Chemical Looping Partial Oxidation Process for Syngas Production

DISSERTATION

Presented in Partial Fulfillment of the Requirements for the Degree Doctor of Philosophy in the Graduate School of The Ohio State University

Dikai Xu

Graduate Program in Chemical Engineering

The Ohio State University

2017

Dissertation Committee:

Liang-Shih Fan, Advisor, David Wood, Barbara Wyslouzil, and Jacques Zakin

Graduate Faculty Representative:

Esperanza Carcache de Blanco

Copyrighted by

Dikai Xu

2017

Abstract

The chemical looping partial oxidation process is developed for the efficient conversion of gaseous and solid fuels into syngas via partial oxidation. The chemical looping partial oxidation process converts the fuels into high purity syngas with flexible

H2:CO ratio that is suitable for downstream fuel or chemical synthesis. In the chemical looping partial oxidation process, the fuels are partially oxidized in the reducer reactor by the oxygen carrier to generate high purity syngas. The reduced oxygen carrier is regenerated in a fluidized bed combustor via the oxidation reaction with air. Compared to the conventional syngas generation processes, the chemical looping partial oxidation process eliminates the need for additional steam or molecular oxygen from an air separation unit (ASU), resulting in an increased cold gas efficiency and decreased fuel consumption. The chemical looping partial oxidation process features the combination of an iron-titanium composite metal oxide (ITCMO) oxygen carrier and a co-current gas- solid moving bed reducer reactor. The ITCMO oxygen carrier is selected for the chemical looping partial oxidation process due to its desired thermodynamic and kinetic properties.

Theoretical analysis aided by a modified Ellingham Diagram illustrates that syngas production is thermodynamically favored in the presence of ITCMO oxygen carrier. The co-current moving bed reducer design provides a desirable gas-solid contacting pattern that minimizes carbon deposition while maximizing the syngas yield. Experimental studies in a fixed bed reactor and a bench scale reactor successfully demonstrate the ii

production of high purity syngas from methane and biomass with the combination of moving bed reducer and ITCMO oxygen carrier. Further scale-up of the chemical looping partial oxidation process is demonstrated in an integrated sub-pilot scale reactor system using non-mechanical gas sealing and solid circulation devices. A dynamic modeling scheme is developed for studying the transient behavior and the control of the chemical looping system. A hierarchical control system based on sliding mode control concept is developed for the chemical looping technologies to simplify process operation.

iii

Acknowledgments

I would like to express my gratitude to all those who gave me the possibility to complete this thesis. I want to first thank my advisor Dr. Liang-Shih Fan for offering me the opportunity to pursue my degree at OSU. I was impressed and motivated by Dr. Fan’s enthusiasm for the research in our group. The unique research experience along with Dr.

Fan’s tuition I received over the past five years will greatly benefit my future.

The research group not only exposed me to the invaluable research experience, but also introduced mentors, collaborators, and friends to me. In particular, I want to thank

Liang Zeng, Andrew Tong, and Siwei Luo for offering me the guidance and help on research. The work presented in this dissertation is, in part, inspired by the discussion with them. I greatly benefited from the discussions with my knowledgeable collaborators on various topics. For that I want to thank Lang Qin, Dawei Wang, Zhuo Cheng, Pengfei

He, and Qiang Zhou from our research group, Umit Ozguner and Arda Kurt from OSU

ECE Department, and Tom Flynn, Tim Fuller, Chris Poling, Luis Velazquez-Vargas, and

Bill Arnold from Babcock & Wilcox. I am grateful to my all lab mates and friends, Tien-

Lin Hsieh, Cheng Chung, Mandar Kathe, Ankita Majumdar, Sam Bayham, Aining Wang,

Yitao Zhang, Mengqing Guo, Sourabh Nadgouda, Mingyuan Xu, Yaswanth Pottimurthy,

Cody Park, Fanhe Kong, and Yu-Yen Chen, for their support in my research as well as my daily life.

Vita

July 2012 ...... B.S. Chemical Engineering, Tsinghua University, Beijing, China

Publications

Xu, D., Zhang, Y., Hsieh, T.-L., Guo, M., Qin, L., Chung, C., Fan, L.-S., & Tong, A.

(2017). A Novel Chemical Looping Partial Oxidation Process for Thermochemical

Conversion of Biomass to Syngas. In second review at Energy & Environmental

Science.

Luo, S., Zeng, L., Xu, D., Kathe, M., Chung, E., Deshpande, N., Qin, L., Majumder, A.,

Hsieh, T.-L., Tong, A., Sun, Z., & Fan, L.-S. (2014). Shale gas-to-syngas chemical

looping process for stable shale gas conversion to high purity syngas with a H 2:

CO ratio of 2: 1. Energy & Environmental Science, 7(12), 4104-4117.

Qin, L., Cheng, Z., Fan, J. A., Kopechek, D., Xu, D., Deshpande, N., & Fan, L.-S.

(2015). Nanostructure formation mechanism and ion diffusion in iron–titanium

composite materials with chemical looping redox reactions. Journal of Materials

Chemistry A, 3(21), 11302-11312.

Luo, S., Majumder, A., Chung, E., Xu, D., Bayham, S., Sun, Z., Zeng, L., & Fan, L.-S.

(2013). Conversion of Woody Biomass Materials by Chemical Looping Process

Kinetics, Light Tar Cracking, and Moving Bed Reactor Behavior. Industrial &

Engineering Chemistry Research, 52(39), 14116-14124.

Cheng, Z., Qin, L., Guo, M., Fan, J. A., Xu, D., & Fan, L.-S. (2016). Methane adsorption

and dissociation on iron oxide oxygen carriers: the role of oxygen vacancies.

Physical Chemistry Chemical Physics, 18(24), 16423-16435.

Qin, L., Guo, M., Cheng, Z., Xu, M., Liu, Y., Xu, D., Fan, J. A., & Fan, L.-S. (2017).

Improved cyclic redox reactivity of lanthanum modified iron-based oxygen carriers

in carbon monoxide chemical looping combustion. Journal of Materials Chemistry

A. DOI: 10.1039/C7TA04228K.

Fields of Study

Major Field: Chemical Engineering

Table of Contents

Abstract ...... ii

Acknowledgments...... iv

Vita ...... v

Table of Contents ...... vii

List of Tables ...... x

List of Figures ...... xi

Chapter 1: Introduction ...... 1

The Production of Syngas ...... 1

Chemical Looping Technologies ...... 6

Chemical Looping Combustion ...... 7

Chemical Looping Partial Oxidation ...... 11

Chemical Looping Systems using Moving Bed Reactors ...... 13

Chapter 2: Thermodynamics of Chemical Looping Partial Oxidation ...... 17

Thermodynamics of Oxygen Carriers ...... 17

Thermodynamics of the Iron-Titanium Composite Metal Oxide (ITCMO) ...... 27

vii

Reactor Design Considerations ...... 31

Mass Balance and Thermodynamics in Moving Bed Reactors ...... 38

Conclusion ...... 49

Chapter 3: Experimental Studies of Chemical Looping Partial Oxidation ...... 52

Reactivity of the ITCMO ...... 52

Syngas Generation in a Fixed Bed Reactor ...... 56

CH4 Partial Oxidation in a Moving Bed Reducer ...... 61

Solid Fuel Gasification in a Moving Bed Reducer ...... 70

Scale-up of Chemical Looping Partial Oxidation ...... 88

Sizing of Reactors ...... 90

Design of the Sub-pilot Reactor System ...... 92

Operation of the Sub-pilot Reactor System ...... 99

Conclusion ...... 101

Chapter 4: Dynamic Simulation of Chemical Looping Systems ...... 102

Physical Model ...... 103

Components of the System ...... 103

Mass Balance ...... 104

Pressure Balance and Hydrodynamics ...... 105

Assumptions for Simulation ...... 107

viii

Mathematical Description ...... 108

Mass Balance in the Reactors ...... 110

Pressure Drop and Pressure Distribution along a Moving/Packed Bed ...... 111

Gas Flow Rates through Valves ...... 114

Pressure Drop in the Fluidized Bed Reactor and The Riser ...... 117

Pressure Drop And Gas Flow in Zone Seals ...... 122

Solving for the System Dynamics ...... 123

Conclusion ...... 125

Chapter 5: Hierarchical Control of Chemical Looping Systems ...... 126

Reactor Pressure Control using HLC-SMCs ...... 127

Automatic Operation of Chemical Looping Systems ...... 133

Conclusion ...... 142

Chapter 6: Conclusions and Future Researches ...... 143

Conclusions ...... 143

Future Researches ...... 147

References ...... 150

List of Tables

Table 1. Solid fuel gasification technologies ...... 4

Table 2. Early development and testing of chemical looping processes ...... 7

Table 3. Iron phases in each zone of Figure 6 ...... 26

Table 4. Proximate and ultimate analysis results for the wood pellets ...... 72

Table 5. Summary of experiment conditions for biomass gasification reducer tests ...... 76

Table 6. Comparison between bench scale reducer experiment results and ASPEN

RGibbs reducer model results for biomass gasification ...... 82

Table 7. Process inlets and outlets on the sub-pilot reactor ...... 95

Table 8. Major instrumentation on the sub-pilot scale unit ...... 97

Table 9. Pressure control valves in the chemical looping system ...... 116

List of Figures

Figure 1. Concept of chemical looping ...... 6

Figure 2. Conceptual design of chemical looping processes with a counter-current (a) and

co-current (b) moving bed reducer reactor ...... 15

Figure 3. The modified Ellingham Diagram ...... 20

Figure 4. Carbon distribution in the equilibrium product of the CeO2-CH4 system at

900°C. [O]/CH4 represents the molar ratio between usable oxygen in CeO2 and

CH4 feedstock...... 23

Figure 5. Carbon distribution in the equilibrium product of the NiO-CH4 system at

900°C. [O]/CH4 represents the molar ratio between usable oxygen in NiO and

CH4 feedstock...... 24

Figure 6. Equilibrium product distribution of the Fe2O3-CH4 system at 900°C with a CH4

flow rate of 1 mol/s...... 26

Figure 7. Phase diagram of the Fe2O3-CO system ...... 27

Figure 8. Phase diagram of the ITCMO-CO system...... 28

Figure 9. Oxidation reaction of Fe with and without TiO2 ...... 30

Figure 10. Weight change of a Fe2O3-based oxygen carrier pellet in a redox cycle ...... 32

Figure 11. Internal age distribution of particles in a perfectly mixed reducer ...... 34

Figure 12. Thermodynamic equilibrium lines and operation lines for moving bed reducer

and oxidizer using Fe2O3 oxygen carrier at 900°C ...... 39

Figure 13. Thermodynamic equilibrium lines and operation lines for moving bed reducer

and oxidizer using the ITCMO oxygen carrier at 900°C ...... 42

Figure 14. Operation region of the co-current moving bed reducer in chemical looping

partial oxidation ...... 45

Figure 15. Operation region of the co-current moving bed reducer at complete CH4

conversion ...... 47

Figure 16. Recyclability test on the ITCMO oxygen carrier ...... 53

Figure 17. Fixed bed experiment setup for biomass tar cracking ...... 55

Figure 18. MS data from fixed bed biomass pyrolysis with and without oxygen carriers 56

Figure 19. Fixed bed experiment setup for CH4 conversion ...... 57

Figure 20. Product gas composition of fixed bed reactor with CH4 injection. (1) lowered

gas flow rate; (2)increased temperature...... 59

Figure 21. Product gas composition of fixed bed reactor with CO2 injection. (1) increased

CO2 concentration ...... 60

Figure 22. Bench scale co-current moving bed reducer ...... 62

Figure 23. Schematic electrical connection diagram of the bench scale reactor system .. 64

Figure 24. Exemplary temperature distribution in the bench scale reactor ...... 65

Figure 25. Syngas composition from the bench scale reducer at 1040°C using methane as

feedstock ...... 67

xii

Figure 26. Quality of syngas from the bench scale reducer at 1040°C using methane as

feedstock ...... 68

Figure 27. Gas composition at different locations of the bench scale reducer at 1020°C 69

Figure 28. XRD pattern of the reduced ITCMO oxygen carrier ...... 70

Figure 29. Gaseous product from BTS reducer at equilibrium ...... 73

Figure 30. Effect of H2O injection to the reducer on H2:CO ratio at 1000°C ...... 74

Figure 31. Bench scale co-current moving bed reducer for solid fuel conversion ...... 75

Figure 32. (a) Molar ratio between steam and carbon in wood pellet and (b) N2 free

syngas composition and dry syngas purity under Test Condition 1 ...... 77

Figure 33. (a) Molar ratio between steam and carbon in wood pellet and (b) N2 free

syngas composition and dry syngas purity under Test Condition 2 ...... 81

Figure 34. (a) Oxygen carrier conversions and (b) XRD spectrums of oxygen carrier

samples from different locations ...... 85

Figure 35. Reactions occurring in different locations of the reducer for biomass

gasification ...... 87

Figure 36. 3D drawing (left) and picture (right) of the sub-pilot reactor on structure ..... 94

Figure 37. Supporting seat for the reducer separator ...... 94

Figure 38 Process inlets and outlets on the sub-pilot reactor ...... 96

Figure 39. Control system of the sub-pilot scale unit ...... 98

Figure 40. N2-free product gas composition at the reducer outlet of the sub-pilot unit ... 99

Figure 41. Syngas generation performance of the sub-pilot unit ...... 100

xiii

Figure 42. Components and Physical Quantities in the Model of Chemical Looping

System ...... 109

Figure 43. Flow characteristics of an actual equal percentage globe valve ...... 117

Figure 44. Relation between and ...... 121

Figure 45. One-tank reactor model ...... 127

Figure 46. Steps in the pressurization of the one-tank reactor ...... 128

Figure 47. Controller for pressurization of the one-tank reactor ...... 129

Figure 48. Simulation result for pressure regulation using HLC-SMCs ...... 132

Figure 49. System status trajectory in the phase plane during simulation ...... 133

Figure 50. Performance of the HLC-SMC hybrid controller on the pilot plant for

pressurizing and depressurizing the reactor ...... 134

Figure 51. (a) Schematic of the chemical looping sub-pilot unit; (b) Sensor inputs and

manipulated variables for process control...... 136

Figure 52. Performance of the HLC-SMCs hybrid controller during heating ...... 139

Figure 53. Performance of the HLC-SMCs hybrid controller during fuel injection ...... 141

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Chapter 1: Introduction

The Production of Syngas

Syngas refers to a mixture of H2 and CO. It is widely used as a chemical intermediate for the production of valuable commodity chemicals from a variety of carbonaceous feedstock, such as natural gas, coal, and biomass. The products that can be produced from syngas include synthetic liquid fuels, methanol, dimethyl ether, acetic acid, hydrogen, and ammonia. The downstream process may require a syngas composition with a specific H2:CO molar ratio and certain upper limits on the impurities, such as CO2, hydrocarbons, and H2S. For example, the methanol synthesis process and the cobalt-based Fischer-Tropsch process for liquid fuel production requires a H2:CO ratio of 2:1. However, the composition of syngas varies widely depending on the nature of the feedstock and the process that is used. Syngas produced from solid fuels, such as coal and biomass, typically contains more CO than H2 when air or O2 is used as an oxidant in the process, while that produced from natural gas contains a higher H2 content.

Syngas conditioning processes, such as steam reforming, water gas shift (WGS), sulfur removal, and CO2 removal, are usually required upstream the chemical synthesis step.

The state-of-the-art processes for the production of syngas can be classified into two categories, i.e. reforming processes for gaseous fuel conversion, and gasification processes for solid fuel conversion. All these processes involve the partial oxidation of a carbonaceous fuel using an oxidant, such as air, O2, H2O, and CO2.

One of the most commonly used process for the conversion of natural gas to syngas is the steam methane reforming (SMR) process.[1-9] The main reaction in SMR is:

3 Δ≅225⁄ 1‐1

To maximize the CH4 conversion in reaction (1-1), excess amount of steam is usually injected into the reactor.[10,11] Thus, the WGS occurs at the same time as a side reaction:

1‐2

Typical SMR reactors are tubular fixed bed reactors, which are filled with Ni-based catalysts and are operated at a pressure of 14 to 20 atm and a temperature of 800 to

1000°C.[12] Because the main reaction (1-1) is strongly endothermic, the tubular reactors are usually embedded in a furnace fueled by natural gas and/or combustible tail gases from downstream processes, which results in significant amount of CO2 emission from the process. The syngas produced from a SMR reactor typically consists of 66% H2, 12%

CO, 12% CH4, and 10% CO2. The high H2 content in the product syngas makes SMR a widely employed process for large scale H2 production. The product syngas from the

SMR reactors can be further converted to H2 in a WGS reactor, followed by a pressure swing adsorption for H2 purification.

Another commonly used natural gas reforming process is the autothermal reforming

(ATR) process, in which O2, obtained from an air separation unit (ASU), is used as the oxidant in the CH4 reforming reactors along with H2O.[13-15] The ATR process uses the thermal energy from the exothermic reactions between CH4 and O2 to compensate the energy requirement for reaction (1-1). Partial and full oxidation reactions of CH4 may occur in the reactor:

2 24 1‐3

2 2 1‐4

Large scale ATR processes are designed with a two-step configuration.[13] The natural gas is partially oxidized by H2O in a catalytic “pre-reformer” before being mixed with O2 in the following catalytic “autothermal reformer”. The ATR process can directly produce a syngas with a H2:CO of 2, which is suitable for methanol and liquid fuel synthesis. Commercial scale natural gas to liquid fuel (gas-to-liquid, or GTL) plants using the ATR process, including the Oryx GTL plant, have been built and operated.[16]

In addition to the SMR and ATR process, the partial oxidation (POX) process for natural gas conversion is also commercialized.[17-19] The POX process uses O2 from an

ASU as the oxidant to convert CH4 via reaction (1-3) without the use of catalysts. As determined by the stoichiometric ratio of reaction (1-3), the syngas produced from the

POX process has a H2:CO ratio close to 2, desirable for liquid fuel and methanol synthesis. The POX process is used for commercial scale production of syngas in the world’s largest GTL plant, the Pearl plant at Qatar.

The conversion of solid fuels, such as coal and biomass, has been of great interest in areas with limited natural gas resources and abundant solid fuel reserve. A variety of solid fuel gasification technologies have been developed to convert coal and/or biomass into syngas via partial oxidation reactions. Table 1 summarizes several solid fuel gasification technologies that are developed.[20-33]

Gasifier Name Oxidant Feedstock Reactor Type Lurgi Dry Ash Steam/O2 Dry coal Moving bed Slagging BGL Steam/O2 Dry coal Moving bed High Temperature Winkler Steam/O2 Dry coal Fluidized bed KBR TRIG Steam/O2 Dry coal Fluidized bed Koppers-Totzek O2 Dry coal Upward entrained bed E-Gas O2 Coal slurry Upward entrained bed Texaco O2 Coal slurry Downward entrained bed Shell SCGP O2 Dry coal Upward entrained bed Foster Wheeler Air Biomass Fluidized bed Battelle Steam Biomass Fluidized bed Table 1. Solid fuel gasification technologies

The type of gasifier and oxidant for the solid fuel gasification process varies depending on the downstream application. For example, in the integrated gasification combined cycle (IGCC) system, the syngas produced from the gasification process will be burned in the turbine for power generation. Because high H2 content and low N2 dilution in syngas is not required for gas combustion turbines, air or O2 blown gasifiers can be used. Steam and/or O2 are used as the oxidant for high purity syngas generation with low N2 content when the product syngas will be used in chemical synthesis.

