conducted at 850 °C

12

and pure CO2 (Sun et al., 2008). Since the conditions were all exaggerated when compared to actual and optimal conditions, the results obtained may differ greatly from results obtained under more realistic experimental conditions.

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CHAPTER 3

CALCINATION OF LIMESTONE

3.1 Introduction

The calcination of limestone is a highly studied reaction due to its importance in the construction industry, and more recently, the environmental sector. The calcination of limestone, which decomposes calcium carbonate into calcium oxide and carbon dioxide, has been used for centuries to produce CaO as construction material. Evidence suggests limestone calcination occurred as far back as 10,000 years ago (Felder-Casagrande et al.,

1997). Currently, the dominant use of lime occurs within the chemical and industrial sector, with only a small fraction of produced lime used for construction (Miller, 2007).

3.2 Mechanism

Heat and mass transport phenomenon govern the calcination of limestone. Figure 3.1 depicts the individual steps of calcium carbonate calcination (Boateng, 2008).

Figure 3.1 Thermal and mass transport phenomena governing calcium carbonate dissociation (adapted from Boateng, 2008)

14

The calcination of calcium carbonate has been established to occur through a five-step process (Kumar et al., 2007; Boateng, 2008). They are as follows:

1.) Heat of sufficient quality reaches the calcium carbonate surface

2.) Heat is transferred from the surface to the reaction interface

3.) Calcium carbonate dissociates into calcium oxide and CO2

4.) CO2 migrates from the reaction interface through the layer of lime formed to the

surface of the particle

5.) CO2 migrates away from the surface

Given the numerous steps required for calcium carbonate calcination, a number of process parameters greatly influence the calcination rate and quality of products. The parameters that will affect the calcination in a calcium oxide/calcium carbonate cycle for

CO2 removal include particle size distribution, temperature profile, and calcination atmosphere.

3.3 Factors Affecting Calcium Carbonate Calcination

3.3.1 Particle Size Distribution

The particle size of the limestone can limit both heat and mass transfer. At larger particle sizes, the calcium oxide product layer can inhibit mass transfer (Borgwardt, 1985). For the same starting material under identical calcination conditions, experimental results, independent of reactor type, show an increase in calcination time with increasing particle size distribution (Borgwardt, 1985; Criado and Ortega, 1992; Kumar et al., 2007). No detailed studies have been performed correlating the influence of particle size to reactivity of calcium oxide. One data point exists, with no significant difference in

15

surface area between a 10 micrometer particle and 50 micrometer particle calcined at

1000 C for approximately 0.35 seconds (Borgwardt, 1985).

3.3.2 Temperature Profile

Sintering is a process where particles coalesce and adhere to one another (Borgwardt,

1989a; Kumar et al., 2007). The closure of pores, which can be measured either directly or inferred from a decrease in specific surface area, is a general result of sintering and directly affects the reactivity of the product calcium oxide (Moropoolou et al., 2001;

Kumar et al., 2007). At higher calcination temperatures, the effect of sintering increases

(Glasson, 1958; Borgwardt, 1989a). Although no chemical composition difference exists between sintered calcium oxide and reactive calcium oxide, calcination of limestone above 1550 C produces a calcium oxide that is unreactive (Kumar et al., 2007). With respect to a calcium oxide/calcium carbonate cycle for CO2 removal, it is highly desirable to obtain a calcium oxide product with a high reactivity. Obtaining a high reactivity calcium oxide then requires low calcination temperatures, which is directly linked to the calcination atmosphere.

3.3.3 Calcination Atmosphere

The calcination temperature and calcination atmosphere are directly linked, where the calcination atmosphere dictates the minimum calcination temperature. More specifically, the partial pressure of CO2 during calcination controls the minimum calcination temperature (Stanmore and Gilot, 2005). Research has also shown that the presence of

CO2 as a sweep gas increases the calcination time and enhances sintering, as measured by a decrease in surface area (Glasson, 1961; Ewing et al., 1979; Borgwardt, 1989b). To minimize sintering, which will maximize calcium oxide reactivity, requires the reduction 16

of the partial pressure of CO2 experienced by the calcium oxide at its reaction interface.

This is achieved through the use of a vacuum pump or a diluent (sweep) gas.

When calcium carbonate is calcined under a hard vacuum between 450 C and 750 C, a high surface area calcium oxide is produced (Glasson, 1958; Glasson 1961, Borgwardt,

1989a). The surface area of the calcium oxide has been found to be directly proportional to the extent of calcium oxide reactivity towards CO2 (Sakadjian et al., 2007).

The most popular diluent gas for calcination has been steam. Since the early 1900’s, research has been ongoing with respect to steam calcination. However, to date, the effect of steam is still largely unknown. One of the earliest studies compared the calcination of limestone in steam versus air under identical conditions and did not observe an improved rate of calcination (Berger, 1927). Differing results were obtained by MacIntire and

Stansel (MacIntire and Stansel, 1953). The results obtained comparing the rate of calcination in steam versus a rotary kiln atmosphere clearly indicated faster kinetics under a steam atmosphere at identical temperatures (MacIntire and Stansel, 1953). The contradictory results were attributed to a difference in particle size distribution. More recently, a mechanism and explanation supporting the catalytic effect of steam during limestone calcination was given (Wang and Thompson, 1995). The results of their experiments clearly indicate the addition of steam can both lower the calcination temperature and increase its kinetics (Wang and Thompson, 1995). The results were explained by the rate of adsorption of water vapor on calcium carbonate being greater than the rate of adsorption of carbon dioxide; therefore, the water vapor replaces the carbon dioxide and weakens the CaO-CO2 bond (Wang and Thompson, 1995). If Wang and Thomson are correct, a competition between lower calcination temperatures and

17

calcium oxide sintering exists since prior research has shown steam to enhance calcium oxide sintering more than carbon dioxide (Borgwardt, 1989b; Mai and Edgar, 1989;

Stanmore and Gilot, 2005). Since lower calcination temperatures form a higher reactivity calcium oxide, while sintering of calcium oxide produces a lower reactivity calcium oxide, the use of steam as a diluent gas is still inconclusive due to its ability to lower the calcination temperature, yet it is known to enhance sintering.

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CHAPTER 4

SIMULTANEOUS CO2 AND SO2 REMOVAL OVER MULTIPLE CYCLES

4.1 Introduction

The Ohio State University has developed the CCR Process for simultaneous removal of both sulfur dioxide and carbon dioxide from a flue gas stream using a calcium-based solid sorbent. Initially, a mesoporous Precipitated Calcium Carbonate (PCC) developed at

The Ohio State University was intended solely for sulfur dioxide capture from a coal combustion flue gas stream (Mahuli et al., 1997; Wei et al., 1997). The high reactivity of the PCC led to a demonstration at the pilot-scale, in a process known as OSCAR, that utilized PCC to remove SO2 generated from a coal-fired boiler (Gupta et al., 2007).

Further research showed promising results for the capture of CO2 over multiple cycles using PCC (Gupta and Fan, 2002). Testing at the lab-scale using a TGA showed little degradation of the capture capacity of PCC towards CO2 over multiple cycles and the ability for PCC to simultaneously react with both CO2 and SO2 (Iyer et al., 2004). The high capture capacity of PCC towards CO2 over multiple cycles, along with the successful demonstration of PCC reactivity towards SO2, indicated further testing at larger scales should be performed. A sub-pilot scale demonstration unit was eventually built. The sub-pilot scale facility utilizes a 20 pound per hour coal-fired combustion

19

unit and was successfully operated for single-cycle studies (Wong, 2007). With the facility initially set-up without the installation of a calciner to decompose calcium carbonate into calcium oxide, single cycle studies were conducted utilizing commercially available calcium hydroxide and calcium oxide. Calcium hydroxide showed promising results, allowing for greater than 90% CO2 removal with a low Calcium:Carbon (Ca:C) mol ratio and virtually complete SO2 removal, which was independent of the Ca:C mol ratio (Wong, 2007). The SO2 removal is virtually complete (greater than 99+%) due to the significant disparity between carbon and sulfur content in coal. For example, a 1:1

Calcium:Carbon mol ratio for a typical bituminous coal with 75% carbon content will possess a 50:1 Calcium:Sulfur ratio for a 4% sulfur coal. The successful single-cycle results indicated the need to renovate the existing facility to allow for the installation of process equipment necessary for cyclic studies.

4.2 Thermodynamics and Reactions

In the CCR process, a multitude of reactions are feasible, and multiple reactions occur within the temperature window of the CCR process. A balance between the kinetics of the reactions and its corresponding thermodynamics exists. In order to ensure reactions occur at a reasonable rate without violating the laws of thermodynamics, the CCR

Process can only operate within a narrow temperature range. The decomposition of the calcium hydroxide (Ca(OH)2), the carbonation of calcium oxide, the sulfation of calcium oxide, the calcination of calcium carbonate, and regeneration of calcium hydroxide are the reactions in the CCR Process.

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4.2.1 Hydration of Calcium Hydroxide Reversible Reaction

The calcination of calcium hydroxide is endothermic, and the decomposition reaction,

shown in Equation 4-1, begins at 512 °C in 1 atmosphere partial pressure of water vapor.

Figure 4.1 depicts the equilibrium relationship between the partial pressure of water and

temperature.

Ca(OH)2  CaO  H 2O (g) [ Eq 4-1 ]  H 298  109 kJ/mol

H2O Equilibrium Pressure versus Temperature for Ca(OH)2 1000

100 CaO + H2O (g) Ca(OH)2 10 1 0.1 0.01 0.001 Ca(OH)2 CaO + H2O (g)

0.0001

(atm) O O 2 1E-05 H P 1E-06 1E-07 1E-08 1E-09 1E-10

1E-11 1E-12 0 100 200 300 400 500 600 700 800 900 1000 1100 1200 Temperature (°C)

Figure 4.1 Equilibrium pressure versus temperature diagram for calcium hydroxide

21

With typical water vapor concentrations in flue gas between 3% and 7% (0.03 to 0.07 atm partial pressure water vapor), it is not expected for CaO to react with the water vapor in the flue gas stream to form calcium hydroxide. At a partial pressure of 0.07 atm H2O, the equilibrium temperature is approximately 400 °C. Above 400 °C and a H2O partial pressure of 0.07 atm, calcium hydroxide will spontaneously decompose into calcium oxide and water vapor. The kinetics of calcium hydroxide decomposition varies widely, based on temperature and reactor configuration; however, the millisecond time-scale is sufficient for calcination at temperatures greater than 700 °C (Ghosh-Dastidar et al,

1995).

The formation of calcium hydroxide is a crucial reaction in the CCR process. The calcium hydroxide is able to regenerate a reactive calcium oxide that maintains a high capture capacity towards CO2 (Fenell et al., 2007; Manovic and Anthony, 2007; Zeman,

2008). Traditional dry calcium hydroxide is formed at atmospheric pressure and ambient temperature using liquid water and calcium oxide (Miller, 1960; Boynton, 1980). An excess amount of water (3 to 4 times stoichiometric) is fed to the reaction vessel to control the process temperature of the highly exothermic reaction (Miller, 1960).

In order to be properly incorporated into the CCR Process, a high-temperature hydration is necessary to minimize the overall energy penalty encountered when integrating the

CCR process into a coal-combustion power plant. The high-temperature hydration provides two advantages over traditional hydration. First, the calcium oxide exiting the

Calciner possesses a minimum temperature of 900 °C, while the Carbonator/Sulfator operates at a minimum of 450 °C. In order to minimize heating the calcium hydroxide upon exiting the hydration reaction vessel (Hydrator), the Hydrator should operate near

22

the operating temperature of the Carbonator/Sulfator and the Calciner. If the traditional route for calcium hydroxide formation were used, the calcium hydroxide would require significant energy for heating the solids to the operating temperature of the

Carbonator/Sulfator. Second, the hydration reaction, as mentioned earlier, is a highly exothermic reaction. With respect to heat utilization and process efficiency, higher temperature heat is always more valuable than heat at lower temperatures. For example, at 500 °C, the exothermic heat of reaction can be utilized for generating steam for input into a steam turbine or heat exchanged with cold streams. The traditional route of dry calcium hydroxide formation would provide heat at 100 °C, which is minimal for steam generation or heat exchange (Miller, 1960). At temperatures greater than 512 °C, which would be more favorable in terms of heat integration and efficiency, the Hydrator would require a partial pressure of H2O greater than 1 atm. At elevated operating temperatures and pressures, the cost, safety, and reliability of the Hydrator would significantly increase the capital and Operations and Maintenance (O&M) costs.

4.2.2 Carbonation of Calcium Oxide Reversible Reaction

The carbonation reaction, shown in Equation 4-2, is exothermic and can occur at a reasonable rate between 450 °C and 750 °C (Abanades and Alvarez, 2003; Kuramoto et al., 2003; Lee, 2004; Wong, 2007).

CaO  CO2 (g)  CaCO3 [ Eq. 4-2 ]  H 298  -178 kJ/mol

Higher temperatures thermodynamically prohibit significant CO2 removals. At 750 °C, the equilibrium CO2 partial pressure is 0.099 atm (HSC Chemistry, 2008). A flue gas stream containing 15% CO2 would only be capable of 34% CO2 removal. 23

Previous research has shown a temperature range between 600 °C and 700 °C as optimal

(Iyer et al., 2004). For commercial application where a minimum target of 90% CO2 removal is expected, a typical flue gas stream containing 10% to 15% CO2 would require a final CO2 concentration between 1% and 1.5% (Ciferno et al., 2009a). A CO2 concentration of 1.5% corresponds to an equilibrium temperature of 660 °C, while a concentration of 1% corresponds to a temperature of 643 °C (HSC Chemistry, 2008).

Although the optimal temperature range for the CCR process lies between 600 °C and

700 °C, 90% CO2 removal is not thermodynamically possible at those upper temperature limits. Given the dilute concentration of CO2 in flue gas, the maximum temperature for carbonation is 660 °C, with lower maximum temperatures at lower CO2 concentrations.

This reduces the temperature window for the carbonation reaction by at least 40 °C, from a maximum optimal temperature of 700 °C to 660 °C. Figure 4.2 shows the relationship between the equilibrium CO2 partial pressure and temperature.

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CO2 Equilibrium Pressure versus Temperature for CaCO3 100

1

0.01

0.0001 CaCO CaO + CO 3 2 0.000001

0.00000001

1E-10 2

CO P 1E-12 1E-14 CaO + CO2 CaCO3 1E-16

1E-18 1E-20

1E-22

1E-24 0 100 200 300 400 500 600 700 800 900 1000 1100 1200 Temperature (°C)

Figure 4.2 Equilibrium Pressure of CO2 versus Temperature for Calcium Oxide. The solid black line shows the equilibrium temperature and pressure at 1 atm (~900 °C). The dotted black line represents the partial pressure of CO2 and corresponding temperature for 90% CO2 removal from a power plant flue gas stream (~660 °C).

4.2.3 Sulfation of Calcium Oxide

One of the advantages of the CCR Process is its unique ability to perform multi-pollutant

capture. The current leading commercial technology for CO2 removal is amine scrubbing.

However, the amine solvent is easily degraded by SO2 and should be maintained below

10 ppm (Rao and Rubin, 2002; Supap et al., 2009). The CCR process has the capability to

simultaneously remove both CO2 and SO2 without sorbent degradation.

Several technologies have been developed for post-combustion SO2 control. Worldwide,

87% of installed SO2 control technologies use a wet process with 97% using a calcium-

based sorbent. The other main technology injects a dry sorbent into the flue gas stream,

25

where the sorbent is still predominantly calcium-based (Srivastava, 2000). In any case, even for the lowest sulfur coals, a minimum of 98% SO2 removal is required to achieve the 10 ppm target; however, typical SO2 removal technologies are currently designed only for 90% SO2 removal (Stultz and Kitto, 1992; Srivastava, 2000).

The wide use of calcium-based sorbents for SO2 removal has provided a vast amount of information with respect to all aspects of the reaction. The thermodynamics, kinetics, and mechanism have all been widely studied (Marsh and Ulrichson, 1985; Mai,1987; Tullin and Ljungström, 1989; Bruce et al., 1989; Kuburović et al, 2002). Figure 4.3, located on the following page, depicts the equilibrium partial pressure of SO2 versus temperature for the thermal decomposition of calcium sulfate under its own self-generated atmosphere.

Lower decomposition temperatures can be achieved in a reducing atmosphere.

