novel flex fuel oxidation for distributed ... research and development division final project report...

E n e r g y R e s e a r c h a n d D e v e l o p m e n t D i v i s i o n F I N A L P R O J E C T R E P O R T

NOVEL FLEX FUEL OXIDATION FOR DISTRIBUTED GENERATION Base Load Combined Heat and Power with Carbon Capture

APRIL 2017 CE C-500-2017-024

Prepared for: California Energy Commission Prepared by: ZERE Energy and Biofuels, Inc.

PREPARED BY: Primary Author(s): Marisa Zuzga Charles Coronella, Ph.D. . Mingheng Li, Ph.D. Evan Hughes, Ph.D. Helal Uddin ZERE Energy and Biofuels, Inc. 851 Cherry Ave. #27-202 San Bruno, CA 94066 Phone: 562-760-8827 http://www.zerepower.com Contract Number: PIR-11-016 Prepared for: California Energy Commission Michael Kane Contract Manager Aleecia Gutierrez Office Manager Energy Generation Research Office Laurie ten Hope Deputy Director ENERGY RESEARCH AND DEVELOPMENT DIVISION Robert P. Oglesby Executive Director

DISCLAIMER This report was prepared as the result of work sponsored by the California Energy Commission. It does not necessarily represent the views of the Energy Commission, its employees or the State of California. The Energy Commission, the State of California, its employees, contractors and subcontractors make no warranty, express or implied, and assume no legal liability for the information in this report; nor does any party represent that the uses of this information will not infringe upon privately owned rights. This report has not been approved or disapproved by the California Energy Commission nor has the California Energy Commission passed upon the accuracy or adequacy of the information in this report.

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ACKNOWLEDGEMENTS

First we would like to dedicate this report to the memory of George L. Touchton III without whom this project never would have been envisioned. We hope our work has made him proud. We would like to thank the team of people who supported and were vital to the success of this project namely Chuck Coronella, Ph.D., Mingheng Li, Ph.D., Evan Hughes, Ph.D., and Helal Uddin. Also, we thank Reggie Mitchell, Ph.D., Eli Goldstein, Ph.D., Bob Aldrich Ph.D., and Eli Yagor for their advice and encouragement. Finally, we hope that the California Energy Commission staff and specifically our Program Manager Michael Kane will accept our thanks for their assistance that helped make this project possible.

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PREFACE

The California Energy Commission Energy Research and Development Division supports public interest energy research and development that will help improve the quality of life in California by bringing environmentally safe, affordable, and reliable energy services and products to the marketplace.

The Energy Research and Development Division conducts public interest research, development, and demonstration (RD&D) projects to benefit California.

The Energy Research and Development Division strives to conduct the most promising public interest energy research by partnering with RD&D entities, including individuals, businesses, utilities, and public or private research institutions.

Energy Research and Development Division funding efforts are focused on the following RD&D program areas:

• Buildings End-Use Energy Efficiency

• Energy Innovations Small Grants

• Energy-Related Environmental Research

• Energy Systems Integration

• Environmentally Preferred Advanced Generation

• Industrial/Agricultural/Water End-Use Energy Efficiency

• Renewable Energy Technologies

• Transportation

Novel Flex Fuel Oxidation for Distributed Generation-Base Load Combined Heat and Power with Carbon Capture is the final report for the Novel Flex Fuel Oxidation for Distributed Generation project (contract number PIR-11-016) conducted by ZERE Energy and Biofuels, Inc. The information from this project contributes to Energy Research and Development Division’s Environmentally Preferred Advanced Generation Program.

For more information about the Energy Research and Development Division, please visit the Energy Commission’s website at www.energy.ca.gov/research/ or contact the Energy Commission at 916-327-1551.

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ABSTRACT

Californians require cost effective fuel flexible technology to reduce or eliminate the problems associated with biogas, bio-liquid, and biomass waste disposal. There is an immediate need for eliminating or minimizing flaring and venting of gases from dairy wastes, landfills, and waste water treatment facilities. Combining heat, power, and cooling that uses fuel flexible biogas technology will benefit California’s utility ratepayers by ensuring price stability and lowering electric rates.

This project designed, constructed, operated, and tested a ZERE Independent Internal Oxidation system at laboratory and prototype scales to produce heat and power using fuel flexible biogas. The prototype demonstrated that a ZERE system fueled with untreated biogases (such as dairy digester and waste water treatment) and natural gas could produce electricity and heat while also functioning as a carbon capture technology that can emit almost no pollution and surpass California Air Resources Board standards.