Similar to the SMR reaction (1-1), the gasification reactions of coal and biomass using steam are endothermic. In order to compensate the heat requirement, air or O2 are injected into the gasifier to partially oxidize the fuels in most gasifiers. Alternatively, the

Battelle indirectly heated biomass gasifier supplies the required thermal energy by circulating sand as a heat carrier between two reactors.[32] In the gasification reactor, the biomass feedstock is heated by the high temperature sand and reacts with H2O to generate syngas. The unconverted solid char is circulated with the sand to the combustion reactor, where the char is combusted in air to re-heat the sand. By circulating the sand between the reactors, thermal energy released in char combustion is carried over to the gasification reactor for biomass conversion.

In most of the state-of-the-art reforming and gasification processes, ASUs are required to supply the molecular O2 needed for the generation of high purity syngas with low N2 dilution. However, the operation of ASUs is energy intensive, which significantly lowers the energy efficiency of the gasification plant and increases the cost. The processes that use H2O as the only oxidant, such as the SMR process and the Battelle indirectly heated gasification process, are unable to fully convert the hydrocarbons in the fuels into syngas. Thus, additional hydrocarbon conversion units may be required to increase the process efficiency.

Chemical Looping Technologies

In recent efforts, the chemical looping technologies are used for the conversion of carbonaceous fuels into syngas. As shown in Figure 1, the chemical looping processes decouple the oxidation reaction of carbonaceous fuels into two steps, which typically occur in two separate reactors. In the reducer, the fuel is oxidized by a solid oxygen carrier MeOx. The reduced oxygen carrier MeOy is separated from the gaseous product and circulated to the combustor, where air is used to regenerate the oxygen carrier via reaction (1-5):

→ 1‐5

Figure 1. Concept of chemical looping

Although the term of “chemical looping” was not used until the late 1980s,[34] chemical processes based on the principle of chemical looping have been under development for over a century. Early chemical looping processes were developed for the purpose of chemical production, including H2, CO2, syngas, and other chemicals, from carbonaceous fuels as opposed to today’s focus on carbon emissions control. Table 2 summarizes several representative chemical looping processes.[35] However, all of the earlier processes are no longer in practice due to limitations in the oxygen carrier and reactor performance making them unable to compete with rival technologies at the time.

Process/ Lewis and Lane HYGAS ARCO GTG Developer Gilliland Year 1910s 1950s 1970s 1980s

Feedstock Syngas Solid Fuel Syngas CH4

Products H2 CO2 H2 C2H4 Redox pair / CuO-Cu2O or Fe3O4-Fe Fe3O4-Fe supported Mn looping materials Fe2O3-Fe3O4

Process/ Solar water DuPont Otsuka Steinfeld Developer splitting Year 1990s 1990s 1980s 1990s

Feedstock C4H10 CH4 H2O CH4, iron ore

Products C4H2O3 Syngas H2, O2 syngas, iron vanadium Redox pair / ZnO-Zn or phosphorous supported CeO Fe O -Fe looping materials 2 Fe O -FeO/Fe 3 4 oxide 3 4 Table 2. Early development and testing of chemical looping processes

Chemical Looping Combustion

Since 1990s, research efforts in chemical looping technologies have been motivated by the demand for an efficient and cost effective CO2 capture technology in the power 7

generation industry. Thus, the most widely studied application of the chemical looping technology was the chemical looping combustion for thermal energy generation with in- situ CO2 capture.[36] In chemical looping combustion systems, the full oxidation of the carbonaceous fuels is desired. The main reaction in the reducer is:

→ 1‐6

After condensing the H2O, a stream of high purity CO2 is obtained, which can be compressed for sequestration or other applications. The overall reaction in the system is identical to the normal combustion processes. Thermal energy is extracted from the reducer and combustor for steam/power generation. Compared to other carbon capture technologies, such as the oxy-combustion and the post combustion capture, the chemical looping combustion technology eliminates the need for energy intensive ASUs and CO2 separation units. Thus, a higher energy conversion efficiency can be achieved with CO2 captured. To date, several chemical looping combustion units have been constructed and demonstrated worldwide at scales up to 3MWth.

The oxygen carriers used in the chemical looping combustion processes typically consist of synthesized composite or naturally occurring metal oxide material. A large number of oxygen carrier materials using the Fe2O3, NiO, Mn2O3, CuO, CaSO4, and more complex perovskite materials as the active ingredients have been synthesized and studied.

When Fe-based or Ni-based oxygen carriers are used in the chemical looping combustion system, the conversion of the oxygen carriers in the reducer relies on the gas- solid reaction between the reducing gases and the solid oxygen carriers. As a result, the conversion of solid fuels has to occur in two steps:

The solid fuel is gasified via pyrolysis and its reactions with H2O and CO2

to produce combustible gases;

The combustible gases are oxidized by the oxygen carriers.

Char gasification is usually the rate limiting step due to its slow reaction kinetics compared to other reactions. These systems are sometimes referred to as chemical looping combustion with in-situ gasification, or CLC-iG.

When Mn-based or Cu-based oxygen carriers are used, molecular O2 can be released from the oxygen carriers in the reducer. O2 can react with the gaseous and solid fuels rapidly and fully convert them into CO2 and H2O. Thus, oxygen carriers that can release

O2 are of particular interest for the chemical looping combustion of solid fuels. The systems using these oxygen carriers are sometimes referred to as chemical looping with oxygen uncoupling (CLOU) systems.

Ni-based metal oxides are among the earliest candidates for oxygen carrier materials because of their superior reactivity. Researchers from the Vienna University of

Technology have extensively tested Ni-based oxygen carriers with various supporting material compositions for gaseous fuel combustion in a 120-140kWth pilot scale dual circulating fluidized bed (DCFB) test unit.[37][38] Fuels tested include natural gas and syngas, and >95% gas conversions were achieved in multiple test campaigns. Despite the favorable kinetics and potential catalytic capability of Ni-based oxygen carrier, interests have shifted due to growing concerns over its high cost and toxicity to health and environment.

Recently, natural occurring ores have been widely reported as the oxygen carrier material for chemical looping combustion systems. Natural ores minimize the cost associated with oxygen carrier synthesis. Iron-based natural ores such as ilmenite have been of particular focus. Researchers from the Chalmers University of Technology operated a 100kWth circulating fluidized bed (CFB) unit for solid fuel conversion.

Multiple types of coals were tested using ilmenite as the oxygen carrier with >96% carbon capture efficiencies reported.[39] Mixing a manganese-based ore with ilmenite showed a significantly increase of the reducer product gas conversion from 84% under ilmenite-only condition to 91.5%.[40] Researchers from Technische Universität

Darmstadt recently demonstrated the world’s first autothermal operation of metal oxide based chemical looping combustion on a 1MWth CFB unit using ilmenite as oxygen carrier and coarse hard coal as fuel. A carbon capture efficiency of 44–52% was reported.[41]

Cu-based oxygen carriers are less commonly demonstrated at larger scales due to its low melting point temperature of metallic copper and the cost of synthesis. However, many researchers are analyzing the effect of addition of different support materials to copper on the redox reaction pathway. Researchers from University of Utah recently finished construction and commissioning of a 200kWth unit using the CLOU reaction schemes with Cu-based oxygen carrier supported by SiC.[42] The conversion of coal in this system is under investigation.

The CaS/CaSO4 redox couple is also demonstrated for chemical looping combustion application. Alstom Clean Power Plant Laboratory (currently part of GE) achieved the

world’s first autothermal operation of a CaSO4-based chemical looping combustion process on a 3MWth pilot unit.[43] Different from the other metal oxide based chemical looping combustion processes, CaSO4 oxygen carrier introduces the additional challenge of oxygen carrier recyclability as loss of sulfur occurs during the redox cycles.

Chemical Looping Partial Oxidation

Chemical looping technologies can be used for not only the full oxidation, but also the partial oxidation of carbonaceous fuels. Depending on the feedstock used in the processes, these systems are sometimes referred to as chemical looping reforming (CLR) or chemical looping gasification (CLG). The primary reaction in the reducer of chemical looping partial oxidation is:

→ 1‐7

Similar to the chemical looping combustion processes, the reduced oxygen carriers are transported to the combustor for regeneration. The overall reaction in the system is the partial oxidation of the fuel by O2 in air to generate syngas.

For carbonaceous fuels with a high heat of combustion, such as CH4, the exothermic heat released from the partial oxidation reaction can maintain the autothermal operation of the system. However, when fuels with low heating values, such as biomass, are used, a portion of the fuels need to be converted to CO2 and H2O in order to generate sufficient thermal energy. Efficient heat integration among the reactors and the inlet/outlet streams are required to maintain autothermal operation while supplying the energy for the auxiliary equipment in the process.

Compared to chemical looping combustion, research efforts on chemical looping partial oxidation processes are less reported. Researchers from the Chalmers University of Technology reported chemical looping partial oxidation of CH4 in a 300 Wth two- compartment fluidized bed reactor system using nickel-based oxygen carriers.[44-47]

Significant carbon deposition on the reduced Ni-based oxygen carrier in the reducer is observed when oxygen supply to the system is lowered to inhibit the full oxidation reactions. In order to reduce the carbon deposition, 30% of H2O or CO2 is fed into the reactor along with CH4. The system was able to produce a high purity syngas that is suitable for chemical synthesis.

Researchers from the Vienna University of Technology reported chemical looping partial oxidation of CH4 in a 140 kWth DCFB reactor using nickel-based oxygen carriers.[48,49] By reducing the air supply, the system was able to produce a syngas with

40% CO2 (dry base) without significant carbon deposition in the reducer. Further reducing the air supply may cause carbon deposition, as seen in the Chalmers University of Technology unit.

Chemical looping partial oxidation of solid fuels using Fe-, Ni-, and Cu-based oxygen carriers are also reported. Researchers from the Guangzhou Institute of Energy

Conversion, Chinese Academy of Sciences reported the conversion of biomass to syngas using iron-based oxygen carrier in a 10 kWth CFB unit.[50-52] Researchers from the

Southeast University, China reported a similar biomass conversion unit at 25 kWth using

Ni- and iron ore-based oxygen carriers.[53,54] The syngas produced from these chemical looping partial oxidation systems typically consists of 40% CO, 20% H2, 20% CO2, 10%

CH4, and 10% other gases. Thus, a steam reforming reactor is required to further convert the hydrocarbons in the product syngas. CO2 emission observed from the combustor increases with an increasing biomass feeding rate, which is attributed to the insufficient char residence time in the reducer and the slow reaction kinetics of char gasification.

Increasing the reactor temperatures promotes the full conversion of hydrocarbons and char in the reducer. However, full oxidation of the syngas to CO2 and H2O is also accelerated.

Chemical Looping Systems using Moving Bed Reactors

The majority of the chemical looping processes developed to date utilize fluidized bed designs for fuel conversion due to the advantages of rapid reaction kinetics and optimal heat transfer characteristics. Alternatively, The Ohio State University has developed a unique moving bed reducer design the conversion of carbonaceous fuels. As illustrated in Figure 2, a counter-current gas-solid contact pattern is employed for full fuel combustion to CO2/H2O and steam-iron reaction to H2, while a co-current contact pattern is used for fuel gasification/reforming to syngas.[35, 36, 55] The packed moving bed reactor design allows the oxygen carrier and reactant/product to flow along the axial direction while preventing over conversion due to gas-solid back mixing and under conversion due to either mass diffusion limitation in gas phase bubbles in the fluidized bed or short circuiting of unreacted oxygen carriers and char. With proper oxygen carrier to fuel ratio and sufficient reactant residence time, uniform oxygen carrier and fuel conversion can be achieved in a moving bed reducer.

The selection of co-current or counter-current contact pattern is corroborated with the thermodynamic equilibrium between desired product gas and oxidation state of the oxygen carriers. For fuel combustion to CO2 and H2O with an iron-based oxygen carrier, the counter-current gas-solid contact pattern ensures the product gas is in contact with

Fe2O3 prior to exit, which then allows the reducer to achieve almost complete fuel conversion to CO2/H2O while maximizing the conversion of the Fe2O3-based oxygen carrier. The reduced oxygen carrier contains Fe and/or FeO phases, which allows hydrogen production in the subsequent moving bed oxidizer. Thus, high purity H2 can be produced from a chemical looping system with a 3-reactor configuration, including a reducer, an oxidizer, and a combustor. For the partial oxidation of fuel, the co-current gas-solid contact pattern ensures the product gas is in contact with the reduced oxygen carrier in the FeTiO3 and/or Fe oxidation state prior to exit which then allows the reducer to produce >90% syngas.

(a) (b)

Figure 2. Conceptual design of chemical looping processes with a counter-current (a) and co-current (b) moving bed reducer reactor

This dissertation presents the effort to develop and scale-up the chemical looping partial oxidation process using co-current moving bed reducer and the iron-titanium composite metal oxide (ITCMO) oxygen carrier. Chapter 2 of this dissertation discusses the theoretical rationale in the design of the chemical looping partial oxidation process using co-current moving bed reducers. The thermodynamic properties of the oxygen

carriers, and their impact on the design and operation of the moving bed reactor are discussed. Chapter 3 presents the experimental studies that demonstrate the feasibility of high purity syngas generation using the ITCMO oxygen carriers and the co-current moving bed reducer. The conversion of both gaseous fuel and solid fuel are discussed.

Chapter 4 presents a dynamic simulation scheme for studying the transient hydrodynamic behavior in the integrated chemical looping reactor. Chapter 5 presents the effort to develop a hierarchical control scheme for the automation of the chemical looping process.

Chapter 2: Thermodynamics of Chemical Looping Partial Oxidation

The thermodynamic equilibria between the solid and gaseous reactants significantly affect the performance of a chemical looping system. In case of chemical looping combustion systems, the equilibrium between the oxygen carriers and the fuel or flue gas sets the upper limit for the fuel conversion in the reducer. In CLOU systems, the equilibrium between the oxygen carriers and oxygen sets the lower limit for the air flow in the combustor. This chapter examines the thermodynamics of the chemical looping partial oxidation systems. In particular, the equilibria between the oxygen carriers and the gaseous reactants in the system, and the implication of these equilibria properties on reactor design, are discussed.

Thermodynamics of Oxygen Carriers

The chemical reactions involved in the chemical looping process include the following:

Reducer:

Δ2‐1 17

Δ 2‐2

Δ 2‐3

Oxidizer:

Δ 2‐4

Combustor:

Δ 2‐5

For chemical looping combustion systems, the reactions (2-1), (2-2), and (2-3) should all be spontaneous with negative values of standard Gibbs free energy change

(Δ , which enables the complete conversion of carbonaceous fuel to CO2 and H2O.

However, for chemical looping partial oxidation systems, reactions (2-2) and (2-3) are undesirable because CO and H2 are the products of the process. Thus, an oxygen carrier

material that gives rise to positive Δ for reactions (2-2) and (2-3), i.e. Δ 0 and

Δ 0, is advantageous for the chemical looping partial oxidation systems.

The thermodynamic properties of reactions (2-2) and (2-3) can be determined by

examining the property of the oxygen carrier, specifically, the Δ of its oxidation reaction. For example, reactions (2-2) can be regarded as the “difference” between reaction (2-6) and (2-8):

2 2 Δ 2‐6

2 2 Δ 2‐7

Δ 2‐8

Thus, the Δ of reaction (2-2) is the difference between those of reaction (2-6) and (2-

8):

Δ Δ Δ 2‐9

A positive value of Δ requires Δ Δ . Similarly, a positive Δ requires

Δ Δ . Thus, the oxygen carrier material with desirable thermodynamic property

can be selected by examining the value of Δ .

A modified Ellingham Diagram, as shown in Figure 3, is developed to facilitate the

selection of oxygen carrier material. The modified Ellingham Diagram plots the Δ of the oxidation reactions of metals, metal oxides, and other substance over a range of temperature, normalized to consume 1 mole of O2 in the reaction.

The location of a line in the modified Ellingham Diagram reflects the relative tendency for a compound to be reduced. Specifically, consider the following two reactions:

Δ 2‐10

Δ 2‐11

If the line for reaction (2-10) locates above that for reaction (2-11), it can be inferred that:

compound AOy is a stronger oxidant than compound BOz;

compound BOw is a stronger reductant than compound AOx;

the reaction between AOy and BOw is spontaneous:

Δ Δ Δ 0 2‐12

If AOx and AOy are solids while BOw and BOz are gases, then reaction (2-

12) can reach chemical equilibrium under a molar ratio of BOw:BOz

determined by the difference between Δ and Δ:

exp 2‐13

Figure 3. The modified Ellingham Diagram

The modified Ellingham Diagram can be divided into several zones based on the lines for reactions (2-6), (2-7), (2-14) (and (2-15), when CH4 is used as the fuel).

2 2 2‐14

2 24 2‐15

The zone above all these lines, labeled as “Full Oxidation”, covers the reactions for the formation of highly oxidative compounds. These compounds can thermodynamically convert CH4, CO, and H2 into CO2 and H2O, which is required for chemical looping combustion applications. Thus, compounds in this region, including Fe2O3, NiO, and

CaSO4, have been widely studied and used as chemical looping combustion oxygen carriers. The compounds near the top portion of the modified Ellingham diagram, such as

CuO and Mn2O3, can release molecular O2 at elevated temperatures. These compounds are used in the CLOU processes.

The zone below the line for reaction (2-14), labeled as “Inert”, covers the reactions associated with the compounds that cannot be used as the oxygen carrier in chemical looping systems. The metal oxides in this zone are weak oxidants that are unable to oxidize carbon to form CO. Equivalently speaking, reaction (2-16) for carbon deposition from CO is spontaneous:

2‐16

When CH4 is introduced into a reactor filled with AOy, the only spontaneous reaction that occurs above 600°C is the decomposition of CH4:

2 2‐17

Thus, AOy as the “oxygen carrier” cannot supply any oxygen to the fuel, and the chemical looping reactions cannot be completed. Nevertheless, metal oxides in this zone, such as MgO, Al2O3, SiO2, and TiO2, are widely used in a variety of chemical looping processes as the supporting material for another active metal oxide in the oxygen carrier material.

The zone between the lines for reaction (2-7) and (2-14) is labeled as “Partial

Oxidation”. The metal oxides associated with the reactions in this region are moderate oxidants that are suitable for chemical looping partial oxidation reactions. When these metal oxides are used as the oxygen carrier in the chemical looping system, reaction (2-1) is spontaneous, while reactions (2-2) and (2-3) are inhibited thermodynamically. Thus, these metal oxides are capable of converting the fuels into syngas (CO and H2) with minimal formation of full oxidation byproducts (CO2 and H2O).