The maximum temperature encountered in the CCR Process where calcium sulfate is present is in the Calciner, which can be heated either in an indirect-fired manner or direct-fired manner. In an indirect-fired calciner, where the purity of CO2 is maintained by placing the heating source on the exterior of the calciner, the atmosphere in the calciner will be pure carbon dioxide, which does not alter the calcium sulfate decomposition temperature. In a direct-fired calciner, where the heat is generated through combustion with oxygen and fuel, ideally a sulfur-free fuel using stoichiometric ratios, the atmosphere in the calciner would consist of the hydrocarbon fuel, oxygen, CO2, and

H2O. In this instance, the possibility of calcium sulfate decomposing exists. If calcium sulfate does decompose, the resulting calciner gas exhaust mixture would contain CO2,

H2O, SO2, and O2. While the CO2, SO2, and O2 are easily separable from the H2O by condensing out the H2O, the separation of CO2 from SO2, and O2 is difficult. For this

26

reason, it is preferred to operate an indirect-fired calciner for the CCR process, or ensure that calcium sulfate does not decompose in a direct-fired calciner.

SO2 Equilibrium Pressure versus Temperature for CaSO4 0.1 0.001 0.00001 0.0000001 CaO + SO2 + 1/2 O2 CaSO4 0.000000001 1E-11 1E-13 1E-15 1E-17 1E-19 1E-21 1E-23

1E-25 (atm)

2

1E-27 SO

P 1E-29 CaSO4 CaO + SO2 + 1/2 O2 1E-31 1E-33 1E-35 1E-37 1E-39 1E-41 1E-43 1E-45 1E-47 1E-49 0 100 200 300 400 500 600 700 800 900 1000 1100 1200 Temperature (°C)

Figure 4.3 Equilibrium Partial Pressure of SO2 versus Temperature for Calcium Sulfate

4.3 Experimental

4.3.1 Chemicals, Sorbents, and Gases

In order to generate the flue gas stream, coal and natural gas are co-fired in a stoker. The high-sulfur (greater than 3.0%) stoker-grade coal was obtained from the Sands Hill Coal

Company located in Hamden, Ohio. The results of the chemical analyses, performed by

CONSOL Energy, are located in Tables 4.1 and 4.2. It must be noted that CONSOL only provided the Proximate Analysis on a dry basis and the percent moisture. Thus, the wet basis results were calculated from the dry basis and the moisture content. Natural gas,

27

provided by The Ohio State University, was delivered to the laboratory at 5 psig. A typical composition for natural gas is located in Table 4.3 (McGurl et al., 2005).

Proximate Analysis Weight % Weight % (Dry Basis) Moisture 6.065 Ash 7.036 7.49 Volatile Matter 38.626 41.12 Fixed Carbon 48.273 51.39 BTU/lb 13,311 MAF BTU/lb 14,389

Table 4.1 Proximate Analysis of as-received stoker-grade coal

Ultimate Analysis Weight % (Dry Basis) Carbon 73.91 Hydrogen 4.79 Nitrogen 1.43 Chlorine 0.00 Sulfur, Total 3.73 Ash 7.49 Oxygen (Difference) 8.65

Table 4.2 Ultimate Analysis of as-received stoker-grade coal

Component Chemical Formula Volume Percentage Methane CH4 93.1 Ethane C2H6 3.2 Propane C3H8 0.7 n-Butane C4H10 0.4 Carbon Dioxide CO2 1.0 Nitrogen N2 1.6 BTU/ft3 @ 1 atm 1032

Table 4.3 Typical composition of Natural Gas

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Two types of sorbents were tested, commercial-grade calcium hydroxide and commercial-grade calcium oxide. Graymont, in Bellefonte, Pennsylvania, provided both high calcium ground lime and high calcium hydrated lime. Carmeuse Lime & Stone provided hydrated lime that was used to complete the cyclic studies. Table 4.4 lists an approximate chemical composition for Graymont ground lime and Graymont calcium hydroxide, as provided by the manufacturer.

Component Graymont Calcium Hydroxide Graymont Calcium Oxide CaO Minimum 72.0% Minimum 94.0% MgO Minimum 0.4% Minimum 0.5% CaCO 3 Maximum 1.1% ------MgCO3

Table 4.4 Composition of Graymont Calcium Hydroxide and Graymont Calcium Oxide

Theoretical flue gas streams, which assumes complete combustion, generated from coal combustion is listed in Table 4.5. The range of flue gas composition arises from the wide variability that occurs in the composition of coal and amount of excess air utilized during combustion. The theoretical values were calculated using a value of 20% excess air, which is a reasonable assumption (Stultz and Kitto, 1992). Table 4.6 lists a typical flue gas stream generated from the sub-pilot scale facility using the coal composition determined from the analysis.

29

Theoretical l Flue Chemical Species Gas Composition (Volume %) Nitrogen (N2) 70-80 Carbon Dioxide (CO2) 10-15 Water (H2O) 3-7 Oxygen (O2) 2-4 Sulfur Dioxide (SO2) 0.05-0.3 Nitrogen Oxides (NOx) --- Carbon Monoxide (CO) --- Sulfur Trioxide (SO3) ---

Table 4.5 Theoretical flue gas compositions derived from coal combustion

Typical Flue Gas Composition Chemical Species (Dry Volume %)

Nitrogen (N2) Balance

Carbon Dioxide (CO2) 10%

Water (H2O) 0%

Oxygen (O2) 8%

Sulfur Dioxide (SO2) 1500 ppm Nitrogen Oxides (NO ) 75 ppm x Carbon Monoxide (CO) 200 ppm

Sulfur Trioxide (SO ) Not Measured 3

Table 4.6 Typical flue gas concentration generated from sub-pilot scale facility

4.3.2 Facility Design and Set-up

The sub-pilot plant facility must meet several design criteria in order to conduct the CCR process on a larger scale. First, a coal-combustion unit is required to generate the flue gas stream. Second, since the carbonation and calcination reactions occur at high temperatures, 450 °C and 900 °C (minimums), respectively, ductwork and process equipment capable of withstanding and generating those temperatures are necessary.

Third, solids feeding equipment are necessary to transport the solid sorbent into the high-

30

temperature flue gas stream and calciner. Particulate capture devices are required to eventually separate the solids from the flue gas stream. Finally, process measurement equipment is necessary to monitor the temperature, pressure, and gas composition throughout the facility. Appropriate safety control measures are also essential for operating equipment around potentially poisonous gases and chemicals.

The original design and layout of the experimental facility was capable of single cycle testing only. The initial experimental set-up has been explained in detail elsewhere

(Wong, 2007). No major modifications were made to the ductwork, equipment, control systems, data acquisition systems, or insulation. However, to facilitate cyclic testing, additional pieces of equipment were added, and adjustments to the ductwork were made in order to integrate the equipment into the existing facility. The following description provides a basic outline of the experimental process and provides in detail the modifications made to the set-up for simultaneous CO2 and SO2 removal over multiple cycles in the 20 pound per hour facility. Figure 4.4 shows a process flow diagram of the

20 pound per hour facility, while Figure 4.5 is an actual photograph of the experimental layout with key equipment labeled.

An underfeed stoker donated by Babcock & Wilcox co-combusts approximately 20 pounds per hour of stoker-grade coal and 3 actual cubic feet per minute (ACFM) natural gas. The coal is stored in a hopper and delivered to the stoker via a variable-speed screwfeeder. The natural gas is available through The Ohio State University supply line and is directly fed into the stoker unit. Two Forced Draft (FD) fans provide combustion air necessary to generate the flue gas stream. A 30 horsepower, variable-frequency

Induced Draft (ID) fan, which can pull 3000 ACFM and generate 45 inches water static

31

pressure, is located at the end of the process. The ID fan is necessary to generate a negative pressure to transport the flue gas stream through the ductwork and towards the exhaust. The ID fan also ensures that all flue gas and sorbent in the ductwork will not be released into the laboratory atmosphere. A zero-pressure point is maintained in the stoker, where the positive pressure of the FD fans balances the negative pressure of the ID fan.

A Schenck-Accurate mid-range volumetric hopper, with a capacity of 300 pounds, stores the sorbent and is connected to a screwfeeder, with a maximum feed rate of 500 pounds per hour, which injects the solid sorbent into a FEECO rotary calciner. The feed rate of the sorbent is controlled by the revolutions per minute of the screw.

The FEECO rotary calciner is indirectly-heated using electric heaters with a maximum operating shell temperature of 982 °C. Due to the rotation and length of the calciner, the internal temperature of the calciner is not directly measured and can only be roughly estimated; however, the internal temperature will always be lower than the shell temperature due to heat losses and heat transfer resistance. The residence time is controlled by a variable frequency drive that determines the revolutions per minute of the rotary calciner and can be varied between 30 minutes and 45 minutes. For both carbonation and calcination, the residence time was never varied and maintained a residence time of 35 minutes. The sorbent, while in the calciner, can be either preheated to minimize the temperature drop that occurs in the carbonation reactor or calcined, depending on the calciner temperature setpoint. The calciner is positioned at a 2 degree downward slope to allow the solids to flow through the calciner via gravity while rotating. A Sunco double-dump valve, which acts as a gas-solid separator, and an exhaust are located at the outlet of the calciner. The double-dump valve allows the pressure and

32

atmosphere in the rotary calciner to be independent of the pressure and atmosphere in the flue gas stream, while also allowing the solids to enter the carbonation reactor, where the sorbent contacts the flue gas stream.

The Carbonator/Sulfator contacts the flue gas stream and the calcium hydroxide in the temperature range between 450 °C and 700 °C. The calcium hydroxide is injected in the downer of the Carbonator/Sulfator and is entrained by the natural flow of the flue gas stream. In the Carbonator/Sulfator, the calcium hydroxide simultaneously decomposes into calcium oxide and steam while reacting with both carbon dioxide and sulfur dioxide present in the flue gas stream to form calcium carbonate and calcium sulfate. The observed residence time in the entrained bed reactor can be varied between 0.3 seconds and 0.6 seconds by altering the location of the downstream gas analyzer.

Following the carbonation reactor is an air injection valve to significantly cool the flue gas temperature before entering a Donaldson Torit downflow baghouse, where solids exit into a 55-gallon drum and the CO2-and-SO2-lean, particulate-free flue gas stream is vented to the outdoor atmosphere.

At the completion of each carbonation cycle, the calciner outlet is disconnected from the

Carbonator/Sulfator and connected directly to a 55-gallon drum. The solids collected in the baghouse are then placed into the Schenck-Accurate mid-range hopper. The Calciner is pre-heated to a safe operating temperature of 950 °C. Upon completion of heating, the solids are screwfed into the Calciner for calcination.

In the calciner, the calcium carbonate decomposes into calcium oxide and carbon dioxide.

As previously discussed, due to the stability of the calcium sulfate, it remains as calcium sulfate in the calciner. The pure, dry CO2 gas exits through the exhaust of the calciner,

33

while the solid mixture, consisting of calcium oxide derived from calcium hydroxide decomposition, calcium oxide derived from calcium carbonate calcination, un-calcined calcium carbonate, and calcium sulfate, exits into a 55-gallon drum. The collected solids are then shipped to Carmeuse Lime and Stone (Pittsburgh, PA) where they are hydrated at atmospheric conditions to produce a dry hydrate, which completes the cycle. The dry hydrate formed is used as the feed for the next cycle.

To monitor the gas composition and analyze the percent removal of both carbon dioxide and sulfur dioxide, two sets of gas analyzers are employed. One set of gas analyzers is located upstream of sorbent injection and is used as the baseline. The other set of gas analyzers is located downstream of the sorbent injection. The difference between the two measurements determines the percent removal. The gas analyzers are CAI 600 analyzers and continuously monitor the concentrations of CO2, SO2, and CO. In addition, a CAI

NOxygen analyzer monitors the upstream oxygen and nitrogen oxides concentration, while a Teledyne Analytical 3000P analyzer monitors the downstream oxygen concentration. All data is continuously recorded by a computer.

In addition, over 20 thermocouples continuously measure the temperature throughout the system. Several manometers are also used to measure the static pressure and pressure drop throughout the system. The pressure drop is used as an indicator of normal operation. If the pressure drop increases greatly in one section, it is an indication that solids have accumulated in that section and that solids are not flowing properly. In such an instance, compressed air generated from an Atlas Copco GA7FF air compressor is available for providing aeration air. In addition, under normal operating conditions, the pressure drop is an indication to the cost of the CCR process at the commercial-scale. In a

34

power plant, the greater the pressure drop of the flue gas, the greater the cost to operate the plant as larger pumps and fans are required to transport the flue gas.

In order to ensure the safety of all personnel, two QRAE Plus Multi-gas alarms, with forced pumps, are placed around the stoker and ductwork to continuously monitor the concentration of carbon monoxide, sulfur dioxide, and methane. In addition, multiple carbon monoxide monitors are placed around the facility.

Figure 4.4 Process flow diagram of sub-pilot scale facility. D denotes duct, C denotes calciner, TT denotes temperature thermocouple, GS denotes gas sampling

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7 4

5 3 2 8

6 9 1

Figure 4.5 Image of 20 pound per hour coal combustion sub-pilot scale facility. 1: Stoker. 2: Upstream gas analyzer. 3: Flyash/SO2 sorbent injection hopper. 4: Main sorbent injection hopper. 5: Rotary Calciner. 6: Entrained bed reactor. 7: Downstream gas analyzer. 8: Air dilution valve. 9: Baghouse.

4.4 Results and Discussion

The installation of the calciner allows for multiple cycles of a sorbent to be tested. This was the sole purpose and motivation of the sub-pilot scale facility experiments. The results from the single cycle study experiments have been detailed elsewhere (Wong,

2007). Two types of sorbents were tested. First, due to the success of the single cycle studies, calcium oxide derived from calcium hydroxide was tested. Second, calcium hydroxide was regenerated after every cycle and tested. Mechanical problems, equipment maintenance, and time required for hydration severely limited the number of sorbents and number of cycles tested. However, consistent solids utilization and CO2 removal was achieved for a limited number of cycles.

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4.4.1 Initial Testing

The initial tests conducted were used to confirm proper integration of equipment and verify that the integrated facility produced replicable results of the single cycle studies.

The initial set-up consisted of a cyclone located before the baghouse to separate the sorbent from the flue gas stream. Below the cyclone was a Sunco rotary valve, whose function was to allow the solids from the cyclone to enter the calciner while independently maintaining the pressure in the calciner and cyclone. Beneath the rotary valve was a screwfeeder for fresh feed injection. Figure 4.6 depicts the initial set-up that included the rotary valve. However, the rotary valve did not create an airtight seal, which allowed the injected solids to be pulled upwards through the cyclone and into the baghouse. In effect, no sorbent was fed into the calciner and the ductwork, as the negative pressure of the ID fan was greater than the negative pressure located in the Calciner.

Thus, the rotary valve was removed and the cyclone bottom disconnected from the

Calciner, which then forced all solids to be collected in the baghouse.

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Cyclone

Rotary Valve

Sorbent Injection

Calciner Inlet

Figure 4.6 Initial set-up of calciner feed integration

4.4.2 Calcium Oxide Carbonation-Calcination Cycles

The goal of the CCR Process is to utilize calcium oxide/calcium carbonate in a cyclic process to capture CO2 and SO2. In order to verify that the additional equipment added and facility renovation required for cyclic studies did not alter the results of the single cycle studies, the initial sorbent tested was Graymont high-calcium calcium hydroxide.

The targeted Ca:C mol ratio was 0.75 because reliable results for CO2 and SO2 removals were obtained for single cycles. A 50:50 weight percent mixture of calcium hydroxide:pulverized ground lime was utilized to ensure proper flow of the solids. Before

38

entering the ductwork, the sorbent must first pass through the rotary calciner. Based on the results of the single cycle studies, a temperature drop in the duct would be expected if the sorbent were injected at ambient temperature. Thus, to avoid the temperature drop, the sorbent was preheated to 500 °C in the rotary calciner. This is below the spontaneous decomposition temperature of calcium hydroxide, which is 580 °C. The spontaneous decomposition temperature is different, and higher, than the equilibrium temperature since the decomposition reaction ensures the prevention of calcium hydroxide formation.

By significantly increasing the temperature over the equilibrium temperature, the kinetics of the decomposition reaction is increased while also increasing the equilibrium partial pressure of H2O. By increasing the temperature, the decomposition is spontaneous and complete.