ZERE generated life cycle, techno-economic, thermodynamic, and computational fluid dynamic models with potential for use in various ZERE technologies. Those process models were validated using data collected with lab scale experiments. That data was then used to design and build a 100 kiloWatt thermal input gas fueled prototype system. Facility integration and initial functionality testing of the prototype was completed.

Keywords: biogas, digester gas, renewable energy, chemical looping combustion, carbon capture, combined heat and power

Please use the following citation for this report:

Zuzga, Marisa; Coronella, Charles; Li, Mingheng; Hughes, Evan; Uddin, Helal (ZERE Energy and Biofuels, Inc.). 20016. Novel Flex Fuel Oxidation for Distributed Generation. California Energy Commission. Publication number: CEC-500-2017-024.

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TABLE OF CONTENTS

Acknowledgements ................................................................................................................................... i

PREFACE ................................................................................................................................................... ii

ABSTRACT .............................................................................................................................................. iii

TABLE OF CONTENTS ......................................................................................................................... iv

LIST OF FIGURES ................................................................................................................................. vii

LIST OF TABLES ................................................................................................................................... vii

EXECUTIVE SUMMARY ........................................................................................................................ 1

Introduction ........................................................................................................................................ 1

Project Purpose ................................................................................................................................... 1

Project Process and Results ............................................................................................................... 1

Project Benefits ................................................................................................................................... 5

CHAPTER 1: Introduction ...................................................................................................................... 6

1.1 Problem Statement ..................................................................................................................... 6

1.2 Project Goal ....................................................................................................................................... 7

Thermodynamic Process Screening and Analysis Using Simulation Models ........................... 7

Economic and Life Cycle Analysis .................................................................................................. 8

1.3 Project Results ................................................................................................................................... 8

CHAPTER 2: Thermodynamic Analysis ............................................................................................ 10

2.1 Background ............................................................................................................................... 10

2.2 Design Cases ............................................................................................................................. 11

2.2.1 Three Reactor Design (Biogas Diluted with H2O) .............................................................. 11

2.2.2 Three Reactor Design (Biogas Diluted with CO2) ............................................................... 12

2.2.3 Effect of Al2O3 on System Performance ............................................................................... 13

2.2.4 Two Reactor with Recycle Design ........................................................................................ 13

2.2.5 Sulfur Removal ........................................................................................................................ 16

2.2.6 Final Prototype Design ........................................................................................................... 18

2.3 Experimental Validation of Thermodynamic Design ............................................................... 19

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CHAPTER 3: Economic and Life Cycle Analysis .............................................................................. 24

3.1 Commercial Scale Life Cycle Analysis .................................................................................. 24

3.1.1 Life Cycle Benefits ............................................................................................................ 24

3.1.2 Final ZERE Case with Steam Power Generation ......................................................... 25

3.1.3 Emissions ........................................................................................................................... 25

3.1.4 Parameters of the Final ZERE Case ............................................................................... 26

3.1.5 Cases in the Preliminary LCA Report ........................................................................... 27

3.1.5 Final ZERE Case ............................................................................................................... 28

3.1.6 Three Biogas Cases ........................................................................................................... 30

3.1.7 Summary on CO2 Emissions ........................................................................................... 31

3.1.8 Conclusion on CO2 Emissions ........................................................................................ 32

3.1.9 Criteria Pollutant Emissions (NOx, SO2, PM, CO) ...................................................... 32

3.2 Commercial System Economic Benefit Analysis ................................................................. 33

CHAPTER 4: Computational Fluid Dynamics .................................................................................. 36

4.1 CFD Hydrodynamic Model of Semi Continuous Reactors ................................................ 36

4.1.1 Introduction ...................................................................................................................... 36

4.1.2 ZERE Prototype Reactors Configuration ...................................................................... 37

4.1.3 Simulated Reactors Configuration................................................................................. 37

4.1.4 Results and discussions ................................................................................................... 40

4.1.5 Conclusions .............................................................................................................................. 52

4.2 Model Reaction Kinetics in the Semi Continuous Reactors ............................................... 53

4.2.1 Fluid Dynamics of the Model ......................................................................................... 53

4.2.2 Material balances into the reactor .................................................................................. 54

4.2.3 Kinetic Model for Oxygen Carrier ................................................................................. 55

4.2.4 Break Through Time for Fuel ......................................................................................... 56

4.3 Validate Reactor Models ......................................................................................................... 57

4.3.1 Results of Laboratory Experiments ............................................................................... 57

4.3.2 Comparison of model with experimental results ........................................................ 59

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4.3.3 Conclusions .............................................................................................................................. 64

CHAPTER 5: Lab Scale Experimentation .......................................................................................... 66