A typical metal oxide in the “Partial Oxidation” zone is CeO2, which is widely studied as the oxygen carrier for chemical looping partial oxidation. Figure 4 shows the distribution of carbon in the equilibrium product of the reaction between CeO2 and CH4 at 900°C. Even at high CeO2 input, the carbon from CH4 can only be partially oxidized to form CO instead of being fully oxidized to CO2. Due to the Gibbs’s phase rule, when two different solid phases, i.e. CeO2 and Ce2O3, are present in the system, the equilibrium gas composition has zero degrees of freedom. Namely, the gas composition is not affected by the ratio between CeO2 and CH4 in the system. Thus, the amount of CO and CO2 plateau out when [O]/CH4 molar ratio is greater than 1.1 and both CeO2 and Ce2O3 are present as shown in Figure 4.

Figure 4. Carbon distribution in the equilibrium product of the CeO2-CH4 system at

900°C. [O]/CH4 represents the molar ratio between usable oxygen in CeO2 and CH4 feedstock.

As a comparison, Figure 5 shows the distribution of carbon in equilibrium product of the NiO-CH4 system. The NiO/Ni redox couple falls in the “Full Oxidation” zone of the modified Ellingham diagram. Thus, NiO is capable of converting CH4 to CO2 as desired in chemical looping combustion applications. As shown in Figure 5, when both

NiO and Ni are present in the system, the equilibrium gas product is mostly CO2. CO is the dominant product only with a [O]/CH4 molar ratio below 2.5. With an increasing NiO input, the amount of CO2 increases as CO is further oxidized by NiO. Note that with a

[O]/CH4 molar ratio below 4, NiO will be fully reduced to metallic Ni at chemical equilibrium, which is a catalyst for the decomposition of CH4. Thus, extra care has to be

taken in order to prevent carbon deposition in chemical looping systems using NiO as the oxygen carrier. The prevention of carbon deposition will be further discussed in the

“Reactor Design Considerations” section.

Figure 5. Carbon distribution in the equilibrium product of the NiO-CH4 system at

900°C. [O]/CH4 represents the molar ratio between usable oxygen in NiO and CH4 feedstock.

Despite the thermodynamic potential for the full oxidation of fuels, the oxygen carriers in the “Full Oxidation” zone of the modified Ellingham diagram can still be used for chemical looping partial oxidation system by controlling the kinetic aspects of the reaction. In this case, the operation of the reducer of the chemical looping system must ensure that only a limited amount of oxygen is transferred to the fuel. This can be achieved by reducing the ratio between the fuel and the oxygen carrier entering the

reducer, lowering the reactor temperature, and/or shortening the gas-solid contact time in the reducer. However, incomplete fuel conversion may be caused.

Fe2O3-based oxygen carriers are of particular interest for both chemical looping combustion and partial oxidation applications due to its high oxygen carrying capacity, low cost, and low toxicity compared to other oxygen carriers. The thermodynamics of the

Fe2O3 based oxygen carrier is more complicated due to the existence of multiple oxidation states. As shown in Figure 3, the Fe2O3/Fe3O4 and Fe3O4/FeO redox couples fall in the “Full Oxidation” zone, while the FeO/Fe redox couple falls in the “Partial

Oxidation” zone. The equilibrium product distribution of Fe2O3-CH4 system is shown in

Figure 6. Based on the iron phases present in the equilibrium product, Figure 6 is divided in to six zones labeled as “a” through “f”. The iron phases in each zone are listed in Table

3. Zone “b”, “d”, and “f” each has two solid phases in the equilibrium product. The equilibrium gas compositions in these zones do not vary with the Fe2O3 input flow rate as a result of the phase rule. In contrast, zone “c” and “e” each has only one solid phase, which allows the gas composition to vary between that of the adjacent zones.

When Fe2O3-based oxygen carriers are used for chemical looping combustion applications, the gas composition in zone “f” of Figure 6, with Fe2O3 and Fe3O4 in the solid product, is desirable as the fuels are converted to CO2 with little unconverted CO.

However, for chemical looping partial oxidation, the gas composition in zone “b” with Fe and FeO in the reduced oxygen carrier is suitable. Thus, the gaseous product should to be in contact with the Fe2O3-based oxygen carriers at proper oxidation states prior to exit in order to produce the desired product from the chemical looping process, which can be

achieved by controlling the ratio between the oxygen carrier and fuel feeding rates, and/or designing a proper gas-solid contact pattern in the reactor. The effect of the gas- solid contact pattern in the chemical looping reducer will be further discussed in the

“Reactor Design Considerations” section.

Figure 6. Equilibrium product distribution of the Fe2O3-CH4 system at 900°C with a CH4 flow rate of 1 mol/s.

Zone a b c d e f Iron FeO, Fe3O4, Fe Fe, FeO FeO Fe3O4 phases Fe3O4 Fe2O3 Table 3. Iron phases in each zone of Figure 6

Thermodynamics of the Iron-Titanium Composite Metal Oxide (ITCMO)

Figure 7 shows the phase diagram of the Fe2O3-CO system. At 900°C, the reduced oxygen carrier, consists of Fe and FeO, is in equilibrium with a gaseous product with

72% CO and 28% CO2, corresponding to a CO:CO2 molar ratio of 2.6. When Fe2O3 is used as the oxygen carrier in a chemical looping partial oxidation process, the product syngas composition will converge to the equilibrium CO:CO2 molar ratio of 2.6 with sufficient residence time. In order to generate a syngas with lower amount of full oxidation products (CO2 and H2O), a modified oxygen carrier with different thermodynamic property can be used.

Figure 7. Phase diagram of the Fe2O3-CO system

The iron-titanium composite metal oxide (ITCMO) is developed as an oxygen carrier for the chemical looping partial oxidation process. When the ITCMO is reduced by a fuel, FeTiO3 will be produced as the Fe(II) compound instead of FeO. FeTiO3 is a weaker oxidant compared to FeO. It has a lower tendency to fully oxidize CO and H2 to form CO2 and H2O. Figure 8 shows the phase diagram of the ITCMO-CO system. FeTiO3 is stable over a larger rage of temperature and CO mole fraction as compared to FeO.

When both Fe and FeTiO3 are present in the solid, the gaseous product of the system at

900°C consists of 94% CO and 6% CO2. The CO:CO2 molar ratio is greater than 15, which is significantly higher than 2.6 in the Fe2O3 case. Thus, the ITCMO is a better oxygen carrier candidate for chemical looping partial oxidation applications.

Figure 8. Phase diagram of the ITCMO-CO system

The formation of FeTiO3 as the Fe(II) compound not only affects the equilibrium syngas composition from the reducer of the chemical looping partial oxidation process, 28

but also impacts the H2 production via steam-iron reaction in the oxidizer of a chemical looping system. As described in Chapter 1, the reduced Fe2O3-based oxygen carriers from the reducer, consist of Fe and/or FeO, can convert H2O into H2 in the oxidizer. In this reaction, the oxygen carriers can potentially be oxidized by H2O to the Fe3O4 oxidation state. However, FeTiO3 is more difficult to be oxidized as compared to FeO.

Consider the oxidation reaction of Fe to form Fe3O4. Thermodynamically, this reaction can be considered as the sum of two reactions, namely, the oxidation of Fe to form a Fe(II) compound, and the oxidation of Fe(II) to Fe3O4. Depending on the presence of TiO2 in the system, the intermediate Fe(II) compound can be FeTiO3 or FeO, as shown in Figure 9. Thus, the oxidation reaction of Fe can be regarded as the sum of reactions (2-

18) and (2-19), or the sum of reactions (2-20) and (2-21).

2 2 Δ 2‐18

6 2 Δ 2‐19

22 2 Δ 2‐20

6 2 6 Δ 2‐21

Because the overall oxidation reaction remains unchanged regardless of the presence of

TiO2, the Gibbs free energy changes of the above reactions satisfy:

3Δ Δ 3Δ Δ 2‐22

As shown in Figure 3, the line for reaction (2-20) lies below the line for reaction (2-

18), namely, Δ Δ. Thus,

Δ 3Δ Δ Δ Δ 2‐23

Δ Δ FeO

Fe Fe O 3 4

FeTiO 3 Δ Δ

Figure 9. Oxidation reaction of Fe with and without TiO2

Equation (2-23) indicates that FeTiO3 is more difficult to be oxidized than FeO.

Specifically to the chemical looping oxidizer, the oxidation of the Fe(II) compounds by

H2O is given in reactions (2-24) and (2-25).

3 Δ 2‐24

3 3 Δ 2‐25

The value of Δ and Δ can be determined from reactions (2-7), (2-19), and (2-21):

Δ Δ Δ 2‐26

Δ Δ Δ 3Δ Δ Δ Δ Δ 2‐27

More intuitively, the equilibrium constants for reaction (2-24) and (2-25) is related by:

exp exp exp exp 2‐28

The term exp can be regarded as the indication of the relative stability of

FeTiO3 compared to FeO, because Δ Δ is the Gibbs free energy change of reaction (2-29):

2‐29 30

The value of Δ Δ is positive, which indicates that FeTiO3 is more stable than

FeO. Thus, exp is smaller than unity. For example, at 900°C,

exp 0.15, 0.44, 0.0015. Thus, while FeO can be oxidized to Fe3O4 by excess H2O, it is very difficult to convert FeTiO3 into Fe3O4 using H2O in the oxidizer. The implication of this property on the operation of the oxidizer will be further discussed in the “Mass Balance and Thermodynamics in Moving Bed Reactors” section.

Reactor Design Considerations

Before we can actually conceive the reactor design for the chemical looping partial oxidation system, certain knowledge on the kinetic properties of the reaction between iron-based oxygen carriers and methane is essential. Figure 10 shows the weight change of a Fe2O3-based oxygen carrier pellet in a single reduction-oxidation cycle at 900°C. The pellet is reduced by methane in the reduction phase, and oxidized by air in the oxidation phase.

Figure 10. Weight change of a Fe2O3-based oxygen carrier pellet in a redox cycle

The reduction phase starts with a rapid decrease of weight due to the loss of oxygen in iron oxide, and sharply turned around when the weight reached about 89% of the original value. The weight of the pellet increases until the end the reduction phase as a result of carbon deposition on the pellet from reaction (2-17), the decomposition of methane. The weight gain diminishes at the beginning of oxidation phase as carbon is burned by oxygen in air, and the pellet restores its original weight as it is fully oxidized.

Carbon deposition has been observed in other chemical looping systems using methane as fuel.[44] The decomposition of methane (reaction (2-17)) can be catalyzed by metal such as Fe and Ni. The sharp turnover at the reduction phase in Figure 10 is attributed to the emergence of metallic Fe on the surface of the pellet after it is reduced to a certain degree.

Note that the goal of chemical looping partial oxidation is to convert the fuel into

CO and H2 by partial oxidation. Carbon deposition in the reducer reactor may hinder in the following ways:

Carbon layer formed on the surface of oxygen carrier particles may

deactivate the particle and impede the oxidation reaction;

Carbon formed in the micro channels on the oxygen carrier particles may

destroy the physical integrity of the particle;

Since a portion of methane is converted to carbon, the yield of gaseous

product decreases;

Solid carbon will be carried over to the combustor reactor and be burned to

CO2, increasing the greenhouse gas emission of the process.

Thus, attention must be paid to prevent carbon deposition in the reducer reactor.

There are two simple ways to mitigate carbon deposition:

Add a certain amount of oxidative gas, such as H2O and CO2, in the

feedstock

Control the degree of reduction, such that the amount of metal on the

oxygen carrier particles cannot catalyze a rapid carbon formation.

The first way requires additional capital and operational cost in order to generate the steam or recycle the CO2 from downstream processes. The second way, on the other hand, requires careful consideration on the reactor design and oxygen carrier conversion, as discussed below.

Fluidized bed reducers are widely used in chemical looping systems. Although a fluidized bed reducer design is advantageous for good heat transfer, the oxygen carriers are inherently well mixed in this reactor design, which promotes a wide oxygen carrier residence time distribution when operated as a continuous flow system. Thus, in this reactor, the fuel gas and product syngas may contact and react with oxygen carriers at undesirable oxidation states.

Full Oxidation Fe2O3/Fe3O4

Fe3O4/FeO FeO/Fe Fe Carbon Deposition

Figure 11. Internal age distribution of particles in a perfectly mixed reducer

To further illustrate this effect, consider a perfectly mixed fluidized bed reducer.

Figure 11 shows the internal age distribution of the oxygen carrier particles in such a reducer. As a result of the nature of a mixed flow reactor, a wide internal age distribution of oxygen carrier particles will occur. Applying the oxygen carrier conversion knowledge obtained from TGA studies (Figure 10) to the internal age distribution, two portions of oxygen carrier particles in the reducer are not desirable for the production of syngas:

under reduced particles consisting of Fe2O3/Fe3O4 or Fe3O4/FeO;

over reduced particles consisting of a large amount of Fe.

The first portion of particles will fully oxidize the fuel and syngas into CO2/H2O, while the second portion will cause carbon deposition when it reacts with carbonaceous fuels. Thus, the choice of oxygen carrier residence time in the reducer falls into a dilemma: a shorter particle residence time will result in more under reduced particles and produce more CO2/H2O; a longer particle residence time will increase the risk of carbon deposition.

The existence of Fe2O3 and Fe3O4 in fluidized bed reducers causes the difficulty in selecting the operation conditions. As mentioned in the “Thermodynamics of Oxygen

Carriers” section, when redox couples in the “Full Oxidation” zone of the modified

Ellingham diagram are present in the reducer of syngas generation, the reactor operation must limit the oxygen transfer from the oxygen carrier to the fuels. However, such limit will in turn hinder the conversion of the fuels. This is particularly significant when CH4 and/or solid fuels (which can generate CH4 and other hydrocarbons in pyrolysis) are used as the fuel, as the reaction between CH4 and the oxygen carrier is usually slower than the reaction between CO/H2 and the oxygen carriers. For example, among a number of reports on biomass gasification using fluidized bed reducers and Fe2O3-based oxygen carriers, 5%-15% of the syngas produced from these systems consisted of unconverted hydrocarbons, such as CH4.[50-54] These hydrocarbon compounds in the syngas stream would require additional processing with a steam reformer before the syngas can be used in downstream steps for chemical synthesis. When the operating temperature of the

chemical looping reactor system was raised in order to increase the conversion of the hydrocarbons in the biomass, the CO and H2 purity in the syngas product stream decreased because the Fe2O3 and Fe3O4 in the oxygen carriers further oxidized the syngas to CO2 and H2O.

The inherent solid mixing in the fluidized bed reducers also challenges the conversion of solid fuels. When carbonaceous solid fuels are fed into the reducer, pyrolysis of the fuel occurs and produces gaseous volatiles and chars. Due to the slow rate of solid-solid reaction, the conversion of char in chemical looping systems relies on its gasification via H2O and/or CO2. However, the solid mixing in the fluidized bed reactor causes the unconverted char to leave the reactor before being gasified. Thus, a chemical looping system with a fluidized bed reducer often requires a carbon stripper downstream the reducer. The carbon stripper is a small fluidized bed that separates the oxygen carriers and char particles based on the particle size difference. The unconverted char from the carbon stripper is recycle back to the reducer for further conversion. The construction and operation the carbon stripper will increase the system complexity and cost of the chemical looping system.

In addition to the challenges brought by solid mixing, fluidized bed reducers may suffer from the bypassing of reactive gases. The sizes of the oxygen carrier particles used in fluidized bed reducers usually fall in the Geldart Group B. Thus, gas bubbles occur in the fluidized bed reducer. Due to the inefficient mass transfer between the bubble and the emulsion phases in the fluidized bed, gaseous fuels and/or volatiles may escape from the

reactor before being converted by the oxygen carriers, resulting in incomplete conversion of the fuels.

An effective alternative to the fluidized bed reducer design is the co-current moving bed reducer. The oxygen carriers and the fuel feedstocks are introduced from the top of the reducer and flow downwards co-currently. The product syngas stream will be extracted from the reducer outlet located at the bottom of the reducer. Due to the plug- flow nature of the moving bed reactor, the residence time of the oxygen carriers in the moving bed reducer can be well controlled in a narrow range. Thus, over reduction of oxygen carrier particles can be avoided to prevent carbon deposition. On the other hand, syngas at the gas outlet is in contact with reduced oxygen carrier particles consisting of

FeO/Fe, which thermodynamically inhibits full oxidation of CO/H2.

The co-current moving bed reducer design is particularly advantageous for solid fuel conversion. Due to the absence of axial solid mixing, the char particles generated from solid fuel pyrolysis have a uniform, long residence time in the reducer. By properly designing the residence time of the moving bed reducer, the char can be fully converted before leaving the bottom of the reducer. When full conversion of char is not desired (as given in the “Solid Fuel Gasification in a Moving Bed Reducer” section in Chapter 3), the char conversion can be controlled by the residence time of the moving bed reactor, as well as the size of the solid fuel. In addition, bypassing of reactive gases in can be avoided in moving bed reducers due to the absence of bubbles and the packed nature of the solid bed. Thus, the co-current moving bed reducer is suitable for the full conversion of gaseous and solid fuels into high purity syngas with minimal full oxidation byproducts.

These advantages are confirmed by experiment studies on the co-current moving bed reducer as discussed in Chapter 3.

Mass Balance and Thermodynamics in Moving Bed Reactors

The mass balance and thermodynamics in the moving bed reducer play a vital role in the design and operation of chemical looping system. Figure 12 shows the thermodynamic equilibrium lines for the Fe2O3-H2 system at 900°C. Different oxidation state of the oxygen carrier can reach chemical equilibrium with gas products at different concentration. The oxygen carrier conversion is defined by:

⁄/ 2‐30 / where ⁄ is the molar ratio between usable O and Fe in unreacted oxygen carriers,

while / is this ratio in reduced oxygen carriers. The solid lines represent the equilibrium states: vertical lines correspond to the equilibrium between two solid phases and the gas phase with a fixed composition. The horizontal lines correspond to the equilibrium between one solid phase and the gas phase with a flexible composition.

Previous studies have examined the mass balance and thermodynamics in a moving bed reducer for the conversion of H2 to H2O in a chemical looping combustion process.[56] Assume that the gas and solid flows in the moving bed reducer are perfect plug flow. By considering the mass balance in a segment of the moving bed reducer, it was found that the oxygen carrier conversion and the H2 conversion along the reducer

correspond to a straight line segment in Figure 12. These line segments are called the

“operation line” of the reactor. Along the direction of gas/solid flow, the conversions of gas and oxygen carrier vary along the operating line. The slope of the operating line is proportional to the ratio between gas and oxygen carrier flow. The operating lines for co- current gas-solid flow reactors have positive slopes and that for counter-current gas-solid flow reactors have negative slopes.

Figure 12. Thermodynamic equilibrium lines and operation lines for moving bed reducer and oxidizer using Fe2O3 oxygen carrier at 900°C

For example, consider the red dotted line in the “Reducer Region” of Figure 12 for a counter-current moving bed reducer in chemical looping combustion. In this reactor,

Fe2O3 enters from the top and leaves from the bottom, while H2 is introduced from the bottom and product gas leaves from the top. The right-most point on the line, corresponding to Fe2O3 and almost pure H2O, represents the top of the moving bed reducer. The left-most point on the line, corresponding to a mixture of Fe and FeO in the solid phase and pure H2, represents the bottom of the moving bed reducer. Thus, this operation line represents a reducer that fully converts H2 to H2O, while the Fe2O3 oxygen carrier is reduced to a mixture of Fe and FeO with an oxygen conversion of about 40%.