The results of the first cycle verified the single cycle results. The solids were then collected in the baghouse and fed into the Calciner, which was operating at 850 °C and approximately 10 moles of steam per mol of CO2. By injecting steam, the partial pressure of CO2 in the calciner atmosphere is lowered, which lowers the calcination temperature.

The resulting calcined CaO sorbent was then injected into the carbonation reactor and

20% CO2 removal was obtained. The Ca:C mol ratio increased to 2 since the active calcium sorbent now included calcium oxide from both the flow enhancer and the calcined calcium carbonate. From single cycle studies, a Ca:C mol ratio of 2 for pulverized ground lime is approximately 12%, based on a linear interpolation. Given both

CO2 removals have an approximate standard deviation of 7%, there is no statistical difference between the single cycle removal of 12% from the linear interpolation of the single cycle and the 20% removal from the second cycle. Since the initial starting

39

material for the second cycle is equivalent to the starting material from the single cycle

studies, this verifies that the removals from both calcium hydroxide and calcium oxide

with the calciner integrated are equivalent to the single cycle studies. A third cycle was

also run, but the observed CO2 removal was zero, so no further cycles were conducted.

Figure 4.7 shows the relationship between percent CO2 removal and cycle number.

The decay in CO2 capture capacity over multiple cycles using calcium oxide derived

from naturally occurring calcium carbonate has been described in Chapter 2.

Percent CO2 Removal versus Cycle Number 100

90

80

70

60

Removal 2 50

40

Percent COPercent 30

20

10

0 0 1 2 3 4 Cycle Number

Figure 4.7 Relationship between Ca(OH)2-Cycle 1 and CaO-Cycle 2 versus CO2 removal

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4.4.3 Calcium Hydroxide Carbonation Calcination Cycles

The rapid decrease in capture capacity of calcium oxide over three cycles will lead to a high solids loading and annual O&M cost for a target of 90% CO2 removal from commercial units. Extrapolating from the single cycle studies, assuming a linear trend, a

Ca:C mol ratio of 15.5 would be required for 90% CO2 removal if using pulverized

Graymont ground lime. If Graymont ground lime were used, a 26.5 Ca:C mol ratio would be required for 90% CO2 removal. For a 500 MWe power plant with an efficiency of

33% (subcritical plant) and burning coal with a 70% carbon content, 90% CO2 capture would require nearly 17,000 tons and 28,000 tons of calcium carbonate per hour for pulverized Graymont ground lime and Graymont ground lime, respectively. From single cycle studies, calcium hydroxide can remove 90% CO2 at approximately a Ca:C mol ratio of 1.3. The advantage of using calcium hydroxide is twofold. First, the solids loading is dramatically reduced. Second, the total cost is significantly reduced. Since the solids loading is lower, equipment size is smaller, which leads to a lower capital cost. Since less sorbent is required per year, due to a lower Calcium:Carbon mol ratio required for 90%

CO2 removal, the annual O&M cost is lowered. Together, this leads to a lower overall cost when calcium hydroxide is used over calcium oxide. The question then remains whether calcium hydroxide can maintain reactivity over multiple cycles in a large-scale unit. Chapter 2 has already covered the bench-scale cyclic activity of calcium hydroxide.

A total of 6 cycles were conducted. The first cycle utilized Graymont calcium hydroxide.

The remaining 5 cycles was calcium hydroxide hydrated by Carmeuse Lime & Stone.

The process for the cyclic tests began with a 50 weight percent Graymont calcium hydroxide and 50 weight percent grinded Graymont ground calcium oxide, which was

41

used to enhance solids flow. The grinded Graymont ground calcium oxide was assumed to be inert, while the calcium hydroxide was assumed to be the active sorbent for simultaneous CO2 and SO2 removal. However, the hydration process is non-selective in nature and produced calcium hydroxide from calcium oxide, regardless of the source. In this manner, the amount of calcium hydroxide received upon return from Carmeuse increased with cycle number.

Figure 4.8 shows the percent CO2 removal over multiple cycles. The sorbent is assumed to be calcium oxide derived from calcium hydroxide. Although the starting precursor is calcium hydroxide, given the relative rate of calcium hydroxide calcination over carbonation and sulfation, the active sorbent is assumed to be calcium oxide. From Figure

4.8, calcium hydroxide maintained its reactivity over multiple cycles. Variations in CO2 removal arise from two sources of error. First, error is introduced in the measurement itself. The standard deviation of each average varied between 6% and 10%. Second, error is introduced in the Calcium:Carbon mol ratio. Although a 0.75 ratio was targeted, variations between 0.67 and 0.87 occurred. An exact 0.75 target could not be achieved due to the use of a volumetric feeder, which is based on bulk density of the solids rather than mass flowrate. Any variation in bulk density alters the mass feed rate.

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% CO2 Removal versus Cycle Number for Ca(OH)2 100

90

80

70

60

Removal 50 2

CO % 40

30

20

10

0 0 1 2 3 4 5 Cycle Number

Figure 4.8 Average CO2 removals versus Cycle Number for Calcium Hydroxide

Figure 4.9 shows a typical temperature profile for the duct beginning in the stoker (D-TT-

01) until just prior to the baghouse (D-TT-14).

Around 235 minutes, the sorbent begins to exit the calciner and enters the flue gas stream, which is observed through the slight temperature drop at D-TT-04, located just after the sorbent injection location, while the temperature just before injection, D-TT-03, continued to rise. Due to the calciner pre-heating the solids, the drastic temperature drop that occurred during single cycle studies does not exist. The temperature during injection remained fairly constant between 600 °C and 700 °C.

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Duct Temperature versus Time 1000 D-TT-01 900 D-TT-02 D-TT-03 D-TT-04 800 D-TT-05 D-TT-06 700 D-TT-07 D-TT-08 D-TT-09 600 D-TT-10 D-TT-14 500

400

300 Temperature (° Temperature C)

200

100

0 0 50 100 150 200 250 300 350 400 450 Time (minutes)

Figure 4.9 Typical temperature profile exhibited in sub-pilot scale facility. Thermocouples labeled correspond to thermocouple locations in Figure 4.4

4.5 Conclusions

The sub-pilot scale facility was successfully renovated to include a rotary calciner, which allowed for cyclic testing. The rotary calciner minimized the temperature drop of the flue gas stream during sorbent injection and also allowed for calcination of the calcium carbonate to regenerate the sorbent. Two types of sorbents were tested over multiple cycles. The first sorbent, calcium oxide derived from naturally occurring high-calcium calcium carbonate, was shown to be unsuitable for cyclic CO2 capture. Even under favorable, unrealistic calcination conditions for a commercial process and a high

Calcium:Carbon mol ratio, the calcium oxide exhibited extreme decay in CO2 capture

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capacity. Calcium hydroxide regenerated after every cycle maintained its reactivity over

5 cycles. Under realistic calcination conditions, 950 °C for 35 minutes, the calcium oxide formed was still able to hydrate. The formed hydrate did not decrease in reactivity and was able to maintain its high level of CO2 removal over 5 cycles.

4.6 Recommendations

In order to perform a greater number of cycles, a hydrator should be integrated into the setup. A truly continuous system that includes the hydration step is necessary in order to determine the rate of sorbent degradation, which was not visible over five cycles. A proper gas-solid contactor that does not require the use of a flow enhancer should also be designed. This is important to minimize the amount of solids in circulation, while also removing the amount of calcium hydroxide generated with each cycle. Since the hydrator, by necessity, is placed after the calciner, calcium oxide, regardless of initial source, will be hydrated. By using pulverized Graymont high-calcium calcium oxide as the flow enhancer, the hydration step, by nature, converts the calcium oxide into calcium hydroxide and diluting the initial feed of calcium hydroxide.

Furthermore, the flow enhancer impedes any experiments to determine the effect of sulfate accumulation in the system on the simultaneous CO2 and SO2 capture. The sulfate content in the solids stream, if not removed by a purge stream, will accumulate since it does not decompose at calcium carbonate calcination temperatures. Another area of study is the impact of the calcium sulfate concentration in the solids stream. This is directly related to the purge stream removal percentage. At higher purge stream removals, the final equilibrium concentration of calcium sulfate will be lower, but a greater amount of fresh feed will be required during every cycle. This will increase the total O&M cost in a

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commercial system. However, a lower purge stream removal amounts to a greater amount of inert sulfate being circulated throughout the system, which leads to a higher solids loading, which, in turn, leads to larger processing vessels. An optimum between process size and annual cost for sorbent must be balanced.

Finally, calcium hydroxide has a natural D50 of less than 10 microns. A proper gas-solid separator must be installed to separate the calcium hydroxide from the flue gas stream. A tradeoff occurs between efficiency, size, and power requirements. For example, a cyclone may be able to separate the calcium hydroxide from the flue gas stream, but there will be a significant pressure drop, which will require additional power from the ID fan to overcome that pressure drop. A candle filter has associated with it a lower pressure drop, but it also has a lower reliability, which will increase capital cost. A detailed economic analysis that takes all facets of the process into account must be performed. The commercial considerations will be examined further in Chapter 5.

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CHAPTER 5

COMMERCIAL CONSIDERATIONS FOR THE CCR PROCESS

5.1 Introduction

The CCR Process incorporates several reactions, with each reaction occurring within a narrow temperature range, requires numerous process reaction vessels with specific gas- solid contacting patterns, Particulate Capture Devices (PCDs), and ancillary equipment to ensure proper operation and provide continuous operation of the CCR Process.

The calcination of calcium carbonate initiates the entire CCR Process since calcium carbonate is thermodynamically the most stable calcium compound that exists under atmospheric conditions. For example, natural deposits of calcium carbonate are found in formations such as limestone and dolomite. The calcination of calcium carbonate can occur in varying reaction vessels, known as calciners or kilns, with heat being provided in either a direct-fired or indirect-fired manner. Although calciners have been well- developed for the cement industry, they are a relatively new idea for producing sequestrable CO2. Energy requirements and CO2 purity are two important criteria that will influence the Calciner, as both have a significant impact on the CCR Process economics. The calcination of calcium hydroxide, carbonation of calcium oxide, and sulfation of calcium oxide can all occur within a similar temperature range and is initiated

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by the calcination of calcium hydroxide. This allows for significant flexibility in designing the CCR Process. For example, the multiple reactions can occur in either one reaction vessel, referred to as the Carbonator/Sulfator, or multiple reaction vessels. The

Purge/Recycle can also be placed in multiple locations. The advantages and disadvantages of each scenario will be discussed in greater detail in this chapter.

Depending on the process configuration, the number of PCDs that must be added will vary. In addition, depending on location and particulate being captured, the solids loading, Particle Size Distribution, and temperature encountered in the PCD will vary widely. The stringent requirements of the PCD and its ability to be integrated into the

CCR Process will be a significant factor in the process economics.

Finally, the ancillary equipment is crucial to the operation of the CCR Process.

Additional items such as solid feeders, heat exchangers, pumps, fans, and process monitoring instruments will play an important role in the successful deployment of the

CCR Process into a coal-fired power plant. Furthermore, the ancillary equipment will increase the additional Capital and O&M costs that must be taken into account when performing an overall economic analysis.

As important as process efficiency and energy penalty are to a process, especially for a carbon capture process, the economics of a process will be the driver for determining

CCR Process equipment and configuration. In order to be competitive as a commercially- viable carbon capture technology, the CCR Process economics must, at a minimum, be equivalent to its competitors. Amine scrubbing using MEA would increase the cost of electricity by a minimum of 80% while decreasing power plant efficiency by 30%

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(Ciferno et al., 2009b). For oxycombustion, overall a 60% increase in the cost of electricity is estimated (McCauley et al., 2009; Ciferno et al., 2009b).

The Aspen simulations presented in Section 5.6 have all been optimized for minimizing the energy penalty. However, the cost of installing and operating the considerable number of heat exchangers necessary to achieve the minimal energy penalty through extensive heat integration may be economically prohibitive. Throughout the sections, the impact on the process economics will always be considered.

5.2 Overall CCR Process Description

Figure 5.1 depicts the general Process Flow Diagram (PFD) of the CCR Process integrated into a coal-fired power plant. Specific integration scenarios, which include a mass balance, energy balance, and heat integration, are provided through Aspen simulations using AspenTech Simulation Software. The assumptions, simulations, and results are located in Section 5.6.

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Waste CaO/CaCO /CaSO Clean Flue Gas 3 4 Particulate Energy CaO/CaCO / Recycle CaO/ Capture 3 Purge CaSO4 Device CaCO3/CaSO4

CARBONATOR/SULFATOR Fresh CaCO 3 Dehydration Ca(OH) CaO + H O 2 2 CALCINER Carbonation Energy Calcination CaO + CO CaCO CaCO CaO+CO 2 3 3 2 Sulfation 900 C – 1200 C CaO + SO + 1/2 O CaSO Energy 2 2 4 500 C – 650 C Flue Gas HYDRATOR Particulate Ca(OH) / CaO/ Boiler Capture 2 CaO + H O Ca(OH) CaSO 2 2 CaSO Device 4 4 500 C, P ~ 1 atm

CO2 to sequestration Fly Ash H2O

Figure 5.1 General Process Flow Diagram for CCR Process

In Figure 5.1, the blue arrows indicate the cyclic flow of solids in the CCR Process. The reactions in green are exothermic and produce high-quality heat that may be used in the coal-fired power plant’s steam turbine cycle. The reactions in red are endothermic and require energy to occur. At equilibrium conditions, the total moles of calcium removed from the Purge stream are replaced by an equivalent number of moles of fresh CaCO3. By doing so, the active Ca:C mol ratio is maintained. The active Ca:C mol ratio only accounts for the Calcium that participates in the removal of CO2 and SO2.

The basic reactions have been described in detail in Chapter 4. With the multitude of reactions, the integration options are nearly endless. However, several limitations restrict the number of feasible options that can be implemented.

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5.2.1 Existing and Future Regulations

No regulation currently exists limiting carbon dioxide emissions; however, the United

States Environmental Protection Agency (U.S. E.P.A.) has already placed regulations on several emissions produced from fossil-fuel based combustion. Any carbon capture technology must still abide by the existing regulations. The main emissions of concern with respect to implementing the CCR Process are those in place for sulfur oxides and particulate matter (PM), as those are the two primary emissions that will be affected by installing the CCR Process.

In addition, further studies, through both computational simulations and experimental verification, will be necessary to ensure that the integration of the CCR Process into a coal-fired power plant does not negatively impact the performance of downstream equipment to a significant extent. This is crucial to minimize the total economics of the

CCR Process. Process variables in the CCR Process include solids loading in the flue gas, which will be dictated by the efficiency of the PCDs and targeted CO2 removal, pressure drop through the CCR process, which will be dictated by reactor design, PCD selection, and solids loading, and volumetric flow rate and gas composition, which will be dictated by the level of reduction of CO2 and SO2 removal along with any additional compounds that may be captured by the solid sorbent.

5.2.2 Carbon Dioxide

Although CO2 is not currently regulated in the United States, several states have taken the initiative to begin drafting legislation and forming regional partnerships. For example, several states located in the western United States have formed the Western Climate

Initiative (WCI) and includes Arizona, California, Montana, New Mexico, Oregon, Utah

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and Washington, as well as parts of neighboring Canada (Western Climate Initiative,

2009a). The purpose of the WCI is to form a collaboration in order “to identify policies to reduce greenhouse gas emissions, including the design and implementation of a regional cap-and-trade program” (Western Climate Initiative, 2009b). The WCI is broad in scope and is expected to regulate the majority of greenhouse gases, with the ultimate goal of reducing emissions to 15% below 2005 levels by 2020 (Western Climate Initiative,

2009c).

Similarly, several states located in the eastern United States have formed the Regional

Greenhouse Gas Initiative (RGGI), whose goal is to reduce CO2 emissions through a cap- and-trade program. The participating states include Connecticut, Delaware, Maine,

Maryland, Massachusetts, New Hampshire, New Jersey, New York, Rhode Island, and

Vermont (Regional Greenhouse Gas Initiative, 2009). Federal CO2 emission legislation is also looming. President Obama has briefly outlined a cap-and-trade policy for controlling

CO2 emissions in the 2010 budget (White House Office of Management and Budget,

2009).

5.2.3 Sulfur Oxides

Current SO2 regulations allow a maximum emissions rate of 1.2 pounds SO2 per million

BTU’s combusted for existing electric utilities providing greater than 75 MWe (U.S.