5.1 Introduction and Summary of Experiments ........................................................................ 66

5.1.1 Experimental Setup and Apparatus .............................................................................. 66

5.1.2 Oxygen Carrier Evaluation ............................................................................................. 69

5.1.3 Fluidization Experiments ................................................................................................ 70

5.2 Oxygen Carrier Cycling with Production of Steam and CO2 from Biogas and Natural Gas 72

5.3 Oxygen Carrier Cycling with Production of Steam and CO2 from Biomass Solid Fuel Oxidation ............................................................................................................................................... 75

5.4 Sulfur Capture in a Fluid Bed ................................................................................................ 78

CHAPTER 6: Prototype Flex Fuel CHP Plant Design & Construction ......................................... 82

6.1 Reactor Design .......................................................................................................................... 82

6.1.1 Introduction ............................................................................................................................. 82

6.1.2 Descriptions ............................................................................................................................. 82

6.1.3 Instrumentation and Control ................................................................................................. 89

6.1.4 H2S or SO2 Removal for Gaseous Fuels ................................................................................ 89

6.2 Prototype Plant Design ........................................................................................................... 95

CHAPTER 7: ZERE Prototype Flex Fuel CHP Plant Testing and Operation ............................. 101

CHAPTER 8: Commercial System Plan ............................................................................................ 103

8.1 Design of ZERE Commercial Scale System ........................................................................ 103

8.1.1 Introduction ........................................................................................................................... 103

8.1.2 Commercial System Design .......................................................................................... 103

8.1.3 Scalability ........................................................................................................................ 108

8.1.4 Fully Fluidized Continuous System ............................................................................ 109

8.2 Rate Payer Benefits ................................................................................................................ 111

8.2.1 Ratepayer Benefits of ZERE Gas Only Systems ......................................................... 111

8.2.2 Added Benefits of Solid Fuel Capability ..................................................................... 115

8.3 Commercialization Plan ........................................................................................................ 115

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GLOSSARY ............................................................................................................................................ 117

REFERENCES ........................................................................................................................................ 119

APPENDIX A: Details of Economics Calculation ........................................................................... A-1

APPENDIX B: Bubble Size and Frequency Calculations .............................................................. B-1

APPENDIX C: Prototype Design Details.......................................................................................... C-1

APPENDIX D: Control Philosophy .................................................................................................. D-1

LIST OF FIGURES

Figure 1: Affect of Catalyst Amount on Reactor Duties in the Presence of Al2O3 (Biogas Feed Diluted with CO2) .................................................................................................................................... 14

Figure 2: Affect of Catalyst Amount on Reactor Duties in the Presence of Al2O3 (Biogas Feed Diluted with H2O) .................................................................................................................................... 14

Figure 3: Affect of Catalyst Amount on CO Formation in the Presence of Al2O3 (Biogas Feed Diluted With H2O) ................................................................................................................................... 15

Figure 4: Affect of Catalyst Amount on CO Formation in the Presence of Al2O3 (Fuel Diluted With H2O) ................................................................................................................................................. 15

Figure 5: Flow Rates and Temperatures of Data Collected on 2/22/16 ............................................. 20

Figure 6: Experimental Gas Flows and Emissions .............................................................................. 23

Figure 7: Effect of Baffles on Bubble Break-Up in the Air Reactor Contour of Gas Volume Fraction at T= 6.0 S with Uniform Inlet Gas Velocity = 0.675 M/S .................................................... 40

Figure 8: Effect of Baffles on Bubble Break-Up in the Fuel Reactor Contour of Gas Volume Fraction at T= 6.0 S with Uniform Inlet Gas Velocity = 0.18 M/S ...................................................... 41

Figure 9: Effect of Inlet Gas Velocity Condition in the Fuel Reactor. Contour of Gas Volume Fraction at T= 6.0 S. In Both Cases, 33% Open Area in the Baffle Hole ............................................ 43

Figure 10: Baffle Effect on Bed Expansion in the Air Reactor (Uniform Inlet Gas Velocity = 0.675 M/S) ............................................................................................................................................................ 45

Figure 11: Baffle Effect on Bed Expansion in the Fuel Reactor (Uniform Inlet Gas Velocity = 0.18 M/S) ............................................................................................................................................................ 45

Figure 12: Baffle Effect on Bed Expansion in the Fuel Reactor (Uniform Inlet Gas Velocity = 0.18 M/S and Jet Velocity = 0.475 M/S) .......................................................................................................... 46

Figure 13: Baffle Effect on Bed Pressure Drop Fluctuation in the Air Reactor (Uniform Inlet Gas Velocity = 0.675 M/S) ............................................................................................................................... 48