Note that the solid lines in Figure 12 represent conditions at chemical equilibrium, namely, conditions without any driving force for chemical reactions. Thus, the operation line of a reactor cannot go across the solid lines. Further increasing the H2 flow rate into the reducer, corresponding to increasing the slope of the operation line, can further increase the conversion of oxygen carrier. However, the H2 conversion at the gas outlet, located at the top of the counter-current reducer, will be reduced. On the other hand, the green dash line in the “Reducer Region” represents a co-current reducer. As shown in the figure, in order to fully convert the H2 fuel into H2O, the Fe2O3 oxygen carrier can only be reduced by 11%. Thus, a counter-current moving bed reducer can maximize the oxygen carrier conversion while achieving a full conversion of the gaseous fuel.

The analysis above can be applied to the operation of a moving bed oxidizer, as shown in the “Oxidizer Region” in Figure 12. The operation lines for the oxidizer start at the same oxygen carrier conversion as the reducer outlet, because the oxygen carriers

feeding into the oxidizer come from the outlet of the reducer. For the same oxygen carrier conversion (from Fe/FeO to Fe3O4), co-current moving bed operation produces a mixture of H2 and H2O with about 25% H2 concentration, while H2 concentration in counter- current is about 35%. Thus, the steam flow rate and energy cost required for counter- current operation is lower than the co-current operation.

When the ITCMO is used as the oxygen carrier, the thermodynamics of the system is shown in Figure 13. The performance of the counter-current moving bed reducer using the ITCMO oxygen carrier is similar to the Fe2O3 oxygen carrier as shown in Figure 12.

The H2 fuel can be fully converted to H2O while the oxygen carrier can be reduced to a mixture of Fe and FeTiO3. However, the oxidizer operation is significantly changed. Due to a very small equilibrium constant for reaction (2-25) ( 0.0015, the reduced

ITCMO oxygen carrier cannot be practically converted to Fe3O4 by H2O in the oxidizer.

Thus, when H2 production from the oxidizer is desired in the chemical looping system, the ITCMO oxygen carrier is less efficient than a Fe2O3-based oxygen carrier.

Figure 13. Thermodynamic equilibrium lines and operation lines for moving bed reducer and oxidizer using the ITCMO oxygen carrier at 900°C

The mass balance analysis can also be performed on the co-current moving bed reducer of a chemical looping partial oxidation system using CH4 as the fuel. Assume that the gas and solid flows in the moving bed reducer are perfect one dimensional plug flow, and the chemical reactions are uniform in the radial direction. Then the reactor can be considered as a one dimensional system. Consider a co-current moving bed reducer using the ITCMO oxygen carriers and CH4 as the fuel, both of which are introduced from the top of the reducer and flow downwards. Assume that no carbon deposition occurs in

the reducer. The status of the reducer at a given location can be depicted by six variables: the conversion of the oxygen carrier, , and the molar flow rates of five species, denoted by [CH4], [CO], [CO2], [H2], and [H2O]. At steady state, the molar flow rate of

Fe, denoted by [Fe], and the total carbon flow rate, denoted by [C], are not varying along the reactor. The variables above have to satisfy the equations given by mass balance, including:

1. Oxygen balance:

1.5 ⋅ 2 0 2‐31

2. Hydrogen balance:

2 2 2‐32

3. Carbon balance:

2‐33

By solving the above three equations, the status of the reducer at a given location can be determined by three independent variables. If we choose , [CH4], and [CO] as the independent variables, the other three variables are represented by:

2‐34

1.5 ⋅ 2 2 2‐35

1.5 ⋅ 4 4 2‐36

Thus, given a set of , [CH4], and [CO] values at a location in the co-current moving bed reducer, the values of [CO2], [H2], and [H2O], and thus the status of the moving bed reducer at that location, can be uniquely determined. Moreover, the status variables should be varying continuously along the moving bed reducer. Thus, the status of a moving bed reducer, under the assumptions above, can be represented by a curve in the 43

three dimensional space of , , . If the feedstock is fully oxidized ITCMO oxygen carriers and pure CH4, the curve will start from point , ,

0, ,0, and end at a point that corresponds to the reducer outlet oxygen carrier conversion and gas composition.

Naturally, all the molar flow rates should be non-negative:

0 2‐37

0 2‐38

0 2‐39

1.5 ⋅ 2 2 0 2‐40

1.5 ⋅ 4 4 0 2‐41

In addition, the conversion of oxygen carriers should fall between 0 and 100%:

0 100% 2‐42

The above inequalities, (2-37) through (2-42), define a three dimensional region in the space of , , , as shown in Figure 14. Specifically:

the equation 0 defines a plane that is

parallel to the axis and intercepts the other two axis at

, , 0,0, and 0, ,0;

the equation 1.5 ⋅ 2 2 0 defines a

plane that intercepts the three axis at , , 0, ,0,

0,0,2, and ,0,0;

the equation 1.5 ⋅ 4 4 0 defines a

plane that intercepts the three axis at , , 0, ,0,

0,0,4, and ,0,0.

Figure 14. Operation region of the co-current moving bed reducer in chemical looping partial oxidation

The operation line of a co-current moving bed reducer is a curve within the green region starting from the point , , 0, ,0, such as the red dash line in Figure 14. The actual shape of the line is determined by the mechanism and kinetics of the reactions occurring in the reducer.

Unlike the plots for H2 conversion in a moving bed as shown in Figure 12 and

Figure 13, which contain several connected line segments of equilibrium conditions, the plot for partial oxidation of CH4 has only one equilibrium point, as shown in Figure 14.

Thermodynamically, the status of a co-current moving bed reducer will be attracted by the equilibrium point. Thus, the operation line of the reducer will extend towards and converge to the equilibrium point as the residence time of the reactants in the reducer increases.

Figure 15. Operation region of the co-current moving bed reducer at complete CH4 conversion

Figure 15 shows the operation region of the co-current moving bed reducer at complete CH4 conversion, which is a slice of the region in Figure 14 at 0. The operation line of a co-current moving bed reducer will reach the green region in Figure

15 when CH4 is completely converted. As shown by the magenta lines in Figure 15, the equilibrium point can be decided by intercepting the chemical equilibria lines of the Fe-

Ti-CO system and the water-gas shift (WGS) reaction. The Fe-Ti-CO system involves the following reactions:

3 2 2‐43

3 3 2‐44

2‐45

The equilibrium lines of this system on Figure 15 are similar to the solid lines in Figure

13. The WGS reaction is:

2‐46

The equilibrium line can be plotted by expanding the definition of the equilibrium constant:

.⋅ ≅ 2‐47 .⋅

This equilibrium line intersects the axis at , ,0, and passes through the intersection of the 0 line and the line.

The operation region of a co-current moving bed reducer is affected by the ratio

between CH4 and Fe2O3 in the feedstock, i.e. the value of . If the CH4:Fe2O3 ratio increases, the operation region on Figure 15 shrinks towards the upper right corner. In the meantime, the equilibrium point will move upwards and result in a higher oxygen carrier conversion, . However, the value of [CO] maintains unchanged, which corresponds to a constant product syngas composition. In contrast, if the CH4:Fe2O3 ratio decreases, the equilibrium point may move to the horizontal section of the Fe-Ti-CO system equilibrium lines, which corresponds to the formation of FeTiO3 as the only phase of reduced iron.

The syngas composition may vary because the value of [CO] is no longer a constant with

respect to a varying value of .

Note that the equilibrium point can also be determined by minimizing the Gibbs free energy of the system under given oxygen carrier and fuel input. Thus, the product of a co- current moving bed reducer with sufficient residence time can be found by performing

the Gibbs free energy minimization with the given feedstock, which can be readily achieved by using the RGibbs module in ASPEN process simulation software. In Chapter

3, all thermodynamically expected results are calculated by calling an RGibbs module in

ASPEN under the specific condition.

Conclusion

The thermodynamics of oxygen carriers in chemical looping processes are examined. The modified Ellingham Diagram is constructed for the selection of oxygen carriers. The Diagram is separated into zones for “Full Oxidation”, “Partial Oxidation”, and “Inert” based on the thermodynamic tendency of the oxidation reactions in relation to the oxidation reactions of CH4, C, CO, and H2. The oxygen carriers from the partial oxidation region are thermodynamically favorable for the chemical looping partial oxidation process.

Specifically, the thermodynamic property of the ITCMO oxygen carrier is analyzed.

The addition of titanium in the system stabilizes the Fe(II) compound. When the ITCMO oxygen carrier is reduced to the Fe/FeTiO3 state, it can reach chemical equilibrium with a gaseous product with high CO and H2 content, which avoids the further production of full oxidation byproducts. Therefore, the ITCMO oxygen carrier is suitable for high purity syngas generation in the chemical looping partial oxidation process.

The design of the chemical looping partial oxidation reducer is analyzed. Although fluidized bed reducer design has been used widely for chemical looping systems, it suffers from the inherent solid back-mixing and gas bypassing. Solid back-mixing causes a wide distribution of internal age of the oxygen carriers, and hence, a distribution of oxidation states in the reactor. High oxidation state oxygen carriers promotes undesirable full oxidation reactions, while the over-reduced oxygen carriers may cause carbon deposition. In addition, solid back-mixing results in the bypassing of unconverted char to the combustor. Gas bypassing due to the insufficient mass transfer between the gas and solid phases may also result in low fuel conversion.

A co-current moving bed reducer design is chosen for the chemical looping partial oxidation process. Due to the plug flow nature of gases and solid in the moving bed reducer, the gas and solid residence time at the reducer outlet can be well controlled.

Thus, gaseous product is in contact with oxygen carriers with desirable oxidation states before leaving the reducer, which is advantageous for preventing full oxidation and carbon deposition. The absence of axial solid mixing and gas bubbles also facilitates the full conversion of solid and gaseous fuels in the moving bed reducer.

The effect of ITCMO on the operation of the oxidizer is analyzed. Because the formation of FeTiO3 stabilizes the Fe(II) compound, it is more difficult to be oxidized by

H2O to form H2 in the oxidizer. The operation line analysis based on the mass balance and thermodynamic property of the moving bed oxidizer shows that the reduced ITCMO oxygen carrier can only be oxidized to the FeTiO3 oxidation state in the oxidizer, while the reduced Fe2O3 oxygen carrier can be oxidized to the Fe3O4 oxidation state. Thus,

ITCMO is less efficient for H2 generation from the oxidizer as compared to a Fe2O3 oxygen carrier.

The mass balance analysis is extended to consider the co-current reducer for the partial oxidation of CH4. A feasible operation region is identified in a 3-dimensional phase space that represents the status of the reducer. The operation line of the co-current reducer is a curve within the feasible operation region. Although the actual shape of the curve is determined by the reaction kinetics of the system, the curve extends towards the equilibrium point of the system with sufficient residence time. This analysis indicates that the product from the co-current moving bed reducer with sufficient residence time can be simulated by a single-stage equilibrium model, such as the ASPEN RGibbs module that calculates the equilibrium state under the given reactant input.

Chapter 3: Experimental Studies of Chemical Looping Partial Oxidation

As described in Chapter 2, the ITCMO oxygen carrier is thermodynamically favorable for the partial oxidation of carbonaceous fuels to generate syngas. The co- current moving bed reducer is suitable for the conversion of fuels to high purity syngas with minimal full oxidation byproducts. The feasibility and advantages of using the

ITCMO oxygen carrier and co-current moving bed in the chemical looping partial oxidation process are confirmed by experimental studies performed in fixed bed and moving bed reactor systems. This chapter discusses the experimental studies on the reactivity of the ITCMO oxygen carrier, the conversion of carbonaceous fuels to syngas, and the scale up of the chemical looping partial oxidation reactor.

Reactivity of the ITCMO

The long-term operation of chemical looping process requires excellent oxygen carrier recyclability, namely, the capability of maintaining reactivity and oxygen carrying capacity. Cyclic reduction-oxidation reactions in a thermal gravimetric analyzer (TGA) are carried out in order to confirm the recyclability of the ITCMO material.

A sample of ITCMO oxygen carrier, weighing 56.79 mg, is placed in a SETARAM

SETSYS Evolution TGA. Helium is used as a carrier gas at a flow rate of 50 ml/min. The sample is heated to 950°C, at which 50 continuous reduction-oxidation cycles are carried out. In each cycle, 200 ml/min of CH4 is used to reduce the sample for 8 mins, and 100 ml/min of air is used to regenerate (oxidize) the reduced powder for 15 mins. 50 ml/min

N2 is used to flush the tubes and reactors between stages. The gas flow rates are regulated by ALICAT mass flow controllers. The gravimetric measurement data is used to calculate the oxygen carrier conversion as defined in equation (2-30).

Figure 16. Recyclability test on the ITCMO oxygen carrier

Figure 16 shows the result for the recyclability test. It is observed that during the stage of reduction, the ITCMO lost more than 80% of its usable oxygen. The change of oxygen capacity varies slightly among cycles, and no decrease in reactivity was observed over the 50 cycles. The oxygen capacity is fully restored in every oxidation stage after reduction. It can be confirmed that the ITCMO material possesses excellent recyclability over reduction-oxidation cycles.

The ITCMO can convert the tar and volatiles from the solid fuels, such as coal and biomass, into syngas.[57] The tar and volatile cracking effect of oxygen carriers is studied with fixed bed experiments. The fixed bed reactor used for these experiments is a cylindrical Inconel tube embedded vertically in a tubular furnace and heated to 900°C, as shown in Figure 17. ITCMO oxygen carrier particles are held by a fine net located at the lower section inside the reactor. As shown in Figure 17, biomass powder is stored in the lock hopper section on top of the reactor, namely the “Biomass Container”. At the start of the experiment, the valve to the lock hopper is opened to drop the biomass into the reactor. Gaseous product exiting the fixed bed reactor at the bottom along with the carrier gas He (100 ml/min) is sampled every 100 seconds by a Pfeiffer Vacuum GSD 300 C

Omnistar Mass Spectrometer (MS). To prevent the products from condensing before entering the MS, all tubes between the reactor and the MS are insulated, and the inlet of the MS is heated externally to 200°C. Species in tar with mass/charge ratio ranging from

2 to 200 are be detected by the MS.

Figure 17. Fixed bed experiment setup for biomass tar cracking

Two experiments are conducted in the fixed bed reactor to illustrate the tar cracking effect of the ITCMO oxygen carrier. In the first experiment, the biomass is dropped into the heated empty reactor. In the second experiment, the reactor is filled by 25g of

ITCMO oxygen carriers. The biomass is dropped onto the heated bed of oxygen carriers.

The mass spectra of the gaseous product from the fixed-bed reactor experiments for the two cases, with and without oxygen carriers, are shown in Figure 18. When no oxygen carrier is present in the reactor, a significant amounts of tar, along with CO and

CO2, is generated in the pyrolysis of the biomass feedstock. Comparison of the two mass

spectra shows a dramatic difference in the variety and amounts of chemical species present in the product gases of the two experiments. In the experiment with ITCMO oxygen carriers, the effluent from the fixed bed consisted mainly of CO2. Note that the vertical axes in Figure 18 are on logarithm scale. Thus, the MS signals for most of the species are reduced by at least one order of magnitude. The MS data shows that the

ITCMO oxygen carriers can effectively crack and oxidize the biomass derived tar.

Figure 18. MS data from fixed bed biomass pyrolysis with and without oxygen carriers

Syngas Generation in a Fixed Bed Reactor

The TGA experiments conducted on the oxygen carriers can provide important information on the reaction kinetics of the oxygen carrier reduction by methane and the recyclability during the redox cycles. However, the gas product cannot be effectively

monitored in the TGA experiments. A fixed bed experiment is conducted to investigate the gas product composition during the methane to syngas reaction. In order to simulate the co-current moving bed reducer, the lower section of the fixed bed reactor is filled with 23.1g pre-reduced ITCMO oxygen carrier particles, while the upper section is filled with 8.3g unreduced (oxidized) oxygen carrier particles, as shown in Figure 19.

Figure 19. Fixed bed experiment setup for CH4 conversion

The fixed bed reactor is embedded vertically in an electric furnace and heated externally. When the temperature of the reactor reached 990°C, 50mL/min CH4 and

50mL/min N2 are introduced into the reactor by mass flow controllers. The gas composition from the outlet is analyzed by a CAI gas analyzer (California Analytics

Instrument Infra-red Analyzer ZRE-600). The CAI analyzer enables an instant

measurement of three gas components (CO, CO2 and CH4), with the range limitations of

10% for CO, 100% for CO2 and 20% for CH4. Due to the detection range limitation (10 vol% CO) of the CAI gas analyzer, the gas product is diluted by 400 mL/min nitrogen prior to being sent to the CAI gas analyzer. When the gas composition reaches a semi- steady state and is stabilized for a period of time under the first reaction condition, the total gas flow rate sent to the reactor is decreased to 60mL/min (CH4 30mL/min, N2

30mL/min) in order to investigate the effect of the gas residence time. When another semi-steady state is achieved under the second reaction condition, the temperature of the reactor is increased to 1050°C to evaluate the temperature effect.

The concentration of carbonaceous species in the gaseous product is shown in

Figure 20. The conversion of methane decreases from about 90% to lower than 50%, and remained low for about 20 min before methane concentration starts to decrease significantly. The decrease in CH4 conversion is probably caused by the formation of low reactivity compounds, such as Fe3O4 and FeTiO3.

After the conversion of methane gradually increased to about 90%, the flow rate of methane is decreased to increase the residence time of methane. An immediate increase in methane conversion is observed. A steady production of syngas is maintained for 10 min, until the temperature of the reactor is increased to 1050°C to evaluate the temperature effect. A slight improvement in conversion was achieved at evaluated temperature. It should be noted that during the period of syngas generation, the ratio between CO and CO2 was mildly fluctuating around 10, indicating a high quality of syngas. Thermodynamic analysis indicates that methane could be fully converted at the

experiment condition. However, the improvement in the methane conversion at a longer gas residence/elevated temperature implies the existence of a kinetic restriction.

Figure 20. Product gas composition of fixed bed reactor with CH4 injection. (1) lowered gas flow rate; (2)increased temperature.

The kinetic nature of the reaction between the ITCMO oxygen carriers and the carbonaceous species is further explored by another fixed bed experiment in which the reduced iron oxide from the previous experiment is subjected to oxidation by CO2. A stream of 50% CO2 (balanced by N2) is introduced to the reactor. About 1000s after the experiment starts, the input gas is switched to a pure CO2 stream at the same flow rate

(and diluted by nitrogen before entering the gas analyzer) to examine the effect of kinetic factors. The following reactions occur:

3‐1

3‐2

The concentration of carbonaceous species in the gas product is shown in Figure 21.

The change in the CO2 flowrate had little effect on the product concentration, indicating that the gas composition was controlled by thermodynamics instead of kinetic factors.

Recall that the gas composition is affected by kinetic factors in the fixed bed experiment with CH4 injection. Therefore, the reaction between CH4 and oxygen carrier is likely to be the rate limiting step in the reducer for CH4 partial oxidation.