EPA, 2009). For new plants built after 2010, it is estimated that SO2 emissions from pulverized coal plants will be as low as 0.085 pounds SO2 per million BTU’s combusted

(Klara et al., 2007). In Ohio, the most heavily mined coal seam is Pittsburgh #8 (Wolfe,

2007). With an average heat of combustion of 12,540 BTU/pound of coal and a sulfur content of 2.3%, a 500 MWe power plant with 33% thermal to electric efficiency will

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emit 82,607 tons SO2 per year (Stultz and Kitto, 1992). Given a 1.2 pound SO2 per million BTU limitation translates into a maximum emission of 28,500 tons SO2 per year for the 500 MWe power plant. Overall, only a 65% reduction in SO2 emissions would be necessary, which is well below the average SO2 removal efficiency of 90% and 95% for older and newer SO2 removal installations, respectively (Srivastava, 2000; Taylor et al.,

2005). However, if the SO2 emissions are drastically reduced to 0.085 pounds SO2 per million BTU, the 500 MWe power plant utilizing Pittsburgh #8 coal would now be required to remove 98% of the SO2, which is above the target design for existing SO2 removal technologies. As mentioned previously in Section 4.1, the sulfur removal is not an issue for the CCR Process due to the relative difference between carbon and sulfur content found in coal.

Another advantage is the high probability of also removing virtually all sulfur trioxide

(SO3), which is not regulated by the U.S. E.P.A. but can cause significant corrosion issues with power plant equipment and contribute to acid rain formation when emitted into the atmosphere (Srivastava et al., 2002; Moretti et al., 2006).

5.2.4 Particulate Matter

As a consequence of the CCR Process utilizing a solid sorbent for CO2 and SO2 removal, the solids loading and circulation in the flue gas will necessarily increase. Since no

Particulate Capture Device operates with 100% efficiency and the mean particle size of calcium hydroxide is under 10 microns, the PCD must either have a high capture efficiency for micron-sized particles at high temperatures (greater than 600 °C), which translates into increased costs due to increased power consumption due to a high pressure drop in the PCD and/or high capital cost through the utilization of a specialized gas-solid

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separation unit such as a ceramic filter, or the solids loading in the flue gas will increase

(Mississippi Lime, 2006; Graymont, 2009). The downstream PCD in the power plant, either a baghouse or electrostatic precipitator and sometimes both, must be able to handle the increased solids loading. The downstream PCD must also be able to effectively remove the micron-sized particles to a high degree.

The current primary PM emissions for coal-fired boilers built or modified after 1978 is

0.03 pounds per million BTU’s combusted (NETL, 2009). The current predictions regulate coal-fired power plants built after 2010 to emit 0.013 pounds per million BTU’s combusted (Klara et al., 2007).

For the 500 MWe coal-fired power plant consuming Pittsburgh #8 coal as feed, this amounts to 712.5 tons/year and 308.75 tons/year for coal-fired boilers built or modified after 1978 and boilers installed after 2010, respectively. If the CCR Process were installed for 90% CO2 removal using a 1.33:1 Calcium:Carbon mol ratio, the percentage of the total solids that must be retained in the power plant are 99.996% and 99.998% for the 1978 boiler and 2010 boiler, respectively. Given these percentages, the CCR PCD equipment and existing downstream PCD equipment must operate with a high efficiency to ensure that current and future PM regulations are adhered.

5.3 Calciner Considerations

Since calcium carbonate is thermodynamically more stable than both calcium oxide and calcium hydroxide, calcium carbonate is the source for Fresh Feed. Also, for this reason, the CCR Process, when applied as a post-combustion carbon dioxide capture technology, results in a carbon neutral process, with the resulting carbon dioxide eventually sequestered. Technically, the calcination of fresh calcium carbonate is not the result of

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CO2 capture from the coal combustion flue gas stream but rather the natural limestone, so the CCR Process is slightly worse than carbon neutral. Over multiple cycles, though, the

CO2 produced from the initial calcination will continually contribute an increasingly smaller fraction of the entire CO2 sequestered.

Although calciners for limestone calcination have been in existence for decades, none currently fulfill all the technical requirements necessary for a commercial calciner that would be used in the CCR Process. The minimum requirements for the Calciner include minimal energy consumption, a heat recovery system, a high conversion and reactivity of the product CaO towards hydration, and a highly concentrated stream of CO2 produced.

5.3.1 Energy Consumption

In the CCR Process, the calcination of CaCO3 is the most energy-intensive unit operation.

Due to the endothermic reaction and high temperature required for decomposing calcium carbonate, it is an absolute necessity to minimize the amount of energy the Calciner consumes while producing complete calcination. Table 5.1 on the following page lists the current commercial calciners and their energy consumptions (Kogel et al., 2006; Boateng,

2008; Derby, 2009). From Table 5.1, it can be seen that the energy consumed by the calcination process is significantly higher than is theoretically required to produce one ton of calcium oxide. The most energy-efficient calciner, the new vertical shaft kiln, still consumes greater than 20% of the theoretical minimum energy required. This amounts to wasted energy, which automatically increases the O&M costs of the CCR Process by increasing the fuel costs necessary to operate the calciner.

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Energy Consumption Maximum Capacity Calciner Type Fuel Type (x 106 kJ/ton CaO) (tons per day) Theoretical 2.7 ------Vertical Shaft Kiln, > 5.2 old Oil, 576 Natural Gas, Vertical Shaft Kiln, 3.3-3.9 Coal possible new Oil, Multiple Shaft Kiln 3.5-4.0 900 Natural Gas, Coal possible Oil, Annular Shaft Kiln 6.2 245 Natural Gas Rotary Kiln, old 12.6 Coal, Oil, 1361 Rotary Kiln, new 7.4 Natural Gas Oil, Rotary Preheater Kiln 4.7-6.3 900 Natural Gas, Coal possible Oil, Natural Gas, Rotary Hot Calciner 7.7 245 Coke, Coal possible Oil, Fluidized-bed 7.8 204 Natural Gas Oil, Flash Calciner 6.2-8.5 ------Natural Gas, Coal possible

Table 5.1 Energy requirements, calciner capacity, and fuel used for various commercial calciners

5.3.2 Heat Integration

In addition to minimizing the energy required to operate the Calciner, the energy consumption of the CCR Process can be further minimized through intelligent heat integration. As discussed in Section 4.2.2, the maximum operating temperature for the

CCR Process with 90% CO2 removal is 660 °C. The minimum temperature necessary for calcination in 1 atm partial pressure CO2 is 900 °C. Lowering the calcination temperature 56

of CaCO3 is possible by lowering the partial pressure of CO2 in the Calciner. The means of lowering the partial pressure of CO2 and its feasibility are discussed in Section 5.3.5.

To ensure 90% CO2 removal and complete calcination, thermodynamics requires a temperature below 660 °C and greater than 900 °C for carbonation and calcination, respectively. This translates into a minimum temperature increase of 240 °C for the solids exiting the Carbonator/Sulfator and entering the Calciner. A delicate balance exists between total solids circulation, Calciner feed inlet temperature, and energy required for calcination in the CCR Process. Three scenarios will be examined to further explore the relationship. The first scenario, (S.1), is the most ideal case. The second scenario, (S.2), is ideal with respect to sorbent consumption but under realistic temperatures. The third scenario, (S.3), examines the relationship between solids circulation, temperatures, and energy requirements under both realistic active Ca:C mol ratios and temperatures.

Scenario (S.1)

Under the most ideal conditions, 90% CO2 removal and 100% SO2 removal would occur at 660 °C at an active Ca:C mol ratio of 0.91163:1. Under these conditions, there would be no excess CaO circulating. The active calcium sorbent would be completely converted to CaCO3 and CaSO4. This indicates that solids circulation is at a minimum since no unreacted active sorbent, CaO, is being circulated. Figure 5.2 shows the total theoretical energy required by the Calciner, which includes the thermal energy required for heating the solids from 660 °C to 900 °C and the energy required for calcination, as a function of the Purge removal percentage. The thermal energy required for heating the solids is calculated from Equation 5-1 (Smith et al., 2001).

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T [ Eq 5-1 ] Cp B 2 2 D  τ - 1  H  dT  ATo τ 1  To τ - 1    T o R 2 T  τ   o

 is defined as T/To, where T equals the calcination temperature of 1173 K (900 °C) and

To equals the Carbonator/Sulfator operating temperature of 933 K (660 °C), R is the ideal gas constant, and A, B, and D are parameters specific to a chemical species.

Table 5.2 lists the parameters for A, B, and D for CaO, CaCO3, and CaSO4 (Smith et al.,

2001; Liu et al., 2007). To maintain consistency, A, B, and D for CaSO4 were divided by

R, in J/(mol K), since Cp was provided (Liu et al., 2007). At 900 °C, the calcination energy required is 0.1655 MJ/mol (HSC Chemistry, 2008).

Species Tmax (K) A B D

CaO 2000 6.104 4.43E-4 -1.047E5 CaCO3 1200 12.572 2.637E-3 -3.12E5 CaSO4 Not Available 9.320423 1.106E-2 -7.89E4

Table 5.2 Parameters used to evaluate energy required for heating solids

In order to simplify the analysis, the Purge stream is assumed to be located immediately after the Carbonator/Sulfator, as shown in Figure 5.1. Heat loss through the PCD is neglected, so the Purge stream is exiting at 660 °C and assumed to ideally heat exchange with the incoming Fresh Feed, which allows for the Fresh Feed to enter the Calciner at

660 °C as well. Although not realistic, the simplification still allows for a useful discussion. First, the absolute, theoretical minimum energy required is 538 MWth. Under actual conditions, the minimum energy would be higher due to non-ideal heat transfer and the requirement of a minimum temperature difference between Fresh Feed inlet 58

temperature and Purge stream exit temperature. Second, increasing the Purge stream decreases the total energy required since the Purge stream controls the circulation of

CaSO4. As the Purge stream percentage increases, the circulation of CaSO4 decreases.

Since CaSO4 is inert and does not have an active role in the CCR Process, reducing

CaSO4 circulation reduces the energy required for heating the circulating CaSO4, which then reduces the total energy required for operating the Calciner. Finally, the Calciner energy decreases towards the minimum asympotically. Beyond a 10% purge stream, minimal gains in energy savings are realized. A 20% decrease in energy consumption occurs when the purge stream increases from 1% to 10%. However, there is only slightly greater than a 2% decrease in energy consumption when increasing the purge stream from 10% to 100%. Gains in energy savings must be balanced by the cost of additional limestone required to replace the Purge. The following scenario, (S.3), highlights that relationship.

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Calciner Energy versus Purge Percentage (IDEAL) 700 Calcination Energy 650 Solids Heating 600 Total Energy

550 500 450 400

350 300

250 Calciner (MWth) Energy 200

150 100 50 0 0 10 20 30 40 50 60 70 80 90 100 Purge Percentage

Figure 5.2 Energy required for Calciner, which includes heating solids and calcining

Scenario (S.3)

Under more realistic conditions, a different trend occurs. Since it is highly improbable to operate at exact stoichiometric ratios for 90% CO2 and 100% SO2 removal, excess active sorbent will most likely be required. Using experimental results of a 1.33:1 active Ca:C mol ratio, 90% CO2 removal, and 100% SO2 removal, the Purge and Recycle stream would consist of unreacted CaO derived from Ca(OH)2 decomposition, CaCO3, and

CaSO4. In addition, the Carbonator/Sulfator temperature has been reduced to 625 °C (To

= 898 K) and the Calciner operating temperature increased to 915 °C (T = 1188 K), which are temperatures favoring their respective reactions both thermodynamically and kinetically. Figure 5.3 shows the total Calciner energy required as a function of Purge

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percentage. The ideal heat exchange between Purge solids and Fresh Feed is still assumed for simplification. Thus, fresh CaCO3 is assumed to enter the Calciner at 625 °C.

Calciner Energy versus Purge Percentage (REALISTIC) 900 Calcination Energy Solids Heating 800 Total Energy

700

600

500

400

300 Calciner (MWth) Energy

200

100

0 0 10 20 30 40 50 60 70 80 90 100 Purge Percentage

Figure 5.3 Total energy required for Calciner under realistic conditions

Since the Recycle stream now consists of CaO, which can be regenerated into active sorbent (Ca(OH)2), a minima exists with respect to Calciner energy. The Calciner energy is still a function of solids heating and energy required for calcination; however, the cycling of CaO allows for a reduction in total CaCO3 being calcined, which reduces the total energy required by the Calciner. In the ideal case presented earlier, since no CaO was being cycled, the only way to maintain a constant active Ca:C mol ratio was to maintain a constant flow of CaCO3, which required the calcination energy to remain

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constant. For a given active Ca:C mol ratio greater than stoichiometric, the presence of

CaO allows for the active Ca:C mol ratio to be maintained while varying the amount of total CaCO3 being calcined. Increasing the Purge stream percentage will initially decrease the amount of energy required for solids heating, but due to non-equivalent heat capacities amongst the solids, the energy required for heating will eventually increase as will the energy required for calcination, both of which are responsible for the continual increase in energy requirements. Initially, the rate at which energy required for solids heating decreases more rapidly than the rate at which energy required for calcination increases with increasing Purge stream removal percentage, so there exists a minima.

At a 9% purge removal, the total Calciner energy is at a minimum. However, even at the minimum, a total of 614 MWth is required as compared to the minimum 555 MWth required under the ideal active Ca:C mol ratio and identical Carbonator/Sulfator and

Calciner temperature (S.2).

Comparing the three scenarios-(S.1) Absolute Ideal Conditions for 90% CO2 removal and

100% SO2 removal at the minimum active Ca:C mol ratio and threshold temperatures,

(S.2) Ideal Conditions for 90% CO2 removal and 100% SO2 removal at the minimum active Ca:C mol ratio under realistic temperatures, and (S.3) Realistic Conditions for 90%

CO2 removal and 100% SO2 removal at experimentally-verified active Ca:C mol ratios and realistic temperatures, several variables exist that can be altered to minimize total

Calciner energy requirements, but they don’t necessarily minimize the O&M cost of the

Calciner. The following section using (S.3) further illustrates this point.

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5.3.3 Calciner Operating Costs

For providing heat to the Calciner, many types of fuels can be used. Typical fuels include coal, residual oil, and natural gas (Kogel et al., 2006; Boateng, 2008). To determine

O&M costs for operating the Calciner, prices for fuels, energy content of the fuels, and the price for Fresh Feed must be established. The prices for the fuels have been established using estimates from the first-half of 2009 for bituminous coal, industrial natural gas, and Fuel oil No. 6-residual oil (EIA, 2009a; EIA, 2009b; EIA 2009c; EIA,

2009d). Since the physical properties of the fuels have a variable range, the energy content of the fuels, along with necessary physical properties, are an approximate value that represent an average (Stultz and Kitto, 1992). The price of Fresh Feed, limestone, is approximately $25/ton (Klara et al., 2007). Table 5.3 summarizes the data.

Source Average Cost Heating Value BTU/$ MJ/$ Coal 75 $/ton 12500 BTU/lb 333,333.33 351.69 Natural Gas 5 $/1000 ft3 1032 BTU/ft3 206,400.00 217.76 Fuel Oil No. 6  (lb/gallon) 8 1.25 $/gallon 18000 BTU/lb 115,200.00 121.54 Limestone 25 $/ton

Table 5.3 Average cost and physical properties of chemicals used in calciner

From Table 5.3, coal provides the greatest amount of energy per dollar. From purely a fuel cost standpoint, the minimal O&M cost would occur using coal as the heating fuel source. Figure 5.4 shows the contribution of Fresh Feed to the O&M cost, the contribution of fuel to the O&M cost, and the total O&M cost for operating the Calciner as a function of Purge removal percentage on a one-hour time basis for the Realistic

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scenario (S.3) of a 1.33:1 active Ca:C mol ratio removing 90% CO2 and 100% SO2 at 625

°C and calcination occurring at 915 °C.

Operating Costs for Calciner versus Purge Percentage $60,000 Fuel Cost Sorbent Cost $50,000 Total Cost

$40,000

$30,000

$20,000 Cost Cost per hour

$10,000

$0 0 10 20 30 40 50 60 70 80 90 100 Purge Percentage

Figure 5.4 Operating cost of Calciner as a function of Purge removal percentage using assumptions listed in Table 5.3

The minimal O&M cost for the Calciner occurs at a Purge removal percentage of 2%, which differs from the minimal energy requirement occurring at a Purge removal percentage of 9%. Based solely on energy minimization, an additional $2,000/hour over the minimum O&M cost would be spent, which translates to $17.5 million per year.