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Figure 14: Baffle Effect on Bed Pressure Drop Fluctuation in the Fuel Reactor (Uniform Inlet Gas Velocity = 0.18 M/S) ................................................................................................................................. 49

Figure 15: Baffle Effect on Bed Pressure Drop Fluctuation in the Fuel Reactor (Uniform inlet gas velocity = 0.18 m/s and jet velocity = 0.475 m/s) .................................................................................. 50

Figure 16: Baffle Effect on Solid Flux in the Air Reactor (uniform inlet gas velocity = 0.675 m/s)52

Figure 17: Schematic of the Overall Model Used in this Study ......................................................... 56

Figure 18: Inlet Gas Flow Rates During Successive Redox Reaction of Cuo/Al2O3 Particles With Methane ..................................................................................................................................................... 57

Figure 19: Bed Temperature and Pressure Drop during Successive Oxidation and Reduction of Cuo/Al2O3 Particles .................................................................................................................................. 58

Figure 20: Exit Gas Composition During Reduction of Cuo/Al2O3 Particles with Methane ....... 59

Figure 21: Staging of the Solid Bed Used in the Model Based on the Mass of Solid Oxygen Carrier Used in the Experiment ............................................................................................................. 60

Figure 22: Axial Profile of CH4 Leaving from Different Stages Considered in the Model of the Fluidized Bed Reactor ............................................................................................................................. 61

Figure 23: Axial Profiles of Gas and Solids Concentration During First Looping Condition ....... 62

Figure 24: Concentration Profile - Reduction Period with CH4 as Reducing Gas as Temperature of ~720°C .................................................................................................................................................... 63

Figure 25: Presence of Different Oxides in Oxygen Carrier Supplied by Clariant ......................... 64

Figure 26: ZERE Lab Scale Reactor Schematic ..................................................................................... 67

Figure 27: Lab Scale Reactor During Cold Fluidization Test ............................................................. 68

Figure 28: Lab Scale Reactor With Red Hot (800°C) Copper Particles ............................................. 68

Figure 29: Thirty Percent Copper/Copper Oxides on Alumina Support ......................................... 70

Figure 30: Thirty Percent Copper/Copper Oxides on Alumina Support (30X Magnification) ..... 70

Figure 31: Cold Particle Fluidization Test ............................................................................................ 71

Figure 32: Hot Particle Fluidization Test .............................................................................................. 71

Figure 33: Flow Profile of Fuel and Air Cycles (approximately 750°C) ........................................... 73

Figure 34: Flow Profile of Three Fuel and Air Cycles (approximately 750°C) ................................ 74

Figure 35: Emissions Profile of 3 Fuel and Air Cycles at approx. 750°C .......................................... 74

Figure 36: NOx Emissions During Cycling .......................................................................................... 75

Figure 37: Lab scale Reactor with Solid Fuel Feed .............................................................................. 76

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Figure 38: Solid Fuel Experiment with Corn Stover – Test 1 ............................................................. 77

Figure 39: Solid Fuel Experiment with Corn Stover - Test 2 .............................................................. 78

Figure 40: Baseline Gas Composition of Test Biogas .......................................................................... 79

Figure 41: Biogas Test without Lime ..................................................................................................... 80

Figure 42: Biogas Test with Lime in the Bed ........................................................................................ 81

Figure 43: Reactor Schematic .................................................................................................................. 84

Figure 44: Cyclone Design ...................................................................................................................... 86

Figure 45: Baffle Design .......................................................................................................................... 87

Figure 46: Distributor Plate Layout, with 12 Tuyeres ......................................................................... 87

Figure 47: Tuyere Design ........................................................................................................................ 88

Figure 48: Plenum Chamber ................................................................................................................... 89

Figure 49: Prototype Reactor System .................................................................................................... 96

Figure 50: 15kW Steam Turbine ............................................................................................................. 97

Figure 51: Steam System ......................................................................................................................... 98

Figure 52: Exhaust Gas Heat Exchanger ............................................................................................... 98

Figure 53: Control Panel .......................................................................................................................... 99

Figure 54: HMI Screen Examples ........................................................................................................... 99

Figure 55: Prototype Preheat System .................................................................................................. 100

Figure 56: ZERE Prototype in Shipping Configuration .................................................................... 101

Figure 57: ZERE Prototype with Exhaust Duct and Insulation ....................................................... 102

Figure 58: Simplified Reactor Design .................................................................................................. 105

Figure 59: J-Valve Design ...................................................................................................................... 107

Figure 60: Cyclone 1 Design ................................................................................................................. 107