Figure 21. Product gas composition of fixed bed reactor with CO2 injection. (1) increased

CO2 concentration

A sharp change occurs in the product composition at about 2700s on Figure 21. The flip-over of the concentration of CO and CO2 illustrates the dependence of this reaction on solid composition. As long as the ITCMO oxygen carriers remain a mixture of Fe and

FeTiO3, the concentration of CO is way higher than that of CO2, and syngas could be generated with high quality; when the ITCMO oxygen carriers are oxidized into a mixture of FeTiO3 and Fe3O4, the CO concentration would immediately drop to a very low level. These two sets of experiments indicated that methane reaction with oxygen carrier is relatively slow. Once methane is converted, the thermodynamics of this system will determine the gas composition in the product gas.

CH4 Partial Oxidation in a Moving Bed Reducer

The chemical looping partial oxidation using the co-current moving bed reducer is further explored in a bench scale reactor system as shown in Figure 22. The moving bed reactor consists of a 40 inch long, 2 inch (5 cm) ID steel tube, which is heated externally by electric heaters, and a screw feeder installed at the bottom of the reactor. ITCMO oxygen carriers are constantly removed from the reactor by the screw feeder and falls into a container connected to the outlet of the screw feeder, while gaseous product flows out to the gas analysis system. A tilted pipe with a lock hopper is installed at the top of the reactor, from which oxygen carrier particles could be added to the reactor. A glass

window is located below the solid feeding pipe to monitor the solid level. Type K thermocouples are installed along the reactor to measure the temperature profile.

Figure 22. Bench scale co-current moving bed reducer

The electric heaters used on the bench scale reactor system include three pairs of heating elements, located outside the top, middle, and bottom section of the moving bed reactor. The three pairs of heating elements are controlled by the temperature readings

from three thermal couples located at 10, 20, and 30 inches from the top of the reactor, respectively.

Figure 23 shows a schematic electrical connection diagram of the bench scale reactor system. The thermal couple and gas analyzer readings are processed by data acquisition (DAQ) cards, which communicate with a computer. The computers determines the heater output power based on the thermal couple readings, and convert the desired heater power, in percentage, to DC pulse-width modulation (PWM) signals. The

PWM signal consists of a square wave with a 2-second period and a varying duty cycle.

The duty cycle, defined as the ratio of the peak period to the total period, is set to the desired heater power output. The PWM signals control three solid state relays (SSR), which turn on and off the 208V AC circuits that powers the heating elements.

Figure 23. Schematic electrical connection diagram of the bench scale reactor system

Although the reactor temperature is set to a certain value, the actual temperatures along the reactor are unevenly distributed due to the structure of the electric heater and imperfect insulation. The range of variation is usually 100 C from the heater set-point. An exemplary temperature profile, when the reactor temperature is set at 1050°C, is shown in

Figure 24. The setpoint, in fact, represents the highest temperature in the reactor.

30 Distance from (inches)top

800 850 900 950 1000 1050 1100 Temperature (oC)

Figure 24. Exemplary temperature distribution in the bench scale reactor

Opposite to the thermocouples on the reactor, eight gas sampling ports are set to enable the sampling of solid and gas from different parts of the reactor. Gaseous product from the reactor is sent to a gas analysis system. The gas is first cooled with air and dried with desiccant. Afterwards, the gas is sent through a gas analyzer (California Analytics

Instrument Analyzer ZRE–600), which detects CO, CH4, CO2, SO2, and O2, and sampled by a gas chromatography (Varian CP4900 micro-GC) before it is vented. The data collected by the analyzer and the GC is sent to the computer and logged automatically.

The oxygen carrier particles used in the moving bed reactor are spherical ITCMO particles of 1.5mm diameter. Before the experiment, the moving bed reactor is filled with the ITCMO oxygen carriers. The screw feeder is calibrated to convey 20g/min of oxygen carriers before the heaters start. The reactor is flushed with N2 and heated to the desired reaction temperature with the bed moving before the injection of fuel gas. A mixture of 65

90% CH4 and 10% N2 is introduced in a flow of 2 standard liters per minute (SLPM) as the fuel.

There are three parameters that are particularly important to evaluate the performance of natural gas to syngas, i.e., methane conversion, CO:CO2 molar ratio, and

H2:CO molar ratio. These parameters are calculated as follows:

1 ,,

1, , ,

: ,⁄,

: ,⁄, where ,is the molar flow rate of CH4 from the gas outlet, ,is the molar flow rate of CH4 introduced to the reactor, is the molar flow rate of N2 introduced to the reactor, ,is the concentration of component A in gaseous product measured by the micro-GC.

After the experiment, the reduced oxygen carrier particles were examined in a carbon analyzer, where the particles were heated to 900°C in flowing oxygen. The effluent gas was titrated in an electrolytic cell, which determined the cumulative CO2 content.

The product produced in the co-current moving bed reducer is highly dependent on the solids profile in the last stage of the reducer. Initially, the bed is filled with fully oxidized ITCMO oxygen carriers. Thus, CO2 is the dominant product from the reducer 66

after the CH4 injection starts. The production of CO2 is maintained while the bed of

ITCMO oxygen carriers is gradually reduced. After the bottom stage of the reducer is converted to Fe(II), the gaseous product from the reducer gradually transits to a syngas, which is dominated by CO and H2. The gas composition during the steady-state syngas production period with a temperature setpoint of 1040°C is shown in Figure 25. The values of methane conversion, CO/CO2 molar ratio, and H2/CO molar ratio are shown in

Figure 26. Under this experimental condition, the conversion of methane can reach over

99%, with a CO:CO2 ratio of ~9 and a H2:CO ratio of ~2. Throughout the experiment, the methane conversion is increasing as metallic iron formed later could catalyze the methane reaction with ITCMO material. Furthermore, the carbon analysis on reduced particles showed there was no carbon deposition on the particles.

50 H2 N2 CH4 CO

20 Gas Concentration(%) 10

0 0 1000 2000 3000 4000 Time (s)

Figure 25. Syngas composition from the bench scale reducer at 1040°C using methane as feedstock

Figure 26. Quality of syngas from the bench scale reducer at 1040°C using methane as feedstock

The gas composition variation along the moving bed reactor is studied by taking samples from the gas sampling ports. The composition profile at temperature setpoint of

1020°C is shown in Figure 27. Note that the setpoint represents the highest temperature in the reactor. Thus, the actual temperatures in other parts of the reactor are lower than

1020°C. The gas composition profile indicates the need for a longer residence time in the reactor. At higher temperature, this profile is “contracted” towards the inlet of the reactor, and the gas and solid composition reached the final value sooner. Thereby, higher CH4 conversion and/or shorter residence time can be achieved.

90 80 70 60 50 H2 40 N2 30 CH4 Concentration(%) 20 CO 10 0 5 10152025303540 Distance from top (inch)

Figure 27. Gas composition at different locations of the bench scale reducer at 1020°C

The reduced oxygen carriers discharged from the moving bed reducer is characterized using an X-Ray Diffractometer (XRD). Figure 28 shows the XRD pattern of the reduced ITCMO oxygen carrier. FeTiO3 and Fe are detected in the reduced oxygen carriers. The solid composition agrees with the gas analysis results and the post- experiment solid analysis confirms the feasibility of using the co-current moving bed reducer for the partial oxidation of CH4.

Figure 28. XRD pattern of the reduced ITCMO oxygen carrier

Solid Fuel Gasification in a Moving Bed Reducer

Similar to the conversion of gaseous fuel, solid fuels can also be converted in co- current moving bed reducers for syngas generation. Commonly used solid fuels include coal and biomass. When the solid fuels are introduced into the reducer, they are heated by the high temperature oxygen carriers, resulting in the pyrolysis of the fuels. Volatiles and tars produced from the pyrolysis flow downwards through the bed of oxygen carrier, and are partially oxidized to produce a mixture of CO, CO2, H2, and H2O. Solid char produced from the pyrolysis flows downwards along with the oxygen carriers.

Gasification of solid char occurs along the moving bed reducer as the CO2 and H2O react with the char by the reverse Boudouard reaction (3-3) or the water-gas reaction (3-4).

2 3‐3

3‐4

Due to the chemical composition of coal and biomass, the syngas produced from these solid fuels usually contains more CO than H2. However, downstream chemical synthesis processes may require a syngas stream with higher H2 content. For example, the synthesis of acetic acid requires syngas with a H2:CO molar ratio of 1:1, while the synthesis of methanol and Fischer-Tropsch liquid fuel requires a H2:CO molar ratio of

2:1.

In a conventional coal or biomass gasification process, a WGS reactor is usually used to convert a portion of the CO to H2 by the catalytic reaction (2-46) with H2O. In a chemical looping partial oxidation system, H2O can be added into the reducer to adjust the product composition, which reduces the capital and operational cost for the plant.

As described in the “Mass Balance and Thermodynamics in Moving Bed Reactors” section of Chapter 2, the reactants in a co-current moving bed reducer will converge to the equilibrium product composition. Similar arguments can be applied to a moving bed reducer using solid fuels. Thus, the thermodynamic interaction between the ITCMO oxygen carriers and the solid fuel feedstock in the co-current moving bed reducer can be analyzed using an ASPEN RGibbs reactor model.

Consider the biomass partial oxidation in a co-current moving bed reducer. The composition of the biomass used in the simulation as well as the following experimental

study are shown in Table 4. The composition of the as-received biomass feedstock, in terms of atomic ratio of elements, can be written as CH1.57O0.68. The RGibbs model is configured to calculate the reducer gas and sold equilibrium products at 1,000°C and 1 bar (abs). The biomass feed rate in the RGibbs reactor model is set to provide 1 mol/min of carbon, while the feed rate of ITCMO oxygen carrier is varied to reveal the effect of oxygen supply on the product syngas composition.

Proximate Analysis, wt%, as received Moisture 5.52 Volatile 78.06 Fixed Carbon 15.85 Ash 0.57 Ultimate Analysis, wt%, dry basis C 51.55 H 6.09 N 0.15 S 0.12 O 41.49 Table 4. Proximate and ultimate analysis results for the wood pellets

Figure 29 shows the product gas composition from the RGibbs reducer model. The horizontal axis shows the molar flow ratios of the available lattice oxygen in oxygen carriers to the carbon content in the biomass ([O]:CFeed ratio). In the ITCMO oxygen carrier, 1 mol of Fe corresponds to 0.5 mol of Fe2O3 or 1.5 mol of available lattice oxygen. The Fe(0), Fe(II), and Fe(II,III) compounds shown in Figure 29 correspond to metallic Fe, FeTiO3, and Fe3O4, respectively.

When at the thermodynamic equilibrium with both Fe and FeTiO3 phases, the

CO:CO2 ratio in the FeTiO3/Fe system is greater than 14. By maintaining the [O]:CFeed ratio to be less than 1.5, a syngas selectivity of >90% can be achieved from the co-current moving bed reducer. Further increasing the [O]:CFeed ratio hinders the reduction of

FeTiO3, leaving FeTiO3 as the only form of reduced oxygen carrier. Thereby, a larger amount of CO2 and H2O will be produced.

Figure 29. Gaseous product from BTS reducer at equilibrium

Because the atomic ratio of C:H in the biomass feedstock is 1:1.57, the product syngas from the RGibbs reducer model contains more CO than H2. By adding H2O into the reducer, the H2 content in the product syngas can be adjusted to the desired value.

Figure 30 shows the H2:CO ratio of the syngas product from the RGibbs reducer model under different molar ratios between H2O and C from biomass (H2O:C ratio). The H2:CO 73

ratio in the product syngas can be increased to about 1.5 with a H2O:C of 1, and further increased to above 2 with a H2O:C ratio of 2. Note that the H2:CO ratio for the cases with

H2O injection is not a constant, indicating that FeTiO3 is the only phase of the reduced iron in the oxygen carrier, instead of both Fe and FeTiO3.

2.5

1.5 :CO 2 H 1

H2O:C=0 0.5 H2O:C=1 H2O:C=2 0 0.511.522.5

[O]:CFeed

Figure 30. Effect of H2O injection to the reducer on H2:CO ratio at 1000°C

The conversion of biomass to syngas is tested in the bench scale moving bed reducer shown in Figure 31. The reactor system is improved to enable H2O injection and handle the ashes in solid feedstock. The biomass pellets are mixed with the ITCMO oxygen carriers at a designated mass ratio, and introduced into the reactor from the lock hopper at the top. A syringe pump feeds deionized water into reducer. The water is fed into a tube that passes through the electric heater and leads into the top section of the reactor, in

which the water is vaporized to generate steam. The product gas outlet locates at the bottom of the reactor, which is connected to a water trap and a desiccant bed for ash and

H2O removal prior to entering the gas analyzers. An infrared (IR) analyzer (Siemens,

ULTRAMAT 23) for CO, CO2, CH4, and O2 characterization, followed by and a thermal conductivity detector (TCD) analyzer (Siemens, CALOMAT 6) for H2 analysis, are installed in series.

Figure 31. Bench scale co-current moving bed reducer for solid fuel conversion

Wood pellets supplied by BMQ Inc. are used as the feedstock in the experiments.

The cylindrical pellets have a diameter of about 5mm and a length of 3-15mm. The proximate and ultimate analysis of the wood pellet is shown in Table 4.

The moving bed reactor is initially filled with the ITCMO oxygen carriers and heated to 1000°C with the screw feeder set to an oxygen carrier flow rate of 20 g/min.

After the temperature setpoint is reached, the mixture of biomass pellets and ITCMO oxygen carriers are introduced from the lock hopper. Deionized water is injected into the reducer. The water flow rate is adjusted to vary the H2:CO ratio of the syngas produced.

The two experimental conditions are shown in Table 5.

Test Condition 1 2 Temperature 1000 °C 1000 °C [O]:CFeed* 2.0 1.7 H2O:CFeed* 1.5~1.8 1.1~1.8 Table 5. Summary of experiment conditions for biomass gasification reducer tests

The mole fractions of CO, CO2, CH4, and H2 in the dried syngas are measured by the gas analyzers. N2 is introduced at the top of the reducer to serve as an inert flush gas.

The dry, N2 free mole fraction of gas specie i, or , is calculated by:

⁄ (3-5)where is the mole fraction of CO, CO2, CH4, or H2 as measured by the gas analyzers.

For Test Condition 1, Figure 32(a) shows the molar ratio between the H2O and C from biomass that are fed into the reducer, and Figure 32(b) shows the N2 free syngas

composition at the reducer gas outlet. The instantaneous spikes on the gas composition plot are caused by the air entrained into the reactor during solid makeup.

(a)

(b)

Figure 32. (a) Molar ratio between steam and carbon in wood pellet and (b) N2 free syngas composition and dry syngas purity under Test Condition 1

The evolution of oxygen carrier conversion along the moving bed reactor can be inferred from the variation of syngas composition. Before the experiment, fully oxidized

ITCMO oxygen carriers, which are capable of fully oxidizing the feedstock to CO2 and

H2O, are loaded into the reactor. When the biomass pellets are first introduced into the heated zone of the reactor, the products of biomass pyrolysis are fully oxidized by these oxygen carriers. Thus, the initial gaseous product from the reducer is dominantly CO2 as shown in Figure 32(b) from time period 0 to 60 minutes. As more biomass feedstock enters the reducer, the lattice oxygen in the ITCMO oxygen carriers is gradually depleted to produce a lower oxidation state. The lattice oxygen at the higher portion of the moving bed reducer is consumed first while the bottom portion of the bed remains at high oxidation state. Consequently, nearly pure CO2 is observed in the product gas at the outlet of the reducer over the course of the first 60 minutes of operation.

After the ITCMO oxygen carrier near the bottom portion of the moving bed reducer is reduced to FeTiO3, the oxygen carriers are no longer thermodynamically favored for full fuel oxidation. Thus, the gaseous product at the reducer outlet transits from a stream of CO2 to a mixture of syngas and CO2. The transition to syngas generation is observed in

Figure 32(b) from time 60 minutes to 70 minutes. If the operating condition is held steady, the product gas composition from the reducer will reach a steady state consisting predominantly of CO and H2.

At the 107th minute, the water pump is started to inject water into the reducer. The gas outlet composition responds rapidly to the addition of steam as shown by the significant increase in the H2 and CO2 mole fraction and the decrease in CO mole

fraction. The syngas purity, defined by the sum of the mole fraction of H2 and CO in the gas product, decreases from >80% to ~67% due to the increase in CO2. The effect of H2O flow rate on the gas composition can be seen from Figure 32(b). When H2O flow rate is

th th decreased at the 156 minute and the 187 minute, the H2 and CO2 mole fractions decrease while the CO mole fraction increases. Over the course of the experiment, the dry syngas purity is maintained between 65% and 72%. The H2:CO ratio varies between 1.8 and 2.3, while the CO:CO2 ratio varies between 0.6 and 1.0.

The mole fraction of CH4 in syngas also increases after H2O is injected into the reactor. This is attributed to two factors adversely affecting CH4 conversion in the moving bed reducer.

The additional flow of H2O causes an increase in the gas velocity inside the

reactor, which reduces the residence time of the gaseous reactants.

The high concentrations of H2O and CO2 in the reactor inhibit the further

reduction of FeTiO3 to Fe. Metallic Fe is a catalyst for the decomposition of

CH4.[58] The absence of metallic Fe reduces the reaction rate between CH4

and the ITCMO oxygen carriers.

Nevertheless, the mole fraction of CH4 in the product syngas is less than 3%, which indicates an almost full conversion of the gaseous volatiles and tars.

Unconverted wood char is observed in the discharged oxygen carrier particles. The incomplete carbon conversion is a result of the large size of the wood pellets fed into the reducer. The large char size reduces the surface area for the gasification reactions (3-3) and (3-4), resulting in slow reaction kinetics. Thus, the gasification of solid char is the

potential rate limiting step in the full conversion of solid fuels in the co-current moving bed system.

Similar trends of gas composition variation are observed from the experiment with

Test Condition 2, as shown in Figure 33. Initially, CO2 is the dominant gas produced from the reducer before transitioning to the syngas generation phase. After the water pump is started, the mole fraction of the H2 and CO2 species increases with CO mole fraction and syngas purity decreases. Test Condition 2 has lower H2O:CFeed ratios as compared to Test Condition 1. The resulting H2 mole fraction of ~45% is similar to the previous Test Condition 1 results. However, the CO mole fraction of ~28% in Test

Condition 2 is significantly higher than the 22% measured in Test Condition 1.

(a)

(b)

Figure 33. (a) Molar ratio between steam and carbon in wood pellet and (b) N2 free syngas composition and dry syngas purity under Test Condition 2

Table 6 summarizes the typical syngas composition from the two test conditions and the corresponding simulation results from an ASPEN RGibbs model. Two time periods are chosen as indicated by the shaded intervals A in Figure 32(b) and B in Figure 33(b).

Interval A B [O]:CFeed 2.0 1.7 H2O:CFeed 1.57 1.41 Mole Fraction Sim. Sim. Sim. Sim. Exp. Exp. (dry base) XC=100% XC=75% XC=100% XC=90% H2 (%) 45 47.9 46.3 45 49.2 49.1 CO (%) 20 26.7 20.6 25 28.7 26.4 CO2 (%) 32 25.4 33.1 25 22.1 24.6 Syngas Purity 65 74.6 66.9 70 77.9 75.4 (%) H2:CO 2.3 1.8 2.2 1.8 1.7 1.9 Table 6. Comparison between bench scale reducer experiment results and ASPEN

RGibbs reducer model results for biomass gasification

In Table 6, two columns of simulation results are given for each experiment condition. The first simulation column, labeled as “XC=100%”, corresponds to a full conversion of biomass. In this case, all biomass components are sent into the RGibbs module along with the ITCMO oxygen carrier to generate syngas under 1000°C and 1 atm. The simulation results with XC=100% in both intervals deviate significantly from the experiment results. The syngas produced from the reactor has a lower CO mole fraction and higher CO2 mole fraction compared to the corresponding simulation result.