The importance of a cost-effective carbon capture technology cannot be understated. The target goal for carbon capture technologies is a maximum increase of 35% in the cost of

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electricity, as set by the U.S. Department of Energy. (Ciferno, 2009a). Cost minimization will be crucial if the CCR Process is to be commercialized. The complexity of the CCR

Process cannot be simply optimized for minimal energy consumption and maximum overall efficiency. For this reason, all costs, both capital and O&M, for constructing and operating the CCR Process, must be considered, including products processing. For the

Calciner, processing of the CO2 stream for sequestration purposes requires compression to a minimum of 2200 psia and limitations on impurities (Ciferno et al., 2008; McGurl et al., 2005).

5.3.4 Calcination Product-Calciner Design

Depending on how heat is provided to a calciner, it is either classified as either a direct- fired or indirect-fired calciner. In a direct-fired calciner, which comprises the vast majority of calciners, the combustion gas transfers heat directly to the solids feed in either a co-current or countercurrent fashion (Boateng, 2008). In an indirect-fired calciner, the heating source is placed outside of the calciner vessel. In this fashion, the combustion fuel and gas never contact the solids feed. The indirect-fired calciner allows for the atmosphere in the calciner to remain isolated. Due to poor thermal efficiencies, though, indirect-fired kilns are used only for specialized applications (Boateng, 2008).

To obtain a pure stream of CO2, an indirect-fired calciner would be ideal. By also using an indirect-fired calciner, the CO2 stream produced from CaCO3 calcination would be immediately sequestrable. At a minimum, the CO2 stream must possess a moisture content with a dew point of –40 °C in order to prevent carbonic acid formation and pipeline corrosion during transporation (Ciferno et al., 2008). Another advantage of the indirect-fired calciner would be the product stream would be a pure stream of CO2, as

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opposed to a direct-fired calciner where additional gases would be present, thus diluting the CO2 stream and increasing costs for moisture removal and additional energy consumption for compressing a greater volume of gas. One final advantage, assuming innovation can increase the thermal heat transfer efficiency, is the overall efficiency of the CCR Process is higher when compared to direct-fired calciners integrated into a power plant. The results of the Aspen simulations, presented in Section 5.6, for three identical scenarios using an indirect-fired calciner and natural gas oxyfuel combustion direct-fired calciner consistently show the indirect-fired calciner is always superior under the given assumptions.

For calciner heating, a higher efficiency necessarily translates into lower costs. The heat for an indirect-fired calciner would utilize heat from the boiler, with the only associated cost being the cost of rerouting the ductwork to allow the flue gas to exit the boiler at a high temperature, around the calciner vessel, and a return path to the boiler at a lower temperature. For a fossil-fuel fired, oxyfuel combustion direct-fired calciner, the costs would include an Air Separation Unit (ASU), the cost of the fuel, a moisture removal system since H2O is a product of combustion, and higher costs for compression since the

ASU is neither 100% efficient in producing pure oxygen and inerts exist in the fuel, which in turn increases the volume of gas requiring compression and decreasing the purity of the CO2 stream. The underlying assumption requires the development of a high thermally efficient indirect-fired calciner, which does not yet exist.

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5.3.5 Calciner Operating Temperature

As discussed in Section 4.2.2, the temperature at which calcination of CaCO3 occurs is controlled solely by the partial pressure of CO2. Lowering the calcination temperature requires a lowering of the partial pressure of CO2, which can be performed using either a vacuum pump at the exit of the calciner or the addition of a diluent gas, both of which have been discussed in Chapter 3.

Two main benefits arise from lowering the Calciner operating temperature. First, the temperature difference between the Carbonator/Sulfator and the Calciner is lower. The solids heating, based on the scenarios (S.1)-(S.3), can range from 15% to 40% of the total energy required for the Calciner. Lowering the temperature difference between the

Carbonator/Sulfator and Calciner would lower the amount of energy required by the

Calciner, which would lower the overall O&M costs. The second benefit arises from a more reactive CaO product. At lower calcination temperatures, sorbent morphology is better maintained. By obtaining a more reactive solid sorbent, it may be possible to lower the active Ca:C mol ratio to achieve near stoichiometric carbonation/sulfation. Section

5.3.2 has already demonstrated that a lower active Ca:C mol ratio also lowers the energy required by the Calciner. Overall, a lower calcination temperature is favorable.

The advantages of a lower calcination temperature must be balanced by the cost of achieving the lower calcination temperature. The use of a vacuum pump on an indirect- fired Calciner at commercial-scale is economically prohibitive for multiple reasons. The relationship between calcination temperature and partial pressure of CO2 has been shown in Figure 4.2. A miniscule 50 °C reduction, from 900 °C to 850 °C, in calcination temperature requires approximately a CO2 partial pressure decrease of 0.5 atm. Since the

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partial pressure exponentially decreases as a function of temperature, lower temperatures will require even lower CO2 partial pressures. The energy costs required to operate the vacuum pump far outweigh any economic advantage gained by reducing the calcination temperature.

Second, if a vacuum pump were installed, the CO2 purity would decrease due to air infiltration into the Calciner. Since it is impossible to perfectly seal any vessel, placing a suction in the vessel would allow for air to infiltrate into the vessel, thus diluting the generated atmosphere. Since the Calciner is also operating at high temperatures, the material of construction for which the seals can be made becomes increasingly exotic and expensive, since the seals must withstand both high temperature and vacuum pressures.

The ancillary equipment would also become more expensive since independent atmospheres and conditions must be maintained in the Calciner and inlet and outlet connectors, all while operating continuously.

Finally, if a vacuum pump were used, the cost for compressing CO2 would increase. The

U.S. D.O.E. guideline requires CO2 compression up to 2200 psia for pipeline transportation. The CO2 exiting a typical calciner is slightly less than atmospheric, and the CO2 compression consumes approximately 10% of the total power plant’s electrical output (Ciferno et al., 2008). Since the vacuum pump lowers the exit pressure of the CO2, additional electrical energy, which translates into increased operating costs and lost electricity generation revenue, would be required for compression as compared to the case where no vacuum pump is employed.

The use of a diluent gas can be successfully employed to artificially lower the calcination temperature (Boateng, 2008). Most notably, steam has been the most popular diluent gas.

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A liquid water/steam cyclic loop would be formed with the calciner gas, as envisioned in

Figure 5.5. The Water Tank would require double the capacity of the flow rate since liquid water is required for heat exchange with the steam exiting the Calciner. The

CO2/H2O mixture exiting the Heat Exchanger may not be sufficient to reduce the gas dewpoint to –40 °C. In that case, an additional moisture removal system, whether a molecular sieve, dessicant, or condensate chiller, must be added between the Heat

Exchanger and Gas-Liquid Separator.

Water Gas-Liquid CO2 to H O Tank 2 (l) Separator Compression

CO2 H2O(l)

H2O(l) Heat Exchanger H2O(g)

CO2 H2O(g)

Solids PCD Calciner Feed

CaO

Figure 5.5 Process Flow Diagram of Calciner using steam as gas diluent

Three challenges exist when using steam. First, the CO2 stream for compression must be virtually free of moisture. Using steam requires the use of additional equipment to ensure the moisture content of the CO2 stream is acceptable. Second, water possesses a high heat

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of vaporization. A significant amount of energy is required to continuously convert the liquid water into steam. Due to heat transfer limitations and the laws of thermodynamics, the Calciner energy consumption will increase. Third, as discussed in Chapter 3, the effect of steam on product CaO is still relatively unknown.

If a diluent gas, as of yet unidentifed, were added to lower the calcination temperature, the two physical properties that would make it attractive are a high boiling point temperature and a low heat of vaporization. In order to easily separate the diluent gas from the CO2 produced from calcination, the gas mixture should be easily heat exchanged with the diluent in liquid form at ambient temperature. For example, water at 1 atm begins to vaporize at 100 °C, meaning gaseous water will begin to condense below 100

°C at 1 atm. Assuming ambient temperature is 25 °C, only a 75 °C differential exists between ambient temperature and the condensation temperature. This allows for significant amounts of moisture to still exist in the CO2 stream based on Vapor-Liquid

Equilibrium. However, if the diluent had a high boiling point (500 °C – 600 °C, for example), at ambient temperature virtually no vapor would exist in the CO2 stream since the temperature differential between liquid and vapor is so great. The second physical property would be a low heat of vaporization. The lower the heat of vaporization, the less energy is required to convert the diluent liquid into a diluent gas. The main focus when considering a diluent gas would be its effect on cost and CO2 purity.

5.4 Carbonator/Sulfator-CO2/SO2/H2O Management

The reaction vessels for calcium hydroxide calcination, sulfur dioxide removal, and carbon dioxide removal can occur simultaneously in one vessel, independently in three reaction vessels, or coupled in two reaction vessels. If all three reactions occur

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simultaneously in one vessel, the solids stream would consist of calcium oxide derived from calcium hydroxide (unless operating at the minimum stoichiometric ratio), calcium carbonate, and calcium sulfate.

Using two reaction vessels provides two options. One allows the CO2 and SO2 to be removed in the Carbonator/Sulfator while a separate reaction vessel is used for calcining calcium hydroxide. The products from the Carbonator/Sulfator would be identical to the one reaction vessel. Another possibility is to perform upstream SO2 removal while the calcium hydroxide calcination and carbonation reaction occur in the Carbonator. The first option is more advantageous and will be explored further.

If three reaction vessels are used for the three reactions, the SO2 would be removed upstream of the CO2 via Furnace Sorbent Injection. The Dehydrator would cycle the steam continuously between the Hydrator and Dehydrator while producing calcium oxide. The Carbonator products would then be calcium oxide derived from calcium hydroxide (unless operating at the minimum stoichiometric ratio) and calcium carbonate.

The Sulfator would produce calcium sulfate and unreacted solid sorbent and be removed from the flue gas stream prior to CO2 removal. However, in reality, the flue gas cannot provide enough heat to operate an indirect-fired Calciner and Dehydrator and heat the solids necessary for upstream sulfation. It is possible with a direct-fired Calciner.

5.4.1 One Reaction Vessel

If only one reaction vessel is used, the PFD is identical to Figure 5.1. The

Carbonator/Sulfator as a whole could be endothermic or exothermic, depending on the active Ca:C mol ratio. At higher active Ca:C mol ratios, the Carbonator/Sulfator requires an increasing amount of energy input to maintain isothermal operation due to calcium

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hydroxide calcination and solids heating. If the Carbonator/Sulfator requires energy input to operate isothermally, the enthalpy of incoming flue gas could be used by delivering the flue gas to the Carbonator/Sulfator at a higher temperature than the isothermal set-point temperature. Overall, the Carbonator/Sulfator should be operated adiabatically such that the enthalpy of the flue gas cooling to the isothermal temperature exactly offsets the heat required to maintain the isothermal temperature. Another possibility would be operating the Carbonator/Sulfator with heat generation under isothermal operation. To maintain the temperature, heat transfer tubes would be installed into the Carbonator/Sulfator to extract the excess heat. The challenge with operating any reaction vessel in the CCR Process with heat transfer tubes is the fouling of the heat transfer tubes that will occur and the necessity for a blowback system. Figure 5.6 shows the riser section from the 20 pph facility.

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Figure 5.6 Image of Riser section of 20 pph facility

Given the amount of solids circulating, it is not surprising that the walls and internals of the pipes become coated with a layer of solids over time. If left unmitigated, any heat transfer tubes used for heat extraction would quickly become coated as well and lead to a reduction in heat transfer rate. The ability to extract high-quality heat from the CCR

Process is the main advantage the CCR Process possesses over alternative carbon capture processes. By reducing the effective heat available for extraction due to fouling, the advantage of the CCR Process is greatly neutralized. Quantifying the rate of fouling can be achieved through the Cleanliness Factor, which is defined by Equation 5-2.

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Actual Heat Transfer Rate Cleanliness Factor = Clean Surface Heat Transfer Rate [ Eq 5-2 ]

The Heat Transfer Rate is defined by Equation 5-3.

Heat Transfer Rate  m C ΔT [ Eq 5-3 ] w p where mw is the mass flow rate

Cp is the specific heat

T is the temperature difference between inlet and outlet

By quantifying the Cleanliness Factor at different times during sorbent injection, quantifying the rate of heat transfer decrease with respect to time is possible. When the

Cleanliness Factor becomes an unacceptable value, a blowback system to clean the heat transfer tubes would be initiated (Stultz and Kitto, 1992).

5.4.2 Two Reaction Vessels

If two reaction vessels are employed, the calcium hydroxide calcines in one vessel, the

Dehydrator, while carbonation and sulfation occur in the Carbonator/Sulfator. The dehydration reaction spontaneously occurs at temperatures greater than 512 °C. By employing a Dehydrator, the steam can be looped in a cycle, which reduces the amount of water consumption and the necessity of continually generating high-temperature steam for the hydrator. The Hydrator would be charged with the necessary steam once and then make-up steam to replace steam lost through leakage would be required. The Dehydrator, in order to maintain a pure stream of H2O, which is necessary given the Hydrator operating temperature, must be heated indirectly. This also means the Dehydrator cannot operate at a lower temperature through the addition of a diluent gas.

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Figure 5.7 shows a PFD of the set-up. Compared to the one reaction vessel in Section

5.4.2, the two reaction vessels require an additional PCD and an additional reaction vessel. However, the rate of fresh steam is drastically reduced, the necessity for heat exchangers to increase the steam temperature to 500 °C is eliminated, and the volumetric flow of flue gas is reduced, which reduces the energy required by the ID fan. The advantages must be compared with the increased costs of adding additional equipment.

As mentioned previously, the energy required to operate this configuration requires the use of a direct-fired Calciner. The enthalpy contained in flue gas alone cannot provide the energy required to operate both the Calciner and Dehydrator indirectly and still maintain proper operating temperatures for the Carbonator/Sulfator.

Waste CaO/CaCO /CaSO Clean Flue Gas 3 4 Particulate CaO/CaCO / Recycle CaO/ Capture 3 Purge CaSO Device 4 CaCO /CaSO 3 4 Energy

Energy Fresh CaCO3 Energy Particulate CARBONATOR/SULFATOR Capture CaO Device CaSO4 Carbonation CALCINER CaO + CO2 CaCO3 Calcination Sulfation DEHYDRATOR CaCO3 CaO+CO2 CaO + SO2 + 1/2 O2 CaSO4 Ca(OH)2 CaO + H2O 900 C – 1200 C 500 C – 650 C 512 C – 650 C, P ~ 1 atm Ca(OH) 2 CaSO4 Flue Gas H2O HYDRATOR Particulate CaO/ Boiler Capture CaO + H O Ca(OH) 2 2 CaSO Device 4 500 C, P ~ 1 atm

CO2 to Energy sequestration Fly Ash H2O

Figure 5.7 PFD of the CCR Process with a Carbonator/Sulfator and Dehydrator

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5.4.3 Three Reaction Vessels

In the three reaction vessels option, each reaction takes place in its own independent reaction vessel. To accomplish this, the Sulfator must be located upstream of the

Carbonator, otherwise the sulfation reaction would simply occur simultaneously with the carbonation reaction. Since the Sulfator is located prior to the Carbonator, temperatures for sulfation must be greater than the Carbonator operating temperature.

Furnace Sorbent Injection (FSI) is a process by which sulfur is removed via solid sorbents, such as limestone or calcium hydroxide, at temperatures greater than the

Carbonator operating temperature. Operating between 750 °C and 1250 °C, FSI utilizes either limestone or calcium hydroxide to remove the SO2 (Adams et al., 2006). The advantage gained by operating with an upstream SO2 removal system is the elimination of the Purge stream since there will be no circulation of calcium sulfate in the

Carbonator. In essence, upstream SO2 removal behaves as the Purge stream. A Fresh

Feed stream will still be required to replace the sorbent required for SO2 removal.