Figure 61: Cyclone 2 Design ................................................................................................................. 108

Figure 62: ZERE Patented 3-Stage Reactor Design ............................................................................ 110

Figure 63: Cal ISO Hourly Average Net Load 2-22-16 ..................................................................... 112

Figure 64: Renewable Distributed Generation in California (20 MW or Smaller, Includes Self-Generation) by Fuel Type ..................................................................................................................... 112

Figure 65: U.S. EPA Map of Operating and Candidate Landfill Projects ...................................... 114

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Figure 68: Bubble Area in the Fuel Reactor without Baffle at t = 6.0 s. ............................................... 1

Figure 69: Bubble Areas in the Fuel Reactor without Baffle at t = 6.0 s ............................................. 2

Figure 70: Circular Bubble sizes in the Fuel Reactor without Baffle at t= 6.0 s ................................. 3

Figure 71: Circular Bubble sizes in the Fuel Reactor without Baffle at t= 6.0 s. ................................ 4

Figure 72: Circular Bubble sizes > 0.5cm2 in the Fuel Reactor without Baffle at t= 6.0 s. ................. 5

Figure 73: Circular Bubble sizes > 0.5cm2 in the Fuel Reactor without Baffle at t= 6.0 s. ................. 6

Figure 74: System P&ID ............................................................................................................................ 1

Figure 75: System P&ID Section A1 ........................................................................................................ 2



Figure 78: System P&ID Section B1 ......................................................................................................... 5



Figure 81: System P&ID Section C ........................................................................................................... 8

LIST OF TABLES

Table 1: Stream Properties of Three Reactor Design (Biogas diluted with H2O) ............................ 11

Table 2: Heat Duties and Steam Flow Rate in Three Reactor Design (Biogas Diluted with H2O) 12

Table 3: Stream Properties of Three Reactor Design (Biogas Diluted with CO2) ............................ 12

Table 4 Heat Duties and Steam Flow Rate in Three Reactor Design (Biogas Diluted with CO2) . 13

Table 5: Stream Properties in Two Reactors with Recycle Design .................................................... 16

Table 6: Duties of Process Units in Two Reactors with Recycle Design ........................................... 17

Table 7: Stream Properties of Steam System in Two Reactors with Recycle Design ...................... 17

Table 8: Equilibrium Calculation of Biogas CLC with CaO at 800 ºC and 1.2 bar .......................... 17

Table 9: Equilibrium Calculation of Biogas CLC with CaO at 800 ºC and 1.2 bar .......................... 18

Table 10: Duties of Process Units in Final Two Reactor Design ........................................................ 19

Table 11: Per Cycle Temperature Change and Heat Release ............................................................. 21

Table 12: End of Cycle Temperature Drop ........................................................................................... 22

Table 13: ZERE vs. Conventional Biogas-to-Power CHP ................................................................... 34

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Table 14: Physical Properties of Simulation Parameters .................................................................... 39

Table 15: Summary of Bubbles ............................................................................................................... 42

Table 16: Summary of Average Bed Expansion ................................................................................... 47

Table 17: Parameters Used in Model Predictions ................................................................................ 62

Table 18: Bed Particles Evaluated .......................................................................................................... 69

Table 19: Summary of Design Comparisons ........................................................................................ 83

Table 20: Summary of Distributor Design ............................................................................................ 88

Table 21: Findings and Recommendations about H2S Effects on CLC ............................................. 90

Table 22: Advantages and Limitations of Different Technologies to Remove H2S from Biogas .. 93

Table 23: Summary of Solid vs. Gas Only Design Comparisons..................................................... 104

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EXECUTIVE SUMMARY

Introduction Climate change is one of the most significant environmental and engineering challenges faced by society in modern times. A necessary strategy for reducing human contributions to climate change is to cut greenhouse gas emissions from electric power plants, which account for 20 percent of California carbon dioxide (CO2) emissions.0F

1 In addition, Californians require cost effective fuel flexible technology to reduce or eliminate the problems associated with biogas, bio-liquid, and biomass waste disposal. In particular, it is necessary to eliminate or minimize flaring or venting these gases from dairy wastes, landfills, and waste water treatment facilities.

Project Purpose California is looking for innovations in combined cooling, heating, and power and combined heat and power (CCHP/CHP) technology to provide price stability, lower electric rates, and provide environmental benefits. The ZERE team created a flexible fuel Air Independent Internal Oxidation (AIIO) technology solution to generate electric power, heating, and cooling that has the potential to efficiently and effectively reduce or eliminate challenges associated with biogas and biomass waste disposal. This new system is an oxidation technology that chemically combines fuel and oxygen and uses a solid stage oxygen carrier to capture nearly 100 percent of the fuel energy with near zero emissions. The spent oxygen carrier is then regenerated by reacting with oxygen in the air, again forming near zero emissions. The fuel energy is converted to electric power and heat by steam and/or gas turbine power cycles. The system can also be configured to provide process heat and steam only.