Therefore, the H2:CO ratio measured from the experiment is higher than theoretically expected results.

The discrepancy can be explained by the incomplete conversion of biomass char due to the large size of the wood pellets. Because the gasification of char produces CO and H2 while consuming CO2 and H2O, incomplete conversion of char lowers the amount of CO and H2 that can be produced. As shown in Table 6, when the ASPEN RGibbs model is

adjusted to allow a portion of the biomass char (carbon) to leave the reactor unconverted, the simulation results agree with the experiment data.

All the simulations shown in Table 6 produce FeTiO3 as the only form of reduced oxygen carrier, which is confirmed by oxygen carrier conversion measured via post experiment sample thermal oxidation and by the crystal structure observed via XRD analysis.

After the steady state operations are achieved, as verified by a steady gas composition reading from the gas analyzers, the screw feeder motor and reactor heating elements are powered off. A N2 flush is maintained in the stationary bed to flush out the residual syngas in the reactor while preventing the oxygen carriers from being oxidized by air, which thereby preserves the oxygen carrier conversion profile along the bed during the cool down process. Samples of oxygen carriers are taken from nine locations along the height of the moving bed reducer for post experiment solid analysis. The solid samples are re-oxidized by air at 900°C. The oxygen carrier conversion is calculated by

100% (3-6) where and are the mass of the sample after and before re-oxidation, and is the weight fraction of available oxygen in the oxygen carrier particles.

X-ray diffraction (XRD) analysis is also performed using a Rigaku SmartLab X-Ray

Diffractometer (XRD) with a monochrometer to eliminate fluorescence generated by iron. Scans are performed from 15° to 80° at a rate of 0.1°/min with an accelerating voltage and filament current of 40 kV and 44 mA, respectively.

The steady-state oxygen carrier conversion profile along the reactor at the end of

Test Condition 1 is shown in Figure 34(a). At the top section of the reducer, the oxygen carrier conversion is less than 11% due to the short contact time of oxygen carrier with the reducing gases. As shown by the XRD pattern in Figure 34(b), sample 1 contains both

FeTiO3 and Fe2O3. The oxygen carrier conversion increases sharply to ~30% in less than

30 cm from the top of the reactor, and gradually to 33% as the solids approach the reducer gas outlet. The rapid conversion of the oxygen carrier in the initial 30 cm is an indication of a fast reaction kinetics between the gaseous volatiles/tars and the oxygen carriers. In contrast, the gasification reaction of char is the rate limiting step in the conversion of biomass. Note that a 33% conversion of oxygen in the oxygen carrier corresponds to the reduction to FeTiO3. XRD pattern in Figure 34(b) confirmed that the oxygen carriers at the bottom of the reactor, labeled as sample 2, consists of only the

FeTiO3 phase, which is consistent with the prediction of the ASPEN RGibbs model.

(a)

35% 30%

25% ② 20%

15% ① 10%

Oxygen Carrier Conversion 5% 0% 050100 Distance from Top (cm)

(b)

Figure 34. (a) Oxygen carrier conversions and (b) XRD spectrums of oxygen carrier samples from different locations

Based on the experiment results, the reactions occurring at different locations in the co-current moving bed reducer when biomass is used as the fuel can be inferred as shown

in Figure 35. At the upper portion of the reducer, the biomass feedstock is heated by the

ITCMO oxygen carriers and undergoes rapid pyrolysis, producing gaseous volatiles/tars, syngas, and char. The volatiles are rapidly cracked and oxidized by the ITCMO oxygen carriers. Because the majority of biomass is volatile matter, most of the gas-solid reactions occurring between the oxygen carriers and the fuel will take place in this region of the reducer, completing most of the oxygen carrier conversion. The char gasification with H2O and CO2 follows at a slower pace along the remainder of the reducer. The products of these reactions have little contribution to the reduction of the oxygen carrier.

Thus, no substantial change in the solid conversion profile is observed in this region of the reactor.

Figure 35. Reactions occurring in different locations of the reducer for biomass gasification

In a large scale chemical looping biomass gasification plant, full conversion of char can be achieved by the following means:

increasing the reactor volume to provide sufficient residence time for the

biomass char gasification reactions;

reducing the particle size of biomass to increase the surface area of the char

and, thereby, increase its gasification rate;

operating the reducer at an elevated pressure to increase reaction kinetics.

Alternatively, the process can be intentionally designed to leave a portion of the char unconverted in the reducer. The unconverted char will enter the combustor with the reduced ITCMO oxygen carriers and be burned in air. When H2-rich syngas is needed for downstream catalytic processing, bypassing a portion of the unconverted char to the combustor will increase the H2:CO ratio of the syngas produced from the reducer, as illustrated in Table 6. As a result, the reducer steam consumption can be decreased.

Burning unconverted char in the combustor releases thermal energy, which can be used for the heat requirement for the endothermic reactions in the reducer and/or auxiliary power generation.

Scale-up of Chemical Looping Partial Oxidation

The experiment studies in the previous sections have demonstrated the feasibility of producing high purity syngas from gaseous and solid carbonaceous fuels using the chemical looping partial oxidation process with co-current moving bed reducer. In order to further scale up the process, the co-current moving bed reducer has to be integrated with the fluidized bed combustor. In order to design and operate an integrated chemical looping system, the following issues must be properly addressed:

1. The system must be capable of circulating the oxygen carriers between the

moving bed and the fluidized bed reactor, so that the oxygen carriers can

continuously transfer oxygen to the fuels.

2. The gases in the moving bed reactor and the fluidized bed reactor must be

segregated in order to prevent the product gas from leaking and/or being

contaminated/diluted.

The most common way to carry out cyclic reactions on solid particles involves the use of circulating fluidized beds (CFBs). In a typical CFB system, a pneumatic riser is located on the top of the fluidized bed reactor. The solid particles can be conveyed upwards through the riser to the top of the system by a gas flow with suitable gas velocity. The solid particles are then separated from the gas flow by a gas-solid separation device, such as a cyclone, and flow downwards in a stand pipe before being conveyed to the fluidized bed reactor again.

A 15 kWth sub-pilot scale CFB system for the chemical looping partial oxidation process was designed, constructed, and tested. This system integrated a co-current moving bed reducer and a fluidized bed combustor, which are connected using non- mechanical gas sealing devices, a lean phase riser, and a gas-solid separator. During operation, the fuel is introduced into the top of the co-current moving bed reducer and converted to syngas by the ITCMO oxygen carriers. The product syngas is separated from the oxygen carriers at the bottom of the reducer. The reduced oxygen carriers from the reducer are regenerated by air in the combustor and transported back to the top of the reducer via the riser. Solid circulation and gas sealing in the CFB system are achieved without mechanical devices inside the reactor, which is suitable for further scale-up.

Sizing of the reactors and interconnecting gas sealing devices are based on hydrodynamic calculations. The range of operating conditions, including temperature,

fuel capacity, and residence times of gas and solid in each of the reactors, must first be identified for purpose of system design calculations. Based on the operating conditions, the expected syngas composition can be estimated by a performance model of the system, and the gas flow rate in each of the reactors can be determined.

Sizing of Reactors

The dimensions of the cocurrent moving bed reducer are restricted by the following criteria:

The volume of the reducer should be large enough to provide sufficient

residence time for both solid and gaseous species and constrained by

Equation (3-7) and (3-8).

⋅ 1 ⋅ 3‐7

⋅ ⋅ 3‐8

where is the volume of the reducer; is the voidage of the moving bed;

is the volumetric flow rate of oxygen carrier particles; is the required

residence time for oxygen carrier particles; is the volumetric flow rate of

gases; and is the required residence time for gases.

The cross-sectional area of the reducer should be large enough to avoid a

high gas velocity, and the corresponding high pressure drop, in the reducer.

The quantitative criterion is chosen to have a gas velocity lower than 80% of

the minimum fluidization velocity of the particles as given in Equation (3-

9).

0.8⋅ 3‐9

where is the cross-sectional area of the reducer, and is the minimum

fluidization velocity of the oxygen carrier particles.

For the purpose of an externally heated reactor, the diameter of the reactor

should be as small as possible to avoid a significant temperature gradient in

the radial direction.

The sizing criteria for the combustor and riser are different from that of the reducer.

As mentioned previously, the combustor reactor is a fluidized bed reactor where oxygen carrier particles are regenerated via the reaction with oxygen from air. The purpose of the riser is to transport the re-oxidized oxygen carriers back to the reducer. Air introduced into the combustor serves three purposes:

providing oxygen for oxygen carrier regeneration;

fluidizing the oxygen carrier particles in the combustor;

entraining the oxygen carrier particles in the riser back to the reducer.

Therefore, the sizing of the combustor and riser is restricted by the following consideration:

At the designed air flow rate, oxygen flow must be greater than the amount

required by oxygen carrier regeneration. The required amount of oxygen can

be determined based on the fuel capacity and expected syngas composition.

At the designed air flow rate, the gas velocity in the combustor should fall

between the minimum fluidization velocity and the terminal velocity of the

oxygen carrier particles, as given in Equation (4.5.8)

,⁄ 3‐10

where is the minimum fluidization velocity of oxygen carrier particles;

, is the volumetric flow rate of gas in the combustor; is the cross-

sectional area of the combustor; and is the terminal velocity of the

oxygen carrier particles.

At the designed air flow rate, the gas velocity in the riser should be greater

than the terminal velocity, as given in Equation (3-11).

, 3‐11

The volume of the combustor should be large enough to provide sufficient

residence time for the oxygen carrier particles to be completely re-oxidized:

⋅1 ⋅ 3‐12

where is the volume of the combustor; is the voidage of the fluidized

bed at the designed air flow rate; and is the required residence time for

the oxygen carrier particles.

Design of the Sub-pilot Reactor System

Based on the sizing calculation in the previous section, the reducer of the sub-pilot scale chemical looping reactor is designed as a cylindrical column with diameter of 6 inches and height of 78 inches. The combustor reactor is designed to be a cylindrical column with diameter of 6 inches and height of 20 inches. The riser diameter is determined to be 2 inches. The total height of the 15 kWth sub-pilot unit is over 200 inches.

Syngas produced in the reducer must be separated from the oxygen carrier particles.

Thus, a gas-solid separator is designed at the bottom of the reducer. The reducer separator is a cylindrical column with a large diameter compared to the reducer. The separator is sized such that the gas velocity in the annular region between the separator and the reducer is lower than the minimum fluidization velocity of the oxygen carrier particles.

Thereby, syngas can be extracted from the gas outlet on top of the separator without entraining any oxygen carrier particles. The diameter of the reducer separator is determined to be 10 inches.

Figure 36 shows the sub-pilot scale reactor installed on a supporting structure. The

Incoloy reactor is supported at two locations. The bottom of the combustor is standing on the ground floor, while the reducer separator is supported from the second floor of the structure. Figure 37 shows the refractory seat that supports the reactor at the reducer separator. The supporting seat is designed to form a cone-shaped slot for the reducer separator. The seat is made from two refractory pieces to minimize the heat loss from this section. The refractory seat pieces are embraced by two iron holders, which are bolted to the structure. Additional horizontal support is installed around the top section of the sub- pilot unit to provide extra lateral stability.

Figure 36. 3D drawing (left) and picture (right) of the sub-pilot reactor on structure

Figure 37. Supporting seat for the reducer separator

In order to operate the sub-pilot scale reactor system, a number of process inlets and outlets are needed to introduce gases and solids into the reactor, and allow products to leave the reactor. The process inlets and outlets are shown in Table 7 and Figure 38.

Besides, temperature and pressure sensors are required to monitor the status of the reactor system; other instruments, including valves, mass flow controllers (MFCs), and heaters, are required to operate and control the reactor system. A list of major instruments is shown in Table 8.

Process Inlet/Outlet Location Reducer gas inlet Top of reducer Reducer steam inlet Top of reducer Solid fuel inlet Top of reducer Reducer gas outlet Bottom of reducer Combustor gas inlet Bottom of combustor Riser gas inlet Bottom of riser Separation gas inlet Gas-solid separator Combustor gas outlet Top of gas-solid separator Upper zone seal inlet Upper zone seal Lower zone seal inlet Lower zone seal L-valve inlet L-valve Table 7. Process inlets and outlets on the sub-pilot reactor

Figure 38 Process inlets and outlets on the sub-pilot reactor

Instrument Description Location Tag PIT-204 System pressure transducer Combustor gas outlet PDIT-202 Combustor differential pressure transducer Combustor PDIT-203 Riser differential pressure transducer Riser PDIT-350 Upper zone seal differential pressure transducer Upper zone seal PDIT-351 Upper zone seal differential pressure transducer Upper zone seal PDIT-370 Lower zone seal differential pressure transducer Lower zone seal PDIT-371 Lower zone seal differential pressure transducer Lower zone seal PDIT-372 L-valve differential pressure transducer L-valve PDIT-550 Reducer differential pressure transducer Reducer TE-201 Combustor temperature transducer Combustor TE-202 Riser temperature transducer Riser TE-203 Gas-solid separator temperature transducer Gas-solid separator TE-501 ~ Reducer temperature transducers Oxidizer TE-508 TE-600 ~ Pot temperature transducer Pot TE-605 FC-201 Combustor N2 MFC Combustor gas inlet FC-202 Combustor air MFC Combustor gas inlet FC-203 Riser air MFC Riser gas inlet FC-204 Separation air MFC Separation gas inlet FC-320 L-valve MFC L-valve FC-350 Upper zone seal MFC Upper zone seal FC-370 Lower zone seal MFC Lower zone seal FC-501 Solid feeder carrier gas MFC Solid feeder COF500 FC-521 ~ Reducer gas MFCs Reducer gas inlet FC-524 PCV-590 Reducer pressure control valve Reducer gas outlet PCV-700 Combustor pressure control valve Combustor gas outlet PUM-510 Reducer water pump Reducer steam inlet COF-500 Solid feeder Solid inlet CLR-580 Reducer gas cooler Reducer gas outlet CLR-790 Combustor gas cooler Combustor gas outlet Table 8. Major instrumentation on the sub-pilot scale unit

The structure of the sub-pilot unit distributed control system (DSC) is shown in

Figure 39. The sub-pilot unit is controlled by two Allen-Bradley programmable logic controllers (PLCs). The main PLC is connected with 15 input/output (I/O) cards, 97

including digital input, digital output, analog input, analog output, and thermal couple input cards. It processes the input signals from thermal couples, pressure transmitters and gas analyzers, and controls the power output of the electric heaters and the position of the pressure control valve. The MFC PLC handles the serial port (COM port) communication with the mass flow controllers. The two PLCs and the human-machine interface (HMI) computer are connected over an Ethernet.

Figure 39. Control system of the sub-pilot scale unit

Operation of the Sub-pilot Reactor System

The sub-pilot scale chemical looping partial oxidation unit is operated using the

ITCMO oxygen carriers and CH4 as the fuel. The reactor is externally heated to 1000°C while maintaining solid circulation. After the temperature setpoint is reached, CH4 feeding is started. Figure 40 shows the N2-free gas composition at the reducer outlet.

Figure 40. N2-free product gas composition at the reducer outlet of the sub-pilot unit

Similar to the previous experiments on co-current moving bed reducers, the gas produced from the reducer right after the start of fuel injection is predominantly CO2. as the oxygen carriers in the reducer are reduced to lower oxidation states, the mole

fractions of CO and H2 increase gradually. At steady state operation, a CH4 conversion of over 99% and a H2:CO ratio ranging from 1.6 to 2.2 are achieved, as shown in Figure 41.

Figure 41. Syngas generation performance of the sub-pilot unit

Note that during the time period indicated by the dashed block in Figure 40, the H2 mole fraction is increasing while the CO mole fraction is decreasing. The H2:CO ratio is significantly higher than the stoichiometric ratio between H and C in the feedstock. This abnormal gas composition is caused by an insufficient oxygen carrier circulation rate.

The ITCMO oxygen carriers are over-reduced in the reducer, resulting in carbon deposition on the oxygen carriers. When the oxygen carrier circulation rate is increased, the H2 and CO mole fractions changed back to the normal level rapidly.

100

Conclusion

The concept of using the ITCMO oxygen carrier and the co-current moving bed reducer is tested experimentally to confirm the feasibility of the chemical looping partial oxidation process.

Reactivity test on the ITCMO oxygen carrier confirms that the reactivity is sustained over multiple reduction-oxidation cycles. Fixed bed experiment confirms that the ITCMO oxygen carrier can effectively crack and convert volatiles and tars generated from the pyrolysis of solid fuels.

The conversion of CH4 and biomass to syngas is tested in a fixed bed reactor and a bench scale co-current moving bed reducer. High purity syngas with a H2:CO ratio of 2 is produced from both CH4 and biomass. The syngas composition and solid composition agree with the simulation results obtained from the ASPEN RGibbs module.

The chemical looping partial oxidation process is scaled up to a 15 kWth sub-pilot scale unit. The unit is fully integrated with a co-current moving bed reducer and a fluidized bed combustor, connected by non-mechanical gas sealing and solid circulation devices. The control instruments and DSC are designed similarly to a large scale system.

Syngas generation in the sub-pilot unit using CH4 is demonstrated. Steady syngas generation operation is achieved with a CH4 conversion of over 99% and a syngas H2:CO ratio ranging from 1.6 to 2.2.

101

Chapter 4: Dynamic Simulation of Chemical Looping Systems

In an effort to scale-up the chemical looping processes developed at The Ohio State

University, a 250 kWth coal-direct chemical looping (CDCL) pilot unit and a pressurized

250 kWth syngas chemical looping (SCL) pilot unit are constructed and tested. The chemical looping partial oxidation process for syngas generation is also scaled up to a 15 kWth sub-pilot unit. However, due to the lack of automation mechanism, the operation of these units relies largely on the knowledge and experience of the operators. The last portion of the dissertation will focus on the effort for developing a control scheme for the autonomous startup, shutdown, and ramping of the chemical looping processes.

In order to develop the controllers for the chemical looping system, a dynamic model that simulates the transient behavior of the system under control inputs is necessary. This chapter will describe a dynamic model that focus on the hydrodynamic aspects of the chemical looping system with moving bed reactors. The next chapter will discuss the development of a hierarchical control scheme for autonomous operation of the system.

102

Physical Model

Components of the System

The chemical looping system considered here is a CFB system, which consists of a fluidized bed reactor, a lean phase riser, and a packed/moving bed stand pipe. The solid particles in the fluidized bed reactor will be entrained by gas, enter the lean phase riser above, and be transported upwards. After gas-solid separation, the particles will fall into the stand pipe and move downwards in a packed manner. At the bottom of the stand pipe, the particles are transported into the fluidized bed reactor again by a non-mechanical valve, such as an L-valve.