Placing the Sulfator upstream of the Carbonator requires several design considerations, including sorbent composition and PCD arrangement. Overall, the drawbacks of upstream SO2 removal negate any advantage. For SO2 removal, calcium hydroxide is significantly higher in reactivity than calcium oxide (Bruce et al., 1989; Al-Shawabkeh et al., 1994). At a Calcium:Sulfur (Ca:S) mol ratio of 2, calcium hydroxide is capable of

30% to 50% SO2 removal while calcium oxide derived from calcined calcium carbonate can only achieve 20% to 35% SO2 removal (Rostam-Abadi et al., 1990; General Electric,

2003; Adams et al., 2006). Without complete removal of SO2 upstream of the

Carbonator, the Carbonator exit would contain calcium sulfate, which would then require

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a Purge stream and negating the one advantage of upstream SO2 removal. Assuming a linear extrapolation, calcium hydroxide would require a Ca:S mol ratio between 4 and 7, while calcium oxide derived from calcium carbonate would require a Ca:S mol ratio between 6 and 10. Since the sorbent cannot be separated from the calcium sulfate and the

Sulfator solids cannot be cycled through the entire CCR Process without introducing calcium sulfate into the CaO/CaCO3 Carbonator loop, the Sulfator solids must either operate on a once-through cycle or introduce a secondary sorbent/CaSO4 loop that cycles independent of the CaO/CaCO3 loop. Figure 5.8 shows one possibility of upstream SO2 removal with the sorbent used on a once-through basis. Figure 5.9 shows a configuration where the CaSO4 is cycled.

Particulate Flue Gas SULFATOR Capture Boiler Sulfation Particulate CaO/ Device CaO + SO + 1/2 O CaSO Capture CaSO 2 2 4 Device 4 750 C – 1250 C Fly Ash Energy

SPLITTER H O 2 Flue Gas Energy HYDRATOR

CaO + H O Ca(OH) 2 2 500 C, P ~ 1 atm CARBONATOR Carbonation CO to Particulate Energy CaO + CO CaCO 2 Capture 2 3 sequestration Ca(OH)2 Device 500 C – 650 C DEHYDRATOR Energy Ca(OH) CaO + H O CaO 2 2 CaO 512 C – 650 C, P ~ 1 atm CALCINER Fresh CaCO3 Calcination CaO / CaCO CaCO CaO + CO 3 3 2 900 C – 1200 C

Figure 5.8 SO2 removal on a once through CaSO4 loop

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Particulate Flue Gas SULFATOR Flue Gas Boiler Capture Particulate Sulfation Device CaO + SO + 1/2 O CaSO Capture 2 2 4 Device 750 C – 1250 C

Fly Ash Energy CaO/CaSO4 PURGE/ CaO/ SPLITTER RECYCLE H2O CaSO4 HYDRATOR CaO/ CaSO4 CaO + H2O Ca(OH)2 500 C, P ~ 1 atm CARBONATOR Carbonation CO to Particulate Energy CaO + CO CaCO 2 Capture 2 3 sequestration Ca(OH)2 Device 500 C – 650 C DEHYDRATOR Energy Ca(OH) CaO + H O CaO 2 2 CaO 512 C – 650 C, P ~ 1 atm CALCINER Energy Fresh CaCO3 Calcination CaO / CaCO3 CaCO3 CaO + CO2 900 C – 1200 C

Figure 5.9 Recycle loop for CaO/CaSO4 for upstream SO2 removal

In either instance, the PCD must operate with near 100% gas-solid separation efficiency, otherwise calcium sulfate will be circulating in the CaO/CaCO3 loop. Also the Splitter location will differ if the sorbent used is calcium hydroxide instead of calcium oxide derived from calcium carbonate or fresh calcium carbonate itself.

Regardless of sorbent, costs of the CCR Process will increase with upstream SO2 removal. Not only does the complexity of the process increase but also the total sorbent consumed and energy consumption increase. Since high Ca:S mol ratios are required to ensure complete SO2 removal, and the sorbent cannot be utilized for CO2 removal, the unreacted sorbent is essentially useless waste.

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A tradeoff occurs between sorbent type and energy requirement that balance each other.

If calcium oxide derived from calcium carbonate is taken after the Calciner, a high Ca:S mol ratio is required, which means additional Fresh Feed is required to replace the SO2 removal sorbent. Since the SO2 sorbent is at high temperature already, energy is not required to heat the sorbent, but additional energy is required to calcine the Fresh Feed. If calcium hydroxide is the SO2 removal sorbent, a lower Ca:S mol ratio is required compared to utilizing CaO; however, the calcium hydroxide would require heating, which requires energy. In the end, no advantage is really gained from upstream SO2 removal, yet it remains a possible option.

The usefulness of upstream SO2 removal concludes with the waste solids. If there are niche end uses for a CaO/CaSO4 mixture, or CaO/CaCO3, or non-reactive CaO, or even a flyash/CaO/CaSO4 mixture, but not a CaO/CaCO3/CaSO4 mixture, then upstream SO2 removal would become a true option. A flyash/CaO/CaSO4 mixture could occur by eliminating the flyash PCD and capturing it along with the CaO/CaSO4. The final end location of the solids mixture will determine the usefulness of upstream SO2 removal.

5.5 Purge/Recycle and PCDs

The Purge/Recycle stream can be placed between virtually any two pieces of process equipment. The final location will be highly dependent on the process configuration and operating temperatures. One possibility that will be discussed in theory is in a PCD itself.

5.5.1 Purge/Recycle Physical Location

In the one reactor vessel option, the Purge/Recycle can be placed after the

Carbonator/Sulfator, as shown in Figure 5.1, after the Calciner, or after the Hydrator. By placing the Purge/Recycle after the Carbonator/Sulfator, the Fresh Feed can be heated

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near the operating temperature of the Carbonator/Sulfator. In addition, the minimum amount of Fresh Feed is added since the Fresh Feed stream directly enters the Calciner. If the Purge/Recycle stream were located after the Calciner, the Fresh Feed could be heated to near the operating temperature of the Calciner. Another advantage would be the Purge stream would only contain CaO and CaSO4. However, a greater amount of Fresh Feed as compared to the minimum would be required since the Purge stream is removing calcined

Fresh Feed. The worst location for placing the Purge stream is after the Hydrator since the Purge Stream would be removing active sorbent. Based on the advantages and disadvantages of each location, the least obstructive location for the Purge/Recycle is after the Carbonator/Sulfator since the Purge stream is not removing calcined Fresh Feed, which requires energy for calcination and becomes active sorbent.

5.5.2 Purge/Recycle Theoretical Location

The Purge/Recycle stream requires a solids separation equipment based on mass. An example would be a two-screw screwfeeder connected to a hopper. The two screws would rotate at different speeds, with their proportions such that the Purge stream exit is the proper percentage of the total inlet feed.

Since no experiments have yet been conducted to verify PSDs or PCD removal efficiencies, the following possibility is still completely theoretical. The natural D50 of calcium hydroxide is less than 10 microns, with a maximum particle size less than 250 microns. Assuming the calcium oxide derived from calcium hydroxide, calcium carbonate, and calcium sulfate exiting the Carbonator/Sulfator does not increase or decrease dramatically in particle size, the PSD of the solids exiting the

Carbonator/Sulfator would be similar to that of calcium hydroxide. The Fresh Feed

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calcium carbonate can have a variable PSD, ranging from submicron similar to calcium hydroxide up to granules that are millimeters in size. The grinding process of the Calcium

Carbonate dictates the Fresh Feed PSD. In the Calciner, it is possible to obtain a binary

PSD for the solids. The Carbonator/Sulfator solids will be similar to calcium hydroxide while the Fresh Feed can be variable. As an example, Graymont ground lime has a D50 of

600 microns (Wong, 2007). The PCD located at the exit of the Calciner separating the

CO2 stream from the solids entering the Hydrator can be designed to also perform the function of the Purge. Assuming the Fresh Feed is orders of magnitude larger than the solids exiting the Carbonator/Sulfator, the PCD would separate the Fresh Feed particles with ease, with the assumption that massive attrition of the Fresh Feed is not occurring in the Calciner. With respect to the solids from the Carbonator/Sulfator, the PCD can be designed in such a fashion that its separation efficiency is equivalent to the desired Purge percentage. Realistically, PCDs have separation efficiencies based on size while the

Purge percentage is mass based. Also, no PCD performs with 100% gas-solid separation efficiency. Given a large enough margin size distribution between the Fresh Feed and the solids exiting the Carbonator/Sulfator, and assuming the PSD of the solids exiting the

Carbonator/Sulfator are similar to that of calcium hydroxide, statistically the solids exiting the Calciner should be separable such that the entire mass of Fresh Feed enters the

Hydrator, while a fraction of the solids exiting the Carbonator/Sulfator enters the

Hydrator. The calcined Fresh Feed, once hydrated, will be converted into calcium hydroxide and naturally become a fine powder through the hydration process. The critical assumption that must be experimentally verified is the PSD of the solids exiting the

Carbonator/Sulfator is similar to that of calcium hydroxide. If that assumption holds, then

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the PCD efficiency, which separates based on particle size, will be statistically removing the identical percentage, by mass, of each solids component. In that regard, the PCD functions as the Purge. The advantage is the removal of a PCD and a splitter, but the disadvantage is the solids flow into the Calciner increases, and the solids must still be removed from the CO2 stream. Also, the Fresh Feed cannot be preheated if the PCD functions as the Purge. Since the solids will be in a hot gas stream, whose heat is utilized to either generate steam for the steam turbine cycle, or pre-heat other existing streams, the solids are unavailable for Fresh Feed preheating.

The concept of placing the Purge in a PCD by capitalizing on its inefficiency is plausible only if the PSD of the solids in the Purge stream are similar to each other. The PCD located after the Carbonator/Sulfator could also be another location for the Purge. Based solely on temperature, higher efficiency PCDs are generally less expensive at lower temperatures, so it may be more economical to decrease the separation efficiency of the

Calciner PCD to lower costs since the Calciner operates at a temperature greater than the

Carbonator.

5.5.3 Particulate Capture Devices

Several PCDs have been shown in the previous PFDs. The first one is for flyash.

Although prior research has established that flyash poses no performance issues in the

CCR Process, flyash circulation will increase solids circulation and may pose problems for solids disposal or unwanted chemical reaction in the Calciner or Hydrator. In the operating temperature of the Calciner, calcium oxide can react with alumina (Al2O3), silica (SiO2), and iron oxide (Fe2O3) (Boateng, 2008). Depending on the ash composition, all are possible compounds (Stultz and Kitto, 1992; Wong, 2007). An additional problem

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with flyash circulation is its inhibition effect on the calcination rate (Huang and

Daugherty, 1988). With respect to temperature, any PCD should, for ease of operation and lower costs, operate at the lowest operating temperature possible.

A series of contradictions actually complicate the remaining PCDs. Section 5.2.4 detailed the efficiency the PCDs are required to achieve to comply with existing regulations. The high efficiency separations are associated with what are considered high pressure drops across the PCD. To overcome the lost pressure, the power of the ID fans must increase, which then increases cost. To further add to the challenge of minimizing cost are the temperatures, which are advantageous for heat generation, but limit the PCD selection.

The proper equipment to allow for existing compliance at a minimal cost must be designed. The overall design will be plant specific.

5.5.4 Final Considerations

Depending on the integration option, small advantages that work in favor of the CCR

Process are the savings in pressure drop by the flue gas not requiring a post-combustion

SO2 removal system and the lower volumetric flow of flue gas since CO2, SO2, and O2

(required for sulfation) have been converted into solid CaCO3 and CaSO4, respectively.

The lower volumetric flow of flue gas is only realized if a dedicated Dehydrator is located between the Hydrator and Carbonator/Sulfator. Otherwise, the dehydration of calcium hydroxide in the carbonator increases the total volumetric flow based on a 1.33:1 active Ca:C mol ratio, which is required for 90% CO2 removal. An average pressure drop through the most common type of SO2 removal technology, the wet scrubbing forced oxidation of limestone, is approximately 5 inches of water (Stultz and Kitto, 1992). The

500 MWe reference plant would produce approximately 1.1 million Standard Cubic Feet

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per Minute (SCFM) of flue gas, which assumes complete combustion and 20% excess air. A 90% CO2 removal, 100% SO2 removal, and stoichiometric O2 for complete conversion to CaSO4 will reduce the flue gas volumetric flow rate to 0.956 million

SCFM, which is a reduction of 13%. However, without a dedicated dehydrator, the flue gas volumetric flow rate would increase by 5.9%, assuming a 1.33:1 active Ca:C mol ratio.

5.6 Aspen Simulations

The complexity of the CCR Process and the integration options that exist require evaluation. Manual calculations, although possible, would be time-consuming and require simplifying assumptions that could significantly alter the final results. Aspen Plus is a computer software program that models chemical processes. With an extensive thermodynamics and physical properties database, ability to model a wide selection of unit operations, and equation-oriented calculations, Aspen Plus has the ability to quickly and accurately perform mass balances, energy balances, and heat integration. Heat integration, which is crucial to the success of the CCR Process, cannot be easily performed through manual calculations on a complex process-for this reason, Aspen Plus is a key tool for evaluating the multiple CCR Process operations.

5.6.1 Overall Modelling Parameters

In order to properly simulate the CCR Process, chemical species that are present or may be formed must be specified. The source of coal was a typical Pittsburgh #8, and its properties are listed in Table 5.4 (Stultz and Kitto, 1992). Table 5.5 details the chemical components and the method in which Aspen processes the component. The stream class, which determines how each chemical component interacts thermodynamically, chosen is

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MIXCINC. MIXCINC constrains the components defined in a stream class. The thermodynamic behavior of Conventional components are defined by vapor-liquid-solid phase equilibrium. The database determining the vapor-liquid-solid phase equilibrium is

IDEAL, which is a reasonable choice since no electrolytes, aqueous solutions, polymers, or extreme conditions occur in the CCR Process. Solids components are specified for those that can participate in solid chemical equilibrium. Nonconventional components are given to species that have variable physical properties, which the user must input (T-

Raissi et al., 2007). The properties of nonconventional components such as coal and ash are calculated based on the HCOALGEN and DCOALIGT model, respectively.

Proximate Weight % Weight % Ultimate Wt% Weight% Analysis As-Received Dry Analysis As-Received Dry Moisture 5.2 Moisture 5.2 Fixed Carbon 48.1 50.7 Ash 8.6 9.1 Volatiles 38.1 40.2 Carbon 70.2 74 Ash 8.6 9.1 Hydrogen 4.8 5.1 Nitrogen 1.5 1.6 HHV Chlorine 0 0 12,540 13,227.8 (BTU/lb) Sulfur 2.2 2.3 Oxygen 7.5 7.9

Table 5.4 Properties of Pittsburgh #8 Coal

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Component ID Type Component name Formula COAL NC ASH NC CAO SOLID CALCIUM-OXIDE CAO CACO3 SOLID CALCIUM-CARBONATE-CALCITE CACO3 CASO4 SOLID CALCIUM-SULFATE CASO4 CAOH2 SOLID CALCIUM-HYDROXIDE CA(OH)2 C SOLID CARBON-GRAPHITE C CO CONV CARBON-MONOXIDE CO CO2 CONV CARBON-DIOXIDE CO2 N2 CONV NITROGEN N2 NO CONV NITRIC-OXIDE NO

NO2 CONV NITROGEN-DIOXIDE NO2 O2 CONV OXYGEN O2 H2 CONV HYDROGEN H2 HCL CONV HYDROGEN-CHLORIDE HCL CL2 CONV CHLORINE CL2 S CONV SULFUR S SO2 CONV SULFUR-DIOXIDE O2S SO3 CONV SULFUR-TRIOXIDE O3S H2O CONV WATER H2O CH4 CONV METHANE CH4 C2H6 CONV ETHANE C2H6 C3H8 CONV PROPANE C3H8 N-BUT-01 CONV N-BUTANE C4H10-1

Table 5.5 Chemical species identified for simulating CCR Process using Aspen

The physical parameters for Natural Gas follow the Department of Energy guidelines

(McGurl et al., 2005). Table 5.6 lists the composition.

Chemical Volume Stoichiometric Oxygen Requirement Component Formula Percentage for Complete Combustion Methane CH4 93.1 2 Ethane C2H6 3.2 3.5 Propane C3H8 0.7 5 n-Butane C4H10 0.4 6.5 Carbon CO 1.0 Inert Dioxide 2 Nitrogen N2 1.6 Inert

Table 5.6 Composition of Natural Gas 86

A Higher Heating Value (HHV) of 38.46 MJ/m3 was used, per the guidelines (McGurl et al., 2005). Conversion using the ideal gas law translates the HHV into 940.4 MJ/kmol under standard conditions (1 atm pressure and 298 K temperature).