This project produced heat and power using fuel flexible biogas through the design, construction, operation, and testing of a ZERE AIIO combined heat and power system at the laboratory and prototype scale. The AIIO prototype demonstrated that a ZERE system could produce electric power and heat fueled with various forms of biogas and natural gas while emitting almost no pollution, surpassing California Air Resources Board 2007 emissions standards. The success of this project could lead to expanding the portfolio of base load renewable energy systems available in California. Specifically, expanding renewable energy systems that can produce heat, power, and cooling from gaseous or biomass fuels while achieving emissions targets.

Project Process and Results Researchers with ZERE Energy performed thermodynamic, computational fluid dynamic, life cycle, and techno-economic analysis on the patented AIIO. The lab scale experiments helped develop a full understanding of the fluidized bed behavior to design and build a 100 kW thermal input prototype. Finally, ZERE developed a design for a commercial system and assessed the benefits to California ratepayers.

1 2016 Edition of California GHG Emission Inventory. Figure 4, p.2. https://www.arb.ca.gov/cc/inventory/pubs/reports/2000_2014/ghg_inventory_trends_00-14_20160617.pdf

https://www.arb.ca.gov/cc/inventory/pubs/reports/2000_2014/ghg_inventory_trends_00-14_20160617.pdf

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Thermodynamic Analysis Researchers performed a thermodynamic analysis with three different options for operating the AIIO system using code written specifically for this project. The code was validated by comparing results for a few of the cases with results given by Aspen, an industry standard software package. The ZERE team evaluated three systems. The first system included hot gas recycle in the fuel reactor, the second system included an extra air reactor, and the final system evaluated had unequal inlet gas flow rates for the fuel and the air reactors with the fuel flow rate less than the air flow rate. The analysis considered these factors:

• Reactor inlet gas compositions, temperatures, and flow rates.

• Reactor outlet gas compositions, temperatures, and flow rates.

• Fluidized bed characteristic including material, mass, temperature, and reactivity.

• Heats transfer throughout the system.

The system with the unequal inlet flow rates was selected for the prototype system design. Once ZERE selected that design, an additional analysis was preformed on model emissions behavior with inlet gas contaminants such as hydrogen sulfide (H2S) and ammonia (NH3)

present in biogas. The thermodynamic analysis predicted good heat generation for use in power generation and near zero emissions from the fuel and air reactors with and without the biogas contaminates in the fuel.

Computational Fluid Dynamic Analysis Despite their widespread application in the process industries, the design of fluidized bed reactors is still very challenging. Researchers created computational fluid dynamic models to analyze the behavior of their ZERE reactor systems, making them valuable tools in system design. Specifically, the models analyzed the behavior of the reactor beds, the flow rates of the air reactor, and the fuel reactor. The analysis was used to design system elements, specifically the inlet distributor, flow baffles, and internal cyclone. Specific items considered in the analysis were minimum fluidization rates for the bed, gas bubble formation and size, fine particle entrainment, and removal from the gas stream. The inlet distributor design incorporated a system of tuyeres (small individual inlets), to distribute the inlet gas flow. The baffles were designed to prevent slugging (large gas bubbles moving through the bed without reacting with the bed material) and reduce bubble size. For the prototype, an internal cyclone was designed to return entrained bed particles to the bed before the exhaust gases leave the reactor. In the final design (which incorporates solid fuels) two cyclones were designed: one to return bed material back to the fluidized bed and one to remove ash and other particulates from the exhaust gas stream.

Life Cycle and Economic Analysis Taking into account thermodynamic modeling and estimates for system efficiency and auxiliary load, ZERE performed an analysis of a 5 megawatt ZERE system with CO2 capture and sequestration compared to a traditional natural gas power plant. The upstream and downstream emissions were based on the California Greenhouse Gases, Regulated Emissions,

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and Energy Use in Transportation model. This case analyzed these efficiency values that lead to a 21.8 percent overall net efficiency:

• 90 percent efficiency energy content of the biogas (CH4 fraction at 1000 Btu/scf) to hot gas output.

• 30 percent gross thermal (1st Law) efficiency steam-to-electricity. • 95 percent generator efficiency. • 85 percent ratio of net to gross electric generation (15 percent auxiliary power demand).