One or two section(s) of the stand pipe is (are) enlarged to serve as reactors. For example, the pressurized 250 kWth SCL pilot unit has a counter-current moving bed reducer followed by a counter-current moving bed oxidizer. Therefore, two or three reactors occur in the system, including one fluidized bed reactor and one or two moving bed reactor(s). In order to segregate the gaseous species in different reactors, each two of the reactors are interconnected by a thinner stand pipe section, known as a zone seal. The zone seals serve two purposes:

A gas stream (sealing gas) is introduced to the middle of the zone seal, and

flow both upwards and downwards into two different reactors. Therefore,

103

the gas from the reactors cannot flow through the zone seal. The reactors

are, hence, segregated.

The gas flow in the zone seal can cause a pressure difference on the two

ends of the zone seal. This pressure drop can be used to compensate the

pressure drop created by the gas flows in the reactors, and hence maintain

the pressure balance of the system.

Besides sealing gases used in zone seals, gases are introduced into the reactors to carry out the chemical reactions. Each of the reactors has a gas inlet and a gas outlet. For the moving bed reactors, the gas can be introduced either from the bottom or from the top of the reactor, known as the counter-current or the co-current operation mode, respectively. For the fluidized bed reactor, gas is introduced from the bottom of the fluidized bed. The gas maintains particle fluidization in the fluidized bed reactor as well as particle entrainment in the riser. At the gas outlets of each of the reactors, back pressure control valves (PCVs) are used to regulate the pressure balance of the system.

The PCV at the fluidized bed reactor outlet is used to control the overall pressure of the system, while the PCV at a moving bed reactor outlet is used to control the pressure drop across a zone seal.

Mass Balance

In order to describe the variation of pressure in the chemical looping system, the variation of the quantity of gases and solid particles in the reactors must be determined.

Due to the circulation of particles and the flow of gases, the amount of masses in all parts

104

of the system may vary over time. It is natural to require that mass balance is maintained, i.e. for a given reactor, the difference between the mass entering it and the mass leaving it equals the mass accumulating in the reactor.

Consider the mass balance of gases in the upper moving bed reactor. The mass entering the reactor include: (1) gas introduced from the gas inlet; (2) gas flowing upwards from the middle zone seal; (3) gas flowing downwards from the upper zone seal;

(4) gas generated in the chemical reactions. The mass leaving the reactor include: (1) gas flowing outwards from the gas outlet; (2) gas consumed in the gas-solid reaction. The difference between them should equal to the mass accumulation in the reactor.

The considerations of mass balances in other reactors are similar.

Pressure Balance and Hydrodynamics

The accumulation of gas in a reactor will increase the pressure in it and affect gas flow rates in different parts the system. As the gas flow through the bed of particles, a pressure difference between the two ends of the bed (pressure drop) is created as a result of the friction between the gas and the particles. On the other hand, the pressure difference between two points in the system can determine the direction of gas flow between these two points. In general, increasing the gas flow rate at a gas inlet will elevate the pressure at the gas injection point; increasing the valve opening of a PCV will increase the gas flow rate flowing out from the corresponding gas outlet, while decreasing the pressure upstream the PCV.

105

Consider closed loop path of by which the particles travel through the system. The pressure drops on all components of the system along the path should add up to zero, since the system is a closed loop. Therefore, the change of pressure at any point will have a global impact on almost all other points in the system.

In the moving bed reactor(s) and zone seals, the gas velocity is much higher than the solid velocity. Therefore, the hydrodynamic characteristics of a moving bed reactor or a zone seal is considered to be the same as that in a packed bed reactor.

The particle flow rate and pressure drop in the L-valve is determined by the gas flow rate in it.[59] When the gas flow rate is low, the particles will not flow, and the L-valve should behave similar to a packed bed. As gas flow rate increases, the pressure drop will also increase until a critical flow rate where particles start to flow through the L-valve.

The pressure drop is not significantly dependent on the gas flow rate after the critical flow rate. It is known that, as gas/solid temperature increases, this critical flow rate decreases. The flow rate of solid also increases under a given gas flow rate. However, actual correlation among gas flow rate, solid flow rate, critical flow rate, pressure drop, and temperature is not available.

The fluidized bed reactor is considered as a bubbling fluidized bed. At low gas flow rate, the bed of particles will not be fluidized and act as a fixed bed. The pressure drop across the bed increases as the gas velocity increases. As gas velocity increases to above the minimum fluidization velocity, the bed will be fluidized. The pressure drop will stop increasing and start fluctuating significantly. The height of the bed expands as gas velocity further increases, until the top of the bed enters the riser. After this point, if the

106

gas velocity in the riser is sufficient to entrain the particles, the particles will be transported to the top of the riser and returned to the moving bed side. Otherwise, the particles will accumulate in the riser and create a very high pressure drop, as a result of the smaller diameter of the riser compared to the fluidized bed reactor. Besides gas velocity change, the bed height also increases as more particles flow into the vessel.

The transient behavior of particle flow in the riser is not considered here.

Assumptions for Simulation

In order to simplify the simulation, the following assumptions are made:

The gases in the system are ideal gases.

The gas flow rate through all sections of the system can change

instantaneously. Therefore, the pressure drop across the reactors and zone

seals can change instantaneously.

The pressure at zone seal injection points (due to the much smaller volume

of zone seals compared to the reactors, the transient time should be

negligible) can change instantaneously. In contrast, the pressure in the

reactors can only change gradually due to the difference between the gases

entering and leaving the reactor.

The status of the system is described by four integrals, i.e. the mean

pressure in three reactors ( , , and ) and the solid holdup in the

fluidized bed reactor.

107

The pressure along a moving/packed bed reactor is linearly distributed.

Therefore, the mean pressure in a reactor is the numeric average of the

pressures at the two ends of it.

The relationship between the gas flow in a section of moving/packed bed

and the pressure drop across this section follows the Ergun Equation.

The particles in the system are spheres with uniform size and no attrition.

The particle level in the moving bed side is constant.

The particles in the fluidized bed reactor will not enter (and hence not be

entrained by) the riser unless the bed level reaches the top of the reactor.

When entrainment occurs, the particle holdup rate in the riser is a constant,

while the particle velocity is the difference between the gas velocity and the

terminal velocity of the particles.

Mathematical Description

The components and physical quantities in the model of chemical looping system is shown in Figure 42. The components of the system are in blue; the dimensions of the components are in black; the variables that can be manipulated or controlled are labeled in red; the variables that can be measured are labeled in green; the variables that cannot be measured but are estimated by calculation are labeled in magenta. The quantitative relations among the variables are described in this section.

108

Figure 42. Components and Physical Quantities in the Model of Chemical Looping

System 109

Mass Balance in the Reactors

In order for the mass balance to hold in the chemical looping system, the difference between the mass entering a reactor and the mass leaving it should equal the mass accumulated in the reactor:

4‐1

4‐2

4‐3 where , , and are the amount of gases in the upper moving bed reactor, the lower moving bed reactor, and the fluidized bed reactor, respectively. For the purpose of simulation, we define the mean pressures in the reactors as:

4‐4

4‐5

4‐6 where , , and are the void volumes of the fluidized bed reactor, the lower moving bed reactor, and the upper moving bed reactor, respectively. Note that these void volumes of the reactors include the void space in the reactors as well as the volume of the downstream coolers. The variation of the pressures in the reactors can be calculated by:

4‐7

4‐8

110

4‐9

Pressure Drop and Pressure Distribution along a Moving/Packed Bed

. || 4‐10 where Δ is the pressure drop, is the (dynamic) viscosity of the gas, is the bed height,

is the bed voidage, is the particle size, is the gas velocity, and is the density of the gas. The viscosity of the gas depends mostly on the temperature of the gas as calculated from Sutherland’s formula:

. . 4‐11 where is a reference temperature at which the gas viscosity is , and is the

Sutherland’s constant for the gas. For nitrogen, 1.41 10 ⋅ ⋅ . and

111 ; for air, 1.51 10 ⋅ ⋅ . and 120 . Since the Sutherland’s formula is only valid up to 500K, further validation of the viscosity correlation at high temperature may be required.

Equation (4-10) only applies to the steady state condition. Also, it will not apply to conditions with large pressure drop. The dynamics of pressure distribution along a packed bed can be deduced as follows. 111

Assume the gas flow is isothermal, and the Ergun equation applies to every small segment of packed bed.

. || || 4‐12

. where and are constants.

Consider a segment of a packed bed. Mass balance gives:

Δ Δ

or ΔΔ 4‐13 where is the mole number of gas, is the cross section of the packed bed, is the gas constant, is gas temperature, is the height of the reactor, and are the pressure and gas velocity at height , and is time. By ideal gas law,

ΔϵΔh 4‐14

Combining (4-13) and (4-14) gives:

4‐15

Assume in time period Δ, the gas block considered above moved by a distance of

Δ Δ. By Newton’s second law:

,, Δ ΔΔ Δ

4‐16

112

where is the total derivative of gas velocity, and Δ is the frictional force exerted by the particles to the gas. The frictional force should not depend on whether the system is in steady state or not. Therefore, let

, , 4‐17

To determine , consider the steady state condition, where 0.(4-12), (4-15), and (4-16) become:

|| 0

Eliminating gives:

ϵ || 4‐18

Substituting the left hand side of equation (4-16) by equation (4-18) gives:

|| 4‐19

|| 4‐20

The evolution of gas velocity and pressure distribution can be determined by solving equation (4-15) and (4-20):

|| 4‐21

Numeric solution of equations (4-21) shows that the pressure distribution in a packed bed under the normal operation condition of the chemical looping system we are considering is very close to a linear distribution, which validates assumption (5): “The

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pressure along a moving/packed bed reactor is linearly distributed. Therefore, the mean pressure in a reactor is the numeric average of the pressures at the two ends of it.” Thus, the pressures at the two ends of a reactor are related to the mean pressure by:

(4-22)

. || 4‐23

(4-24)

. || (4-25) where the gas densities and should be calculated from the mean pressures. and are the mean gas velocity in the moving bed reactor. Because the gas flow rate entering the moving bed reactor may or may not be equal to that leaving reactor, the mean gas velocities are estimated to be the average value of the gas entering and leaving the moving bed reactor:

(4-26)

(4-27)

Note that and are subtracted from the outbound gas flows because these gas flows do not flow through the moving bed reactor.

Gas Flow Rates through Valves

The estimation of gas flow rates through the three pressure control valves PCV-490,

PCV-590, and PCT-700 are based on the standardized method for compressible fluid in

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the valve sizing instruction document released by Emerson, a major control valve manufacturer.[60]

With all the pipeline friction omitted, the flow rate of gas flowing through a control valve is:

(4-28) where is the valve opening, is the flow coefficient of the valve, 2.63

10 is a numeric constant1, is the pressure upstream the valve

(Pa), is the molecular weight of the gas (Dalton), and is the gas temperature (K). is the expansion factor:

is the ratio of pressure drop to upstream absolute pressure, and can never be greater than the value for the critical flow condition, :

min ,

The downstream pressure of the pressure control valves is assumed to be 101325

Pa. Exemplary valves involved in the chemical looping system are listed in Table 9.

1 The formula at the upper right corner of page 642 of the Emerson document has a typo. The constant used in this 3 formula should be N9 instead of N7. The constant N9=21.2 in the document is used for volumetric flow (Nm /hr) calculation. Here the constant is converted to give the molar flow (mol/s). 115

PCV-490 Valve PCV-700 PCV-590 Body Style Globe Globe Plug Style Micro-Form Cage Guided Flow Characteristics Equal Percentage Equal Percentage Valve Size 1" 2" Port Size 3/4" 2" 8.84 53.8 0.92 0.7 Table 9. Pressure control valves in the chemical looping system

is a function reflecting the effect of the valve opening to the gas flow rate, satisfying 0 ≅0 and 1 1. The actual form of this function depends on the flow characteristics of the valve. For linear valves:

(4-29)

For an ideal equal percentage valve:

(4-30) where is the valve rangeability. However, the flow characteristic of an actual equal percentage valve deviates from the ideal curve at high valve opening as shown in Figure

43. The data is adapted from the valve characterization curve provided by a valve vendor.

The curve can be fitted with the following function:

sin sin sin (4-31) where 7.024, 0.1519, 0.02382,

0.1559, 6.87, 1.666,

0.02012, 15.23, 0.6543,

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Figure 43. Flow characteristics of an actual equal percentage globe valve

Pressure Drop in the Fluidized Bed Reactor and The Riser

The operation status of the fluidized bed reactor and the riser highly depend on the gas velocity. Two critical velocities for the operation status are minimum fluidization velocity in the fluidized bed reactor, , and entrainment velocity (gas velocity above which particles will be entrained upwards by the gas flow) in the riser, .

When gas velocity in the fluidized bed reactor is lower than the minimum fluidization velocity, the particles in the fluidized bed cannot be fluidized, and the bed is, in fact, a packed bed. The pressure drop across the bed can be determined by the Ergun

Equation (4-10). As gas velocity increases, the pressure drop increases, until the pressure drop equals the gravity of the bed. The gas velocity at this point is the minimum fluidization velocity :

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. (4-32)

. or 1 (4-33) where is the weight of all particles in the reactor, is gravity constant, is the cross section area of the reactor, is the height of the bed, and is the density of the particles.

After the gas velocity exceeds , the particle bed will be fluidized, and the pressure drop will start to fluctuate significantly. If gas velocity further increases, the pressure drop (mean value) will not further increase. However, the bed height will expand as the bed voidage increases. The bed expansion can be estimated using equation (4-34).[61]

. . . 1 1 1 14.311 . . (4-34) where is the voidage in the fluidized bed, is the bed voidage at minimum fluidization velocity, is the gas velocity in the fluidized bed, and is the density of gas. SI units should be used in equation (4-34). However, equation (4-34) over estimates the bed expansion at high temperature.[62] Instead, it was proposed that the gas viscosity and density at 20°C should be used in the calculation as given in (4-35).[62]

∗ . . . 11 1 21.365 . . (4-35) ∗

∗ where is the minimum fluidization velocity at 20°C and operating pressure,

29 is the molecular weight of air, is the pressure of the system, and is ambient pressure. Although equation (4-35) gives better consistency with experimental data, the predicted bed expansion doesn’t vanish at minimum fluidization velocity (when

∗ ) since . Further verification may be needed for a better model. 118

The entrainment velocity in the riser is equal to the terminal velocity of particles in the gas, which can be determined by:

(4-36)

or (4-37)

24 2 . where 18.5 2 500 (4-38) 0.44 500 210

is the drag coefficient. is the particle Reynold’s Number at terminal velocity. Solving for gives:

2 . 0.154 . . 2 500 (4-39) . 1.74 500 210

At first glance, iteration may be required when we are using equation (4-39) to calculate the terminal velocity, as and depend on each other. There are a few ways to avoid iterative calculation.

The first way is to consider the dimensionless number , which is independent of or . If 2, the first formula in (4-39) should be used. By definition of :

2 18 18 18

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On the other hand, if 500, the third formula in (4-39) should be used:

. 500 ⋅1.74 1.74√

8.3 10

Therefore, we can calculate based on operation condition first, then determine which formula in (4-39) should be used.

The other method relies on the relation between and . As shown in Figure 44, the drag coefficient at all values is the highest one calculated from the three formulae in equation (4-38). Thus, will be the minimum of the three values calculated from the three formulae in equation (4-39). This method will fail if is greater than

200,000, because will decrease to 0.1 at this condition, which is no longer the largest value predicted by the formulae in (4-38). We can estimate the rough range of expected

. Over the range of system startup and normal operation, the typical values of the

parameters in equation (4-38) are: ≅1~10⁄ , ≅10⁄ , ≅1.5

10 , ≅ 1~5 10 ⋅ . Thus, ranges from 300 to 15,000. Therefore, only the last two formulae in equation (4-39) will be used, and the calculation based on the maximum of (4-38) or minimum or (4-39) will not fail.

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Figure 44. Relation between and

Neglecting the pressure drop caused by the friction between gas flow and reactor walls, the pressure drop in a fluidized bed reactor or a riser is equal to the gravity of all the particles in the reactor/riser divided by its cross sectional area. In the chemical looping system, however, the bed height in the fluidized bed reactor must be considered as the riser is right above the fluidized bed reactor. Let be the mass of particles in the fluidized bed reactor, be the mass of particles in the riser, be the total particle holdup in the fluidized bed reactor and the riser, and be the cross sectional area of the fluidized bed reactor and riser, respectively. Because the transient behavior of particle flow in the riser is not considered here, the total particle holdup is needed to consider the mass balance in the fluidized bed reactor and the riser. Based on assumption

(9), if the particle hold up and bed expansion is so low that the bed height, , is lower

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than the height of the fluidized bed , then 0 and . The pressure drop in the fluidized bed reactor Δand that in the riser Δ are:

Δ , Δ 0 4‐40

As the bed height, , increases and reaches the height of the fluidized bed , only a portion of the total particle holdup will stay in the fluidized bed reactor, while the rest will enter the riser:

, (4-41)

Δ , Δ 4‐42

Under this condition, based on assumption (10), the particle flow rate leaving the fluidized bed reactor or the particle flow rate in the riser is:

4‐43 where is the gas velocity in the riser and 1% is the solid holdup rate in the riser.

Pressure Drop And Gas Flow in Zone Seals

We will use the upper zone seal as an example. As shown in Figure 42, the pressures at the two ends of the zone seal are and . Let be the pressure at the sealing gas injection point, be the total molar flow rate of sealing gas, and and be the molar flow rate of upward and downward gas flow, respectively. Note that the gas flow rates and are directional, i.e. the value is positive when the gas flow is in the direction of the arrows in Figure 42, while the value is negative when the gas flow is in the opposite direction. Based on assumption (2) and (3), the transient behavior of the zone seal is not considered. Mass balance gives: 122

(4-44)

Based on assumption (6), the pressure is related to the gas flow rates by the

Ergun equation:

. || (4-45)

. || (4-46)

, (4-47) where and are the lengths of the two sections of the upper zone seal above and below the injection point, is the cross sectional area of the upper zone seal, and is the gas temperature in the middle zone seal. Gas viscosity and density are all estimated at temperature of and pressure of . The gas velocity is assumed to be constant along the zone seal since the pressure drop is expected to be small compare to the absolute pressure.

Solving for the System Dynamics

The equations listed in the previous section need to be solved in order to obtain the dynamic behavior of the chemical looping system. These equations form a set of nonlinear differential-algebraic equations (DAEs), which includes a number of algebraic equations and four ordinary differential equations (ODEs), i.e. equations (4-7) through

(4-9) for the variation of reactor pressures, and an additional equation for the variation of

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particle mass in the combustor. The ODE for the variation of particle mass in the combustor depends on the solid circulation control mechanism in the system. Therefore, the status of the system can be represented by four variables, i.e. the pressure in the three reactors (, , and ), and the mass of particles in the combustor .

Due to the nature of the chemical looping process, the variables involved in the

DAEs form a “loop”. Thus, the DAEs cannot be solved sequentially. Note that the four

ODEs are dependent on the gas flow rates in different components of the system, including the reactors and the zone seals. To determine the time derivative of the four status variables, a set of nonlinear algebraic equations must be solved to determine the gas flow rates and densities in the system based on the inlet gas flow rates, outlet valve opening, and the current pressure of the reactors.