In order to properly model the CCR Process, parameters for each specific unit operation must be specified. Since not all reactions occur under stoichiometric conditions or completely conform to thermodynamic predictions, the type of reaction vessel chosen is specific to the conditions experienced and experimental results.

Unit operations not modeled include the steam turbine cycle of the power plant (due to its complexity), CO2 compression (due to widely varying compression routes), and power plant equipment downstream of the CCR Process (NOx removal, mercury control, flue gas polishing units, …). The Air Separation Unit (ASU) was simplified from a cryogenic distillation unit to a simple fractional separator. Although not properly modeled, the energy requirement associated with each was taken into account. Table 5.7 lists the major unit operations and their modeling technique.

The basis of comparison for each Aspen simulation is the Energy Penalty, or Parasitic

Energy Consumption, incurred by each integration option. The Energy Penalty is defined as the reduction in net electrical generation due to the installation of carbon capture equipment, expressed as a percentage. Equation 5-4 shows the mathematical formula

(Page, 2009).

 Power Output wit hout Carbon Capture - Power Output wit h Carbon Capture  [ Eq 5-4 ] Energy Penalty  100   Power Output wit hout Carbon Capture 

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Unit Operation Aspen Plus Model Comments / Specifications Virtually decompose coal to elements Coal Decomposition R-Yield (Pre-requisite step for coal combustion) Thermodynamic model of coal Coal Combustion R-Gibbs combustion Adiabatic flame temperature calculated 1.33:1 active Calcium:Carbon mol ratio and total Carbon content in coal, 90% Carbonator/Sulfator Rstoic CO2 removal, 100% SO2 removal. Operating temperature = 625 °C Purge Fsplit Splits streams based on mass fraction Combines Recycle stream and Fresh Recycle Mixer Feed stream in terms of material and heat Models limestone calcination-Operating Calciner Rstoic Temperature = 915 °C and 100% extent of reaction Hydrator R-Gibbs Operating temperature = 500 °C

Gas-Solid Separation Ssplit 100% separation efficiency, generic Unspecified heat exchange medium for Heaters and Coolers Heater cooling and heating Heat exchange gas-solid streams- Heat Exchangers MheatX Operates with minimum 10 °C approach temperature Electrical Power 42% efficiency for thermal to electric Adams et al., 2007 Generation conversion in steam turbine Net Electrical Generation 10% in-house electricity consumption 119 kWh electricity/tonne CO to CO Compression Wong, 2005 2 2 compress to 14 MPa 95% separation, 160 kWh/tonne O ASU 2 produced, cryogenic distillation

Table 5.7 List of Aspen Plus unit operations and references for non-modeled units

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5.6.2 Assumptions and Operating Conditions

The assumptions and operating conditions for each unit operation are listed below:

 Ambient Temperature = 25 ° C  Ambient Pressure = 1 atm  Coal is combusted in the boiler as-received  Coal feed rate of 205 tons/hour of Pittsburgh #8 coal (1507.89 MWth)  Coal-fired boiler with 20% excess air (Stultz and Kitto, 1992)

 Air consists of 21% volume O2 and 79% volume N2  1.33:1 active Calcium:Carbon mol ratio  Carbonator/Sulfator operates at 625 °C  Upstream Sulfator, if present, operates at a 5:1 active Calcium:Sulfur mol ratio

 90% volume CO2 removal and 100% volume SO2 removal  Purge stream set to 3% weight  Hydrator operates at 500 °C  1:1 Calcium:Steam mol ratio  Steam extracted from exit of low-pressure steam turbine  Dehydrator, if present, operates at 530 °C  Calciner operates at 915 °C

 Energy for CO2 compression to 14 MPa is 119 kWh/tonne CO2 (Wong, 2005)  Economizer flue gas temperature = 350 °C (Adams et al., 2007)  All heat exchangers have a minimum of 10 °C between hot inlet and cold outlet  All PCDs operate with an ideal separation efficiency of 100%  Heat extraction from CCR Process is 90% efficient

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5.6.3 Coal-fired Power Plant without Carbon Capture

The reference plant for the CCR Process is a 507.6 MWe coal-fired power plant without carbon capture technology. The Combustion Air is split into two streams, Primary Air and Secondary Air. Primary Air is 15.8% of the total Combustion Air. The Primary Air, which is used to partially dry the coal as it is conveyed into the boiler, is heated to 204 °C

(400 °F) while the Secondary Air, which is used for coal combustion, is heated to 316 °C

(600 °F). The heat for the air is provided through a heat exchanger, the Air Preheater

(APH), which cools the flue gas exiting the economizer from 350 °C down to 111 °C.

The exiting flue gas is then emitted to the atmosphere through the stack. Coal, which enters the Boiler at ambient temperature, is combusted with the Air to generate a flue gas stream exiting the Boiler at its adiabatic flame temperature of 1988 °C. The volume percentages of the flue gas components, minus the ash, is listed in Table 5.8. The trace gas components are also listed in parts per million (ppm) in parentheses.

Component Volume Percentage Nitrogen ( N2 ) 75.15 Carbon Dioxide ( CO2 ) 13.49 Water ( H2O ) 6.45 Oxygen ( O2 ) 3.44 Carbon Monoxide ( CO ) 6.48 x 10-1 (6,475 ppm) Nitrogen Oxide ( NO ) 6.05 x 10-1 (6,045 ppm) -1 Sulfur Dioxide ( SO2 ) 1.64 x 10 (1,645 ppm) -2 Hydrogen ( H2 ) 5.70 x 10 (570 ppm) -4 Nitrogen Dioxide ( NO2 ) 2.63 x 10 (2.6 ppm) -5 Sulfur Trioxide ( SO3 ) 8.94 x 10 (0.89 ppm) Sulfur ( S ) 1.91 x 10-6 (0.02 ppm)

Table 5.8 Flue gas components from Aspen simulation and their concentration

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The flue gas is cooled from its adiabatic flame temperature of 1988 °C to 350 °C. For the purpose of these simulations, the economizer (ECON) represents the heat exchangers in the boiler, which include the superheaters, reheaters, and economizer. Upon cooling

1342.8 MWth is recovered. Since the coal input thermal energy totaled 1506.79 MWth, the boiler heat efficiency is 89%. Assuming a 42% efficiency from thermal energy to electric energy in the steam turbine, the gross electric generation is 563.98 MWe. With

10% of the electricity generated used within the power plant, the net electric generation is

507.58 MWe. The overall efficiency, defined as net energy output divided by the total energy input, of the power plant is 33.7%. Figure 5.10 is the complete Aspen simulation of the reference plant with no CO2 control system.

No CO2Control Temperature (C) Pittsburgh #8 Coal, 1506.79 MWth input S T E A M T U R B I N E Pressure (atm) 89% Boiler Efficiency Mass Flo w R ate (to ns/h r) 1342.8 MWth from B oiler Q 42% Turbine efficiency 563.98 MWe Gross Du ty (MW ) HEAT-1 1342.8 10% in-house consumption Q Du ty (MW ) 507.58 MWe Net ECON

Q=-1343 350 25 316 25 1988 1 1 1 1 1 QBOIFLU 2535 205 BOI-FLU 2535 1962 205

B-BURNIN BOILER SEC-OUT COAL DECOMP APH COAL QB-BURN PRI-OUT 111 1 Q=48 Q=-48 25 Q=-165 205 tons/hour FLU-OUT 2535 -48.1 204 1 1 368 PRI-IN 25 368 1 SEC-IN 1962 CO2/SO2 Lean AIRSPLIT Flue Gas to Stack 25 1 2330

AIR AIR 2330 tons/hour

Figure 5.10 Reference coal-fired power plant with no CO2 control and 507.58 MWe net

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The next two sections will investigate CCR Process integration scenarios. The power plant modeled in this section is the basis of comparison for all Energy Penalty calculations.

5.6.4 CCR Process Integration with an Indirect-fired Calciner

The three CCR Process integrations in this section utilize the specific heat from the flue gas to provide the energy required to maintain the Calciner at 915 °C. The flue gas is first cooled from its adiabatic flame temperature of 1988 °C to a lower temperature, Ta. The heat is used to generate steam for the steam turbine. The Ta flue gas is then cooled to a lower temperature Tb, with the energy used to provide heat for the Calciner. The temperatures Ta and Tb are determined by the amount of energy required by the Calciner, with the thermal energy upon cooling the flue gas from Ta to Tb equivalent to the amount of energy required by the Calciner. The minimum temperature of Tb is 930 °C, which is

15 °C above the Calciner operating temperature.

Cyclical loops, such as hot air preheating the Fresh Feed, steam hydration/dehydration, and the Recycle stream with upstream sulfation, are modeled at steady-state by setting the inlet conditions equal to the outlet conditions with respect to temperature, pressure, and mass flow. Start-up and shut-down of the CCR Process is not discussed although the procedure will be an important factor in the commercial-scale process. Also not modeled in Aspen are the exact number of reaction vessels. Even though the Calciner, Hydrator, and Carbonator/Sulfator, and PCDs are modeled as one vessel, in reality, there will be numerous, identical vessels in parallel or series. This occurs from the vast amounts of gas and solid that must be processed per hour.

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Integration Option 1, shown in Figure 5.11, is the most ideal integration option with respect to energy penalty. The CCR Process begins with the calcination of limestone, which is idealized to 100% calcium carbonate (CaCO3). At steady-state, entering the calciner (SORBMIX) is 1058 tons per hour (tph) of limestone (CaCO3), 303 tph of lime

(CaO), and 613 tph of calcium sulfate (CaSO4). 62 tph of fresh CaCO3 (C-FRESH) enters at 558 °C to replace the Purge solids removed (PURGE). The hot Purge (H-PURGE) stream is cooled from 625 °C to 60 °C in a heat exchanger that heats air from 50 °C to

610 °C (AIR1). The Purge stream is then disposed at 60 °C (PURGE). The hot air (AIR2) at 610 °C is cooled to 50 °C in a heat exchanger that heats the fresh CaCO3 from 25 °C to

558 °C. The air is continuously cycled between the two heat exchangers. The Calciner is indirectly-heated by passing hot flue gas around the Calciner, thus cooling the flue gas from 1680 °C to 930 °C. The products from the Calciner are calcium oxide, calcium sulfate, which is stable at the Calciner conditions, and carbon dioxide. A gas-solid separator produces 466 tph of pure, dry CO2 at 915 °C (CAL-CO2), which is sequestration-ready, and a solids stream consisting of 897 tph of CaO and 613 tph of

CaSO4. The CO2 stream is cooled to 350 °C, with the heat used to generate steam for the steam turbine. The solids stream is cooled to 500 °C, with the heat used to generate steam for the steam turbine, while 288 tph of steam, assumed to be saturated, from the exit of the low-pressure turbine discharged at 38 °C (LPSTEAM) is heat exchanged with the 350

°C CO2 to increase the steam temperature to 289 °C. Steam and the solids enter a hydrator operating at 1 atm and 500 °C. The resulting heat from the exothermic reaction is used to generate steam for the steam turbine. The resulting product stream (CAOH2) consists of 1185 tph of Ca(OH)2 and 613 tph of CaSO4.

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The flue gas is further cooled from 930 °C to 629 °C, where the flue gas enters an ash

PCD to remove the flyash, which produces 18 tph. The flyash-free flue gas enters the

Carbonator/Sulfator and is contacted with the solid sorbent stream consisting of Ca(OH)2 and CaSO4. The Carbonator/Sulfator is operated at 625 °C with the endothermic reaction of calcium hydroxide decomposition balanced by the exothermic reactions of carbonation and sulfation. The resulting gas-solid stream is then sent to a gas-solid separator to separate the CO2-SO2 lean flue gas from the solids. The flue gas is cooled to 350 °C, with the heat used to generate steam for the steam turbine, and then delivered to the APH. The solids stream consists of CaO derived from Ca(OH)2 decomposition, CaCO3, and CaSO4.

Without a continuous purge stream/fresh feed stream, the ultimate result would be a system consisting solely of CaSO4 since it is not regenerable. The purge stream is used to remove a small fraction of the solids such that CaSO4 will be circulated, but not accumulated at equilibrium. The molar quantity of the calcium removed from the purge stream is replaced by fresh CaCO3 to maintain a constant active Calcium:Carbon mol ratio. The overall Energy Penalty is 15.1%. The 431.0 MWe net electricity generation includes 50.3 MWe for CO2 compression, which is common to all post-combustion CO2 capture systems. For carbon capture alone, the CCR Process Energy Penalty is roughly

5% and the remaining 10% arising from CO2 compression.

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Figure 5.11 Integration Option 1-Indirect fired calciner in one reaction vessel

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Integration Option 2, shown in Figure 5.12, is a 2 reaction vessel that cycles the steam through the Hydrator/Dehydrator. The Hydrator/Dehydrator is assumed to operate continuously without any loss of mass through leaks or reaction vessel inefficiencies.

The mass flows are identical to those in Option 1. Except for the temperature of the flue gas prior to the Carbonator, the temperature profiles are nearly identical to those in

Option 1. The Calciner is indirectly-heated by passing hot flue gas over the Calciner, thus cooling the flue gas from 1890 °C to 1153 °C. The temperature of the flue gas is higher than in Option 1 because the energy from the flue gas is also required for the Dehydrator.

The operating temperature of the Dehydrator is 530 °C, which is roughly 20 °C greater than the thermodynamic minimum. The Dehydrator is also indirectly-heated by passing the flue gas over the Dehydrator, thus cooling the flue gas from 1153 °C to 625 °C. At steady-state, steam enters the Hydrator at 1 atm and 530 °C, while the solids enter at 500

°C. The resulting heat from the exothermic reaction is used to generate steam for the steam turbine.

The Carbonator/Sulfator operating temperature is 625 °C. Since the calcination of calcium hydroxide is performed in a separate reactor, the Carbonator/Sulfator generates heat from the exothermic reactions and is used to generate steam for the steam turbine.

Besides the flue gas stream prior to the Carbonator, the only other temperature difference occurs with the CO2 stream exiting the Calciner. The CO2 stream is cooled from 915 °C to 105 °C to generate steam for the steam turbine. The overall Energy Penalty is 16.0%.

The 426.4 MWe net electricity generation includes 50.3 MWe for CO2 compression.

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Figure 5.12 Integration Option 2-Indirect fired Calciner with 2 vessels

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Integration Option 3, shown in Figure 5.13, is a 2 reaction vessel that performs SO2 removal prior to CO2 removal. The assumptions used in Options 1 and 2 are maintained.

The Calciner is indirectly-heated by passing hot flue gas over the Calciner, thus cooling the flue gas from 1639 °C to 930 °C. At steady-state, entering the Calciner is 1099 tph of

CaCO3 and 319 tph of CaO (SORBMIX). A continuous Fresh Feed of 70 tph (C-FRESH) replaces the sorbent consumed by the SO2 removal. The increase in Fresh Feed reflects the excess sorbent consumed by the SO2 removal. The products of calcination include

935 tph of CaO (CAO) and 483 tph of CO2 (CAL-CO2), both increases over Options 1 and 2. The CaO is cooled to 500 °C, with the heat being used to generate steam for the steam turbine. The CO2 is first cooled to 350 °C to generate heat for the steam turbine.

The CO2 is further cooled through heat exchange with steam from the exit of the low pressure turbine. 300 tph of low-pressure steam at 38 °C (LPSTEAM) is heated to 288

°C, while the CO2 is cooled to 50 °C and then compressed for sequestration.

At steady-state, steam enters the Hydrator at 1 atm and 289 °C, while the solids enter at

500 °C. The resulting heat from the exothermic reaction is used to generate steam for the steam turbine. The calcium hydroxide is split into two streams (HYD-SPLIT). The splitter operate such that a 5:1 Ca:S mol ratio is maintained in the Sulfator, while a 1.33:1

Ca:C mol ratio is maintained in the Carbonator.