The indirect fossil CO2 emissions from processes upstream of the ZERE system are taken into account and reduce the negative emission factor. Therefore, the negative emission factor for the ZERE power system is -1166 grams CO2 per kWh.

For the economic analysis, the avoided carbon dioxide from the capture of CO2 by a ZERE system gives a net avoided emission advantage versus conventional of 955 grams CO2 per kWh. Taking into account estimated for total system operation and maintenence (O&M) cost, the total cost of CO2 capture is $35.13 per metric ton of CO2. $35 per ton is a reasonable price for avoiding fossil CO2 compared to estimates of the cost of various options investigated in studies of economics of climate change and fossil emissions. However, it is high compared to the current price of fossil carbon offsets and allowances in the “cap-and-trade” system that exists in California under the AB32 (2006) Climate Change Solutions Act, where the current price is $12 or $13 per metric ton CO2. If the two major items of higher cost for the ZERE system were reduced to the same as conventional systems, then the extra cost of ZERE would be 1.3 cents/kWh instead of 3.3 cents/kWh and the cost of CO2 emission avoided would be close to the current cap-and-trade price range. A target for the ZERE system is $2,000/kW installed capital cost and 2.5 cents/kWh total of fixed plus variable operating cost.

Lab Scale Experimentation A quartz glass reactor capable of supporting a fluidized bed of 100-200 grams was used to perform over 100 separate experiments to evaluate nine different bed materials, evaluate the long term performance of the particles selected for the prototype and assess the emissions performance of AIIO systems. Researchers measured inlet gas flows and outlet gas emissions (including hydrogen, carbon monoxide, carbon dioxide, total hydrocarbons, nitrogen oxides, and oxygen). Emissions of hydrogen sulfide and sulfur dioxide were measured in place of nitrogen oxides to verify AIIO’s ability to capture fuel hydrogen sulfide. Measurements were also taken of inlet, bed, and outlet temperature and differential pressure across the fluidized bed.

The experiments confirmed predictions and results from previous experiments that systems with bed material copper concentration of as low as 15 percent on an alumina substrate can reduce gaseous fuels and produce an outlet gas stream of just CO2 and steam. In addition, experimental results on particle fluidization flow rates were used as parameters in the computational fluid dynamics analysis that was used to design the prototype. Finally, experiments with the addition of lime/calcium oxide proved that ZERE systems can accept untreated biogas while still maintaining zero pollutant emissions in the outlet gas stream. The

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project plan included a series of solid fuel tests with the lab scale equipment. Multiple tests were performed with corn stover as the fuel, but the apparatus could not be adapted to properly add fuel at a steady rate despite multiple attempts and batch tests were only partially successful.

Prototype Design, Construction and Testing A 100 kilowatt thermal input prototype was designed using information gained from the thermodynamic and computational fluid dynamic analysis and the lab scale experimentation. With the exception of the reactor system, the majority of the prototype design was determined by the requirements of the 15 kW steam turbine. The major components of the prototype are the reactors, the steam turbine, the exhaust gas heat recovery steam generator, the steam piping, the preheat system, the control system, and the bed material. The detailed design of the reactors was completed in conjunction with the computational fluid dynamic analysis. The HRSG design was based in large part on the thermodynamic analysis. The prototype design process started with an instrumentation diagram and control philosophy that was used to generate a detailed system design for a skid mounted prototype. The prototype was fabricated by an offsite integrator and then brought to the ZERE test site at Prospect Silicon Valley, where it was assembled and integrated into the facility. Researchers performed preliminary system functionality testing and showed operability of system monitoring instrumentation, such as thermocouples and pressure transmitters as well as functionality of all system control valves and the preheat system.

Prototype testing was restricted to testing the various subsystems and attempting to tune them to achieve stable operation for the fully integrated ZERE system. Unavoidable project delays resulted in insufficient time for the project team to achieve steady state operation of the full system. Steady state operation is needed to be able to test the system’s integrated performance and prove its power production capabilities. Since the end of this project, the project team has continued advancing the ZERE concept on a part time basis as time and funding allow. As of April 2017, they have completed their shake down of most of the subsystems and rewrote a large part of the control program to facilitate future testing. Continuation of this research is contingent upon the ability of ZERE Energy and Biofuels to secure additional funding.

Commercialization Plan Researchers designed a system capable of expanding ZERE’s market by including solid fuels. The design basis is the same, with two bubbling bed reactors that alternate between fuel mode and air mode. The major differences in the design were an increase in bed and reactor size and a change in the cyclone arrangement. The internal cyclone was removed and a system of two external cyclones was added. This allows for the first cyclone to separate and return materials to the bed while the second cyclone can be used for ash and particulate removal.