Iterative methods can be used to numerically solve the algebraic equations.

However, this method is inefficient, especially when chemical reactions, energy balance, and heat transfer in the reactors are added into the model.

An effective approximation for solving the DAEs is to assume that some of the gas density and pressures are not changed from the previous time step. This assumption is valid if the integration step size for the differential equations is sufficiently small. With this assumption, the algebraic equations for different components of the system can be solved sequentially to determine the gas flow rates. By avoiding the iterations in solving these equations, the speed for solving the DAEs is significantly improved.

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Conclusion

A dynamic simulation scheme for a chemical looping system using moving bed reducer and oxidizer is developed to study the transient behaviors of the system when it is subjected to status changes and disturbances. The mass balance and hydrodynamic characteristics of each components, including the moving bed reactors, zone seals, fluidized reactor, and pneumatic riser, are considered. The model consists of a set of

DAEs that depicts the variation of system status variables, such as the pressures in the reactors. The DAEs are solved to simulate the transient behavior of the chemical looping system. The dynamic simulation scheme can be used for control system development and operator training.

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Chapter 5: Hierarchical Control of Chemical Looping Systems

During the startup, shutdown, and capacity ramp-up/ramp-down of the chemical looping system, the temperatures, pressures, and gas flow rates may vary in a wide range.

In order to accommodate these condition variations and maintain steady solid circulation and gas sealing, sealing gases, aeration gases, and control valves must be properly manipulated. However, due to the highly interdependent nature of the process variables, manual manipulation of the chemical looping system may be challenging. The operation experiences gained from the sub-pilot to pilot scale chemical looping units show that the proper operation without autonomous control is highly dependent on the knowledge and experience of the operators. In the course of the development of the large-scale moving bed chemical looping system, the process control architecture applicable to a commercial scale chemical looping system is also explored. A hierarchical control algorithm consisting of a higher level controller (HLC) and sliding mode controllers (SMCs) is developed to facilitate the startup, shutdown, and capacity ramp-up/ramp-down operation of the OSU chemical looping system. The main goal of the autonomous startup sequence is to minimize operator interference and potential human error during the startup and normal operation of the chemical looping system, and to improve the steadiness and robustness of the operation.

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Reactor Pressure Control using HLC-SMCs

Compared to conventional and commonly-utilized controllers such as proportional- integral-derivative (PID), Smith Predictors, and Lead-Lag controllers, SMCs are more robust and versatile to account for modeling uncertainties and disturbances for the control of nonlinear systems.[63-66] By using SMCs, the process variables can slide on a predefined sliding surface towards the desired values. As a result, the status of the system can be regulated to follow a desired trajectory.

The application of the HLC-SMCs hybrid control system can be illustrated by considering a simple model of an isothermal reactor with only one inlet and one outlet, as shown in Figure 45. Gas is flowing into the reactor at flow rate , and flow out of the reactor through an ideal equal percentage control valve. We would like to control the pressure, , by manipulating the valve opening of the control valve.

Figure 45. One-tank reactor model

The dynamics of the system pressure is described by:

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(5-1)

is determined by valve flow equation (4-28) and (4-30):

, (5-2)

1 , min ,

Consider the pressurization of the system in the following steps (Figure 46):

1. Start at and pressurize to with control valve closed;

2. Pressurize from to at a given rate ;

3. Pressurize from to pressure setpoint at decreasing rate, and maintain

reactor pressure at .

Figure 46. Steps in the pressurization of the one-tank reactor

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To control the pressurization of the reactor, a HLC-SMC hybrid controller as shown in Figure 47 is used. The HLC is a finite state machine with a three different states , , and . The lower level controller is a sliding mode controller (SMC), whose sliding surface is decided by the HLC. The sliding surfaces involved are:

: 0 (5-3)

: 0 (5-4)

No sliding surface is used for because the valve is not manipulated in this step.

Figure 47. Controller for pressurization of the one-tank reactor

In order to regulate the rate of pressurization, the dynamics of must be obtained.

Taking the time derivative of (5-4) gives:

′ (5-5)

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Thus, to control the rate of pressurization, the rate of valve opening change, , must be

manipulated. Let be the manipulated variable. A simple control law is used in the

SMC:

⋅ , 0 (5-6)

is chosen to be positive because when 0, we need to open the control valve to depressurize the system, and vice versa.

The goal is to manipulate the system to reach the sliding surface 0.

Specifically, the value of in (5-6) must be chosen such that

0 , 0 0 , 0

or 0 (5-7) at least in a neighborhood of the sliding surface 0. In this case, the trajectories in the state space will converge to the sliding surface, or in other words, the sliding mode exists.

Equation (5-7) is called the reaching condition.

Consider the second step in pressurization, where , and the sliding

surface is 0. Taking the time derivative of gives:

′ (5-8)

If 0, or , we have . The reaching condition requires that

′ 0

or (5-9)

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At constant inlet gas flow, 0, and the right hand side of (5-9) is negative.

Thus, can take any positive value.

If 0, or , we have . The reaching condition requires that

′ 0

or (5-10)

At constant inlet gas flow, 0. Inequality (5-10) becomes:

⋅ ⋅ ⋅ ⋅ (5-11) ⋅ ⋅

If we let

⋅ ⋅ (5-12) then inequality (5-11) will hold at all times.

The HLC-SMCs is tested using a dynamic simulation code in MATLAB. The initial condition is set to: 0, 300, and 0. The inlet gas flow rate increases from 0 to 1000 / at 1⁄ /. The controller is set to ramp up the reactor pressure to 30 at a ramp rate of 1 /. At 45, the gas flow rate is set to suddenly increase from 1000 / to 1300 /.

Figure 48 shows the pressure variation in the simulation. The HLC-SMCs successfully regulates a stable pressurization of the reactor. After the sudden disturbance, the controller responses rapidly to depressurize the system back to the pressure setpoint.

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Figure 48. Simulation result for pressure regulation using HLC-SMCs

Figure 49 shows the trajectory of the system status variables on the phase plane. The arrows indicate the direction of system evolution. Arrows 1 and 2 represents the pressurization, arrow 3 shows the disturbance caused by the sudden increase of reactor inlet gas flow rate, and arrows 4 and 5 shows how the HLC-SMCs regulates the system pressure back to the setpoint. The trajectory of the system status clearly shows the two sliding surfaces that are designed for the system. The simulation confirms that the HLC-

SMCs is capable of manipulating the system to follow a predefined trajectory in the phase space of the variables. 132

Figure 49. System status trajectory in the phase plane during simulation

Automatic Operation of Chemical Looping Systems

A HLC-SMCs hybrid controller is designed to regulate the pressurization and depressurization of the pressurized SCL pilot plant, which involves 4 sliding surfaces corresponding to 4 different stages in the process. Specifically, the controller is designed to ramp up the system pressure at 2 psi/min in stage 1, and maintain a steady pressure of

30 psig in stage 2. In stage 3, the controller depressurizes the system at 1 psi/min before maintaining a pressure of 20 psig in stage 4. As shown in Figure 50(a), the reactor 133

pressure is well regulated and followed the predefined path through each state control.

The trajectory of the system in the phase plane closely followed the four predefined sliding modes as illustrated in Figure 50(b).

Figure 50. Performance of the HLC-SMC hybrid controller on the pilot plant for pressurizing and depressurizing the reactor

An autonomous startup sequence based on the HLC-SMCs hierarchical control architecture is developed and tested in the chemical looping sub-pilot unit. Specifically, the SMCs are designed to automatically establish and maintain desired pressure balance conditions on the moving bed reactor and standpipes to maintain proper solid circulation and gas sealing through process heat up and operation capacity variation thereby facilitating the startup operation of the chemical looping units without affecting the safety interlocks in the systems.

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Figure 51(a) is a schematic of the chemical looping sub-pilot unit. The system consists of a counter-current moving bed reducer, a fluidized bed combustor, and zones seal standpipes. As described in Chapter 4, the standpipes regulate the system pressure balance and serve as gas sealing devices to prevent the product/feed gases present in the two reactors from mixing with each other. As shown in Figure 51(b), the operator manipulates the gas flows (F1 and F3) and pressure control valve (PCV) opening in accordance to the temperature and pressure sensor readings to maintain system pressure balance and gas sealing. Note that the variables in the chemical looping reactor system are highly inter-dependent, i.e. the changes of one variable will significantly affect multiple other variables.

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Figure 51. (a) Schematic of the chemical looping sub-pilot unit; (b) Sensor inputs and manipulated variables for process control.

The operation of the sub-pilot chemical looping system can be divided into two states: the heating state and the fuel injection state. Each state requires a unique pressure balance and gas sealing condition. As shown in Figure 51(b), during the heating state, the differential pressure across the reducer, DP5, will increase as a result of the increasing temperature in the reducer. With the gas flow into the reducer, F2, held constant during heat up, the gas density decreases with increasing system temperature resulting in an increase in superficial gas velocity and, thus, an increase in differential pressure across 136

the reducer, DP3. Further, an increase in DP3 results in a higher local pressure point at the reducer gas inlet, P4, which decreases the bottom zone seal differential pressure, DP5.

In order to maintain proper pressure balance in the chemical looping reactor system during heat, a constant ratio between the top and bottom zone seals (DP2:DP5 ratio) must be maintained by manipulating the PCV at the outlet of the reducer.

In the fuel injection state, the gas flow rate into the reducer, F2, may vary widely between 0% and 100% load capacity. In this state of operation, proper gas sealing must be established and maintained by manipulating the sealing gas flows rates into the top and bottom standpipes, F1 and F3, in addition modulating the reducer outlet PCV to maintain the proper system pressure balance. Gas sealing in the upper zone seal requires the correct gas flow at all times, indicated by a positive differential pressure, i.e. DP1 >

DP2 > 0, while proper gas sealing in the bottom zone seal is indicated by DP4 > 0. Due to the interconnected nature of the variables, manual operation of the F1, F3, and PCV to maintain stable operation through transient conditions such as fuel capacity ramping is highly challenging and potentially hazardous if the operator is unable to maintain a proper system pressure balance resulting in a compromised gas seal between the highly combustible fuel in the reducer and oxidizing air in the combustor. In the case of pressurized chemical looping systems, additional complexity is introduced as the operator is also required to maintain a steady system pressure during all normal operation activities by manipulating the PCV at the outlet of the combustor.

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The performance of the controller during actual operation of the chemical looping sub-pilot unit during the heating state is shown in Figure 52. As illustrated in Figure

52(a), the reducer temperature and differential pressure continuously increase over the course of the heating. The heating state controller is able to accurately adjust the PCV opening according to the pressure sensor readings. Specifically, DP2 and DP5 are successfully maintained at a ratio of 0.6.

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Figure 52. Performance of the HLC-SMCs hybrid controller during heating

Once the reactor system reached a temperature of 550°C, the HLC-SCMs is switched to the fuel injection state, as shown by the dash line in Figure 53(b & c). The

HLC-SMC is able to successfully perform a state transition from heating to fuel injection and established the proper gas sealing in the upper zone seal, i.e. DP1 > DP2 > 0. After

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completing the transition, fuel process capacity ramp up and down are simulated in the sub-pilot chemical looping test unit by creating sudden changes in the reducer gas inlet flow rate, F2, as illustrated in Figure 53(a). For fuel processing capacity ramp up, F2 flow is increased from 35 slpm to 112 slpm in increments of 1, 2, 5, 10, and 30 slpm while for ramp down, F2 flow rate was decreased in increments of 35, 25, and 20. The increments are chosen to observe how the controller responded to the sudden change in system operation conditions. As shown in Figure 53(a), drastic changes in the bottom zone seal and reducer pressure drops, DP3 and DP5, are observed due to the ramping of the gas flow in the reducer. In response, the controller manipulated the sealing gas flows, F1 and

F2, to maintain gas sealing between each reactor as well as the reducer PCV opening as shown in Figure 53(b). During each increment of gas ramp up and down, the controller is able to maintain a steady pressure balance and gas sealing between the reactors. Based on the sub-pilot results, the hybrid HLC-SMC is able effectively and safely respond to the system changes during heat up and fuel capacity ramping eliminating the need for operator intervention to control the system pressure balance and reactor gas sealing.

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Figure 53. Performance of the HLC-SMCs hybrid controller during fuel injection

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Conclusion

A hierarchical control scheme for the autonomous operation of the chemical looping system is developed. The hierarchical control system consists of a HLC, which is a finite state machine, and several SMCs for different states. The SMCs are designed to maintain the system status variables on pre-designed trajectories during the startup, operation, and shutdown of the system. A HLC-SMCs hybrid controller is designed and successfully tested on the chemical looping pilot plant for the reactor pressure control. The controller is able to ramp up the reactor at a pre-determined rate, maintain the pressure at the setpoint, and depressurize the system at a pre-determined rate. Another HLC-SMCs controller is designed to maintain the gas sealing and pressure balance in the chemical looping sub-pilot plant during the heat-up and fuel injection phases. In all stages of the operation, the controller is able to maintain the desired pressure balance in the reactor.

Thus, the hierarchical control scheme can be applied for the autonomous operation of the chemical looping systems to reduce operator work load and facilitate a more efficient startup, ramping, and shutdown.

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Chapter 6: Conclusions and Future Researches

Conclusions

The chemical looping partial oxidation process is studied for the conversion of carbonaceous fuels to high purity syngas that is suitable for liquid fuel and chemical synthesis. Compared to conventional reforming and gasification processes, the chemical looping partial oxidation process eliminates the need for an ASU, WGS reactor, and catalytic reformer. Thus, the capital and operation cost of the process is expected to be decreased while the fuel conversion efficiency is increased.

The rational in the development of the chemical looping partial oxidation process is discussed in this dissertation. In particular, the selection of the oxygen carrier as well as the design of the reducer is examined. The modified Ellingham Diagram is constructed to illustrate the thermodynamic considerations in the selection of the oxygen carrier.

The ITCMO oxygen carrier is selected based on its thermodynamic property. The addition of titanium in the system stabilizes the Fe(II) compound. When the ITCMO oxygen carrier is reduced to the Fe/FeTiO3 state, it can reach chemical equilibrium with a gaseous product with high CO and H2 content, which avoids the further production of full

143

oxidation byproducts. Therefore, the ITCMO oxygen carrier is suitable for high purity syngas generation in the chemical looping partial oxidation process.

A co-current moving bed reducer design is chosen for the chemical looping partial oxidation process. Due to the plug flow nature of gases and solid sin the moving bed reactor, the gas and solid residence time at the reactor outlet can be well controlled. Thus, gaseous product is in contact with oxygen carriers with desirable oxidation states before leaving the reactor, which is advantageous for preventing full oxidation and carbon deposition. The absence of axial solid mixing and gas bubbles also facilitates the full conversion of solid and gaseous fuels in the moving bed reducer.

The effect of ITCMO on the operation of the oxidizer is analyzed. Because the formation of FeTiO3 stabilizes the Fe(II) compound, it is more difficult to be oxidized by

H2O to form H2 in the oxidizer. The operation line analysis based on the mass balance and thermodynamic property of the counter-current moving bed oxidizer shows that the

144

reduced ITCMO oxygen carrier can only be oxidized to the FeTiO3 oxidation state in the oxidizer, while the reduced Fe2O3 oxygen carrier can be oxidized to Fe3O4 oxidation state. Thus, ITCMO is less efficient for H2 generation from the oxidizer as compared to a

Fe2O3 oxygen carrier.

The concept of using the ITCMO oxygen carrier and the co-current moving bed reducer is tested experimentally to confirm the feasibility of the chemical looping partial oxidation process. In particular, the conversion of CH4 and biomass to syngas is tested in a fixed bed reactor and a bench-scale co-current moving bed reducer. High purity syngas with a H2:CO ratio of 2 is produced from both CH4 and biomass. The syngas composition and solid composition agree with the simulation results obtained from the ASPEN

RGibbs module.

The chemical looping partial oxidation process is scaled up to a 15 kWth sub-pilot scale unit. The unit is fully integrated with a co-current moving bed reducer and a

145

fluidized bed combustor, connected by non-mechanical gas sealing and solid circulation devices. The control instruments and DSC are designed similarly to a large scale system.

Syngas generation in the sub-pilot unit using CH4 is demonstrated. Steady syngas generation operation is achieved with a CH4 conversion of over 99% and a syngas H2:CO ratio ranging from 1.6 to 2.2.

The experiment results confirm that the chemical looping partial oxidation process using ITCMO oxygen carrier and co-current moving bed reducer can effectively convert the gaseous and solid fuels into high purity syngas with desirable gas composition.

In order to facilitate the further scale-up the chemical looping processes with moving bed reactors, the automation of the reactor system is studied. A dynamic simulation scheme for the chemical looping system is developed to further understand the transient behaviors of the system when it is subjected to status changes and disturbances.

By considering the mass balance and hydrodynamic characteristics of the system components, a model consisting of a set of DAEs is developed. The DAEs are solved to simulate the transient behavior of the chemical looping system. The dynamic simulation scheme can be used for control system development and operator training.

146

of the system. During the operations, the controllers are able to maintain the desired values for the system status variables, including the system pressure in the pilot plant and the pressure balance in the sub-pilot unit. Thus, the hierarchical control scheme can be applied for the autonomous operation of the chemical looping systems to reduce operator work load and facilitate a more efficient startup, ramping, and shutdown.

Future Researches

To further develop the chemical looping partial oxidation process, further researches are needed in several areas, including oxygen carrier development, reactor design, operation experience, process system engineering, reactor simulation, and controller design. Specific topics of interest are as follows.

Oxygen Carrier Development

Improvement of oxygen carrier mechanical strength;

Improvement of oxygen carrier reactivity;

Study on the effect of high pressure on the gas-solid reaction.

Methods for reusing discharged oxygen carrier fines.

Reactor Design

Design of solid feeder/distributor for solid fuel injection;

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Development of solid flow rate sensors;

Equipment design for pressurized operation;

Novel reactor design for modulization and small scale reactor.

Operation Experience

Solid fuel conversion in sub-pilot scale unit, including various types of coal and biomass;

Study on the effect of H2O and CO2 on char gasification kinetics;

Study on the fate of sulfur;

Study on the fate of mercury.

Process System Engineering

Optimization of the heat integration scheme for energy recuperation from high temperature gas stream;

Optimization of operation pressure to minimize compression cost.

Reactor Simulation

Dynamic simulation of chemical reactions in partial oxidation system;

Incorporation of heat transfer and chemical reactions in the dynamic simulation scheme;

Integration of chemical looping dynamic model with downstream process and heat exchanger network;

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Design of startup equipment and procedure based on dynamic model scheme.

Controller Design

Application of observers or other methods to minimize chattering in SMCs;

Development of fault detection algorithm;

Control logic design for trip and emergency conditions

HLC-SMCs design for fully autonomous operation of the chemical looping system;

Development of optimization scheme to improve efficiency and economics of chemical looping systems.

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References

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[3] Molburg, J. C., & Doctor, R. D. (2003, September). Hydrogen from steam-methane reforming with CO2 capture. In 20th annual international Pittsburgh coal conference.

[4] Rostrup-Nielsen, J. R. (2002). Syngas in perspective. Catalysis today, 71(3), 243- 247.

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