The flue gas is cooled from 1988 °C to 858 °C. The flue gas is then sent to a PCD to remove the flyash. The flyash-free flue gas then enters the Sulfator, where the SO2 present in the flue gas completely reacts with 52 tph of Ca(OH)2 at 850 °C. To maintain

850 °C in the Sulfator requires a higher inlet flue gas temperature due to solids heating and calcium hydroxide calcination. The Sulfator products enter a PCD operating at 850

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°C. Since the solid products of sulfation function as the Purge stream, the maximum amount of Fresh Feed preheat occurs when the PCD operates at the highest temperature possible. Otherwise, it is possible to place the PCD prior to the Carbonator, which operates at 625 °C. Again, the Aspen simulation has been optimized for minimum energy penalty using the given assumptions, and cost has not played a factor in the simulation.

The SO2-free flue gas is further cooled to 622 °C prior to entering the Carbonator. The heat is used to generate steam for the steam turbine. In the Carbonator, 1183 tph of

Ca(OH)2 at 500 °C (CO2-SORB) reacts with the CO2 in the flue gas to remove 90% of the CO2. The Carbonator operates isothermally at 625 °C and does not generate heat. The exothermic reaction of carbonation balances the endothermic reaction of Ca(OH)2 calcination and the incoming flue gas is slightly heated, but overall no heat is output by the Carbonator.

The overall Energy Penalty is 17.9%. The 416.8 MWe net electricity generation includes

52.2 MWe for CO2 compression. Clearly obvious is the decrease in net electricity and increase in Energy Penalty when performing upstream SO2 removal. Without solid justification for performing upstream SO2 removal, the added complexity and higher equipment operating temperature do not warrant its use.

A three reaction vessel option, with a dedicated Dehydrator, Carbonator, and Sulfator was not modeled for an indirect-fired Calciner. The energy consumption and temperature constraints of each process vessel proved to be an impossible situation. Two attempts were made, one with calcium hydroxide as the SO2 removal sorbent and one with calcium oxide as the SO2 removal sorbent.

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Figure 5.13 Integration Option 3-Indirect fired calciner with upstream SO2 removal

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5.6.4 CCR Process Integration with a Direct-fired Calciner

The three integration options in this section are directly analogous to the previous three integration options. The difference being the Calciner is direct-fired using natural gas and oxygen. For the simulation, additional assumptions are required about the performance of the Calciner and required equipment.

The assumptions and operating conditions for each unit operation are listed below:

 Composition of Natural Gas as specified in Table 5.6

 5% excess oxygen sufficient for combustion

 Temperature of inlet Natural Gas and Oxygen high enough for combustion

 Air consists of 21% volume O2 and 79% volume N2

 ASU products simplified to oxygen and nitrogen

 ASU product distribution based on cryogenic distillation

 CO2 purity sufficient for sequestration

 CaO sorbent produced from calcination reactive towards hydration

 Calciner operates at 915 °C and can be stable

 Energy for ASU is 160 kWh/tonne O2 separated (Darde, 2009)

 Energy for CO2 stream compression to 14 MPa is 119 kWh/tonne (Wong, 2005)

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Integration Option 4, shown in Figure 5.14, uses a direct-fired natural gas fueled oxyfired

Calciner. The Calciner operates at 915 °C with the heat of combustion providing the exact required energy for calcination and solids heating and fuel heating. 1056 tph air (C-

AIR) is cryogenically distilled to produce a 95% pure oxygen stream. The oxygen stream

(O2), which flows at 244 tph, is preheated along with 59 tph of natural gas (NG-IN) using heat from the CO2 stream (CO2-OUT) produced from the Calciner. Both the NG-IN and

O2 enter the Calciner at a temperature of 335 °C. The composition of both solids and flue gas are identical to those presented in Option 1. 62 tph of CaCO3 (C-FRESH) at ambient temperature is preheated to a temperature of 558 °C using air preheated from the cooling of the purge stream. The C-FRESH is mixed with solids exiting the Purge splitter

(PURGE), which consists of 998 tph CaCO3, 303 tph CaO, and 613 tph CaSO4

(RECYCLE). The solids exiting the Calciner consists of 897 tph CaO and 613 tph CaSO4

(CAL-OUT). Since the Calciner is direct-fired, the gas stream exiting the Calciner is no longer a pure, dry stream of CO2. By volume, it consists of 65% CO2, 31% H2O, 2% N2,

1.5% O2, and the remaining 0.5% consists of products produced through incomplete combustion. On a dry basis, the CO2 stream is 95% pure requiring compression of 623 tph CO2. The steam for the Hydrator is preheated using the heat remaining from cooling the CO2 stream after preheating O2 and NG-IN. The steam is heated from 38 °C to 200

°C while the CO2 stream is further cooled from 219 °C to 117 °C. The CO2 stream is then further cooled to 35 °C to condense out the water and ready for compression. The equipment for moisture removal has been modeled in the general sense and no specific piece of equipment has been identified. The steam enters the Hydrator at 200 °C and produces 1185 tph Ca(OH)2. The flue gas stream generated from coal combustion has not

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been altered from previous options. The flue gas is cooled from 1988 °C to 629 °C, where it enters the flyash PCD. The flue gas then enters the Carbonator/Sulfator and reacts with the Ca(OH)2 at 625 °C and no heat generation. The products then enter the

PCD with the solids entering the PURGE splitter. Overall, the thermal energy required for the Calciner is approximately half the thermal input of the coal-fired power plant.

While the coal thermal energy input is 1506.79 MWth, the natural gas thermal energy input is 809.53 MWth. With the exact same assumptions, 633.7 MWe is generated after accounting for CO2 compression and ASU energy requirements. Theoretically, 779 MWe can be produced from the fossil fuel inputs. Overall, the Energy Penalty is 18.7%.

Compared to Option 1, nearly a 4% increase in Energy Penalty occurs by installing a natural gas fired, oxyfuel calciner.

Slight alterations were made to Option 4 when compared to Option 1. First, the Calciner reactor model was changed from Rstoic to RGibbs to better predict the products of both combustion and calcination. Second, due to fuel source and input temperatures, oxygen was delivered as one stream instead of creating a Primary and Secondary oxygen delivery system. Third, since the CO2 stream was not 100% pure CO2, compression energy was divided into two components, CO2 and others, where others had a molecular weight of

30. The compression energy was maintained at 119 kWh/tonne. In terms of overall utilization, energy produced from the condensation of water from the CO2 stream was not utilized. The water stream that is produced is also quite significant, totaling 123 tph. If possible, both should be utilized, as they are both useful sources of heat and raw material.

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Figure 5.14 Integration Option 4-One reaction vessel using Direct fired Calciner

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Integration Option 5, shown in Figure 5.15, is analogous to Integration Option 2. The solids flow rate are identical to those of Option 4. The gas flow rate and temperatures differ from Option 4 and Option 2. 1098 tph of Air (C-AIR) is separated in an ASU to produce a 95% volume pure stream of oxygen (O2). With the Dehydrator, a continuous stream of low pressure steam is not required, which allows for the CO2 stream to be cooled to 105 °C with the heat used to generate steam for the steam turbine. In doing so, the available heat for heat exchange from the CO2 stream decreases from 350 °C to 105

°C. The CO2 stream is heat exchanged with the oxygen stream to produce a heated oxygen stream at 95 °C and a cooled CO2 stream at 85 °C. Unlike Option 4, the Natural

Gas stream cannot be preheated using the same CO2 stream, so the hot ash stream is heat exchanged with 62 tph of natural gas at 25 °C. Ash at 625 °C is cooled to 50 °C while the natural gas is heated from 25 °C to 96 °C. Higher natural gas and air flows are required since the fuel gases enter the Calciner at a lower temperature when compared to Option 4.

Prior options have never utilized the heat from the ash. The Hydrator/Dehydrator loop cycles the 288 tph of steam required for hydration (STEAM-1/CAOH2) between 530 °C and 500 °C. Similar CO2 stream compositions to Option 4 are generated with 64% CO2,

32% H2O, 2% N2, 1.6% O2. and the balance consisting of products due to incomplete combustion. On a dry basis, this is a 95% pure CO2 stream. Overall, the Energy Penalty is

18.8%, which is nearly identical to Option 4. The reason for the near equivalency with respect to Energy Penalty arises from greater heat utilization and integration, where all available heat has been utilized, which was not the case in prior situations. The heat generated upon cooling flyash and CO2 to ambient temperatures were previously largely ignored.

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Figure 5.15 Option 5-Two reaction vessel with Direct-fired Calciner and Dehydrator

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Integration Option 6, shown in Figure 5.16, is analogous to Option 3. 70 tph of Fresh

Feed (C-FRESH) and 300 tph of steam (LPSTEAM) are required, which is greater than

Options 4 and 5, since SO2 removal prior to CO2 removal is an inefficient use of sorbent.

However, less energy is required to operate the Calciner since less solids are circulating, even though a greater amount of CaCO3 is being calcined. 1099 tph of CaCO3 and 319 tph of CaO enter the calcienr (SORBMIX). 56 tph of natural gas (NG-IN) and 996 tph air

(C-AIR) are the inlets for the fuel and oxygen. The 996 tph air translates into 231 tph oxygen stream with 95% purity (O2). Both the O2 and NG-IN streams are preheated to

335 °C while cooling down the CO2 stream from 350 °C to 226 °C. The CO2 stream is further cooled to 119 °C by heating up the 300 tph saturated steam exiting the low pressure turbine (LPSTEAM). The steam is heated to a final temperature of 200 °C before entering the Hydrator. Exiting the Hydrator is 1235 tph Ca(OH)2, which is split such that a 5:1 Ca:S mol ratio is delivered to the Sulfator while a 1.33:1 Ca:C mol ratio is delivered to the Carbonator. Compared to simultaneous CO2/SO2 removal, the solid sorbent circulation is actually lower when performing SO2 removal prior to CO2 removal.

Already stated is the lower energy required for the calcination. The overarching setback with independent SO2 removal is the solids heating and high temperature required for

FSI. With lower fuel requirements, the CO2 stream purity increases slightly. The CO2 mixture consists of 66.5% CO2, 28.6% H2O, 2% N2, 1.5% O2, and the remainder being products of incomplete combustion. On a dry basis, the CO2 purity translates into 95% purity. Overall, the Energy Penalty is 20.5%.

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Figure 5.16 Integration Option 6-Direct fired Calciner with upstream SO2 removal

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5.7 Conclusions

The difference between a direct-fired Calciner and Indirect-fired Calciner do not differ greatly. However, the Indirect-fired Calciner is always lower in Energy Penalty. Since the

CO2 generated is always pure, there is no need for an ASU, and since the coal already generates the heat for calcination, no additional energy input is necessary for an indirect- fired calciner. Technological innovation and the economics for the Capital and O&M costs will be the drivers for the Calciner design. The heat integration is essential to minimizing the overall economics given the quantities of high-quality heat available through the CCR Process. Table 5.9 summarizes the Aspen results.

Integration CO /SO Energy CO2 Net Power Calciner 2 2 Dehydrator Option Removal Penalty Purity (MWe) 1 Indirect Simultaneous No 15.1% 100% 431.0 2 Indirect Simultaneous Yes 16.0% 100% 426.4 3 Indirect Independent No 17.9% 100% 416.8 4 Direct Simultaneous No 18.7% ~95% 633.7 5 Direct Simultaneous Yes 18.8% ~95% 641.4 6 Direct Independent No 20.5% ~95% 607.6

Table 5.9 Summary of Aspen simulation results

5.8 Recommendations

The basic framework and outline of several CCR integration options have been explored.

Experiments must be performed and the results applied to the simulations to obtain a more accurate depiction of the integration. In addition, assumptions must be updated to reflect the results. As important as Energy Penalty is to the integration process, more important is the incremental cost to obtain the lower Energy Penalty.

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The Aspen simulations have been completely idealized, with 100% separation efficiencies, the assumption that any heat input into the Steam Turbine Cycle can be fully utilized, and a general 10% electricity consumption for all power plant equipment. These assumptions, although providing a basis, may fall well short of capturing the reality of power plant integration. Finally, several additional integration options exist that have not been further explored. Of importance is the investigation of a coal-fired oxyfueled calciner. Much theoretical attention has been given to the operation of a coal-fired oxyfuel calciner, but no experiments have been conducted to determine its feasibility

(Rodriguez et al., 2008; Romeo et al., 2008; Zhen-shan et al., 2008; Romano, 2009;

Ströhle et al., 2009). The effect of flyash, which will be introduced via the coal, minimum

CO2 purity, excess oxygen requirements, product calcium oxide reactivity, calciner atmosphere, and sulfur management are currently unknown parameters. Without a general idea of its feasibility, simulations will not correlate with reality.

Realistic considerations as well as economic impact have been considered but never quantified. Furthermore, capital costs have never been considered. Due to the non- existence of a Hydrator operating at 500 °C, the theoretical aspect of an oxygen-fueled

Calciner, the unknown PCDs being used, and the specific type of Carbonator/Sulfator reactor with heat recuperation, attempting to determine capital costs will at best provide a rough estimate that is within an order of magnitude of the true price. Prior to estimating capital costs, prototype equipment and discussions with industry and manufacturers must be executed.

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CHAPTER 6

CONCLUSIONS AND RECOMMENDATIONS

The Carbonation-Calcination Reaction (CCR) Process provides several advantages over competing post-combustion carbon capture technologies. The high-temperature process, with numerous exothermic reactions, and high-temperature streams create opportunities for generating high-quality steam that can be supplied to a coal-fired power plant’s existing steam turbine. By doing so, the Energy Penalty associated with the CCR Process will be minimal. Even more encouraging are three factors: 1.) No steam that would otherwise be used to generate electricity from the steam turbine is removed. 2.) The CO2 produced from the Calciner is relatively pure. 3.) The starting material can be obtained for a relatively cheap price. In amine scrubbing, the amine solvent is regenerated through the use of low-pressure steam (Klemeš et al., 2006). In the CCR Process, steam is removed at the exit of the low-pressure steam turbine. The only cost associated with the steam is the replacement feedwater required. Depending on calciner configuration, the

CO2 can be immediately sequestered without additional polishing or a condenser unit must be supplied to remove moisture. In either situation, the CO2 will be high in purity, which reduces additional processing prior to sequestration. Finally, since limestone is the initial feed, the majority of the cost is associated with transportation. A coal-fired power plant near a limestone mine can obtain limestone for as low as $7/ton (Wolfe, 2006).

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Results from the 20 pound per hour coal-fired stoker facility indicate a high level of reactivity is maintained by calcium hydroxide over a number of cycles. Furthermore, decay over multiple cycles did not appear over 5 recycles. Based on Aspen simulations, the possibility of very low Energy Penalty’s is achievable. Ranging from a total of 15% to 21% Energy Penalty, even the maximum is significantly lower than current technologies. In addition, common to all post-combustion capture technologies is the requirement of CO2 compression, which is approximately 10% of the total Energy

Penalty. The true Energy Penalty, which omits compression, associated with the CCR

Process is in the range of 5% to 10%.

In order to realize the potential of the CCR Process, additional work is required in several areas. First, reaction vessels capable of extracting heat on a continual basis with high solids loading must be designed. Pipe fouling can lead to a rapid decrease in heat transfer rate. Second, appropriate Particulate Capture Devices must be identified. The high- temperature operation and the necessity to separate a high solids loading containing a micron-sized Particle Size Distribution poses multiple challenges in terms of cost and design. Third, the Calciner must be designed to operate in a manner that produces highly reactive calcium oxide and a high-purity CO2 stream. Finally, to incorporate a complex process such as the CCR Process into a power plant, a detailed economic analysis must be performed. Preliminary results from Aspen simulations and a 120 kWth fossil-fuel fired combustor shows promising results for integrating the CCR Process into a power plant.

As a whole, a continuous, integrated unit with a Hydrator and appropriate PCDs for sorbent circulation should be installed into the 20 pph unit. Allowing for a continuous

112

hydration cycle will increase the number of cycles performed in a short amount of time, as one cycle is approximately 45 to 50 minutes. The long-term cyclability and effect of sorbent decay and sulfur build-up must be examined.

The research that has been performed at The Ohio State University, along with supporting research worldwide, indicates the CCR Process has significant potential.

Additional considerations that should not be overlooked include end uses for both the

CO2 and solid streams. Research in carbon sequestration is tangential, but crucial, to the success of the CCR Process since the captured CO2 cannot be released into the atmosphere. End uses that will aid in the reduction of the CCR Process economics include Enhanced Oil Recovery and Enhanced Coalbed Methane. Finally, the solids may have many beneficial end uses that will also lower the CCR Process economics.

Environmental areas, such as soil remediation and mine reclamation, along with commercial areas, such as the construction industry, may find value in the waste solids.

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