Researchers also considered a second design for larger scale systems. Mid-size systems can use the same switching design as the current system with the addition of more reactor vessels to allow for multiple units to be in air mode or fuel mode as needed at any given point in time. However, the multi-reactor switching design eventually becomes too complicated and expensive. These complications cause ZERE’s patented three-stage system design to become the

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better option. The three-stage design uses a series of three reactors that feed into each other. The first two reactors will be bubbling bed fuel reactors and the third reactor will be a fully fluidized air reactor with cyclone separation. This new process design allows for larger scale solid fuel systems.

A two phase development plan was created to bring ZERE technology to market. The first phase concentrates on procuring funding for company operations and completing the technical work needed to transition the prototype design into a fully autonomous commercial system with remote operations. The second phase concentrates on deployment of the first commercial system and increasing ZERE market opportunities by completing development of a larger scale, solid fuel system.

Project Benefits The patented ZERE process can establish California as a leader in controlling CO2 emissions. The research has demonstrated that ZERE technology has the potential to be used as a cost effective carbon capture technology while using renewable resources within California.

Liquefying the CO2 and selling it to bottlers, would provide ZERE a secondary revenue stream of roughly $2 million additional per year per 5MW plant. The low nitrogen oxide, sulfur dioxide, and particle emissions from ZERE systems would help meet Environmental Protection Agency and California Air Resources Board emissions targets, especially in areas that have significant amounts of biomethane capacity and high rates of severe to extreme ratings for air quality non-attainment. Furthermore, biogas-fueled ZERE systems operate as a base load power source that can provide non-intermittent renewable electricity that can contribute to stabilizing the electrical grid.

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CHAPTER 1: Introduction 1.1 Problem Statement Fuel flexible distributed combined cooling, heat, and power and combined heat and power (CCHP/CHP) can provide price stability, lower electric rates, and environmental benefits for California’s utility ratepayers. These cost effective fuel flexible technologies must reduce or eliminate the challenges associated with biogas, bio-liquid, and biomass waste disposal. In particular, it is essential to eliminate or minimize flaring or venting of gases from dairy wastes, landfills, and waste water treatment facilities. This project continued work done under the pervious California Energy Commission - Energy Innovation Small Grant (EISG), Grant#55181A-07/05, and work completed with the Copper Alliance support.

Present technology must deal with the immediate biogas challenge of venting, flaring, or cleaning the bio-gas and using it in reciprocating engines or gas turbines for CHP/CCHP production. The first two alternatives have serious shortcomings because neither produces a useful product and both have serious air pollution and greenhouse gas impacts on the environment. The third alternative is hampered by carbon dioxide in the biogas seriously impacting the efficiency of the reciprocating engine or gas turbine. Biogas contaminants such as hydrogen sulfides and silanes must be scrubbed before introducing the gas to the reciprocating engine or gas turbine. The gas cleaning system is expensive, has a significant parasitic load, and is a significant contributor to operation and maintenance (O&M) cost. Even when using Best Available Control Technology (BACT) – lean burn and Selective Catalytic Reduction (SCR) -- both of these end use technologies produce NOx, particulate, carbon monoxide, SOx, Unburned Hydrocarbons (UHCs), and carbon dioxide.

A necessary strategy for reducing human contributions to climate change is to cut greenhouse gas emissions from electric power plants, which account for 40% of U.S. CO2 emissions (Stocker, et al. 2014). The conventional CO2 separation techniques are based on post-combustion processes e.g. amine scrubbing. As opposed to the conventional CO2 separation from exhaust gas in power plants and refineries, ZERE’s Air Independent Internal Oxidation (AIIO) a patented form of chemical looping combustion (CLC) is a technology with inherent separation of CO2. AIIO uses two fluidized beds, an air reactor and a fuel reactor, to mix an oxygen-carrier with a combustion fuel. The “looping” term in CLC refers to the cycle by which the oxygen carrier – typically a metal oxide – undergoes reduction in the fuel reactor and is subsequently re-oxidized in the air reactor. It is an effective method for carbon dioxide capture, as CO2 is separated from other reaction products by the nature of the process (Abad, et al. 2010) (Adanez, et al. 2012)

In addition, at least 12 California counties had severe or extreme 8-Hour Ozone nonattainment in 2015 and eight counties had serious nonattainment for PM2.5 (U.S. Environmental Protection Agency 2015). These environmental factors can also be addressed with implementation of ZERE technology.

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Application of ZERE’s technology would generate electric power and heat, with near zero emissions that beat California Air Resources Board (CARB) 2007 standards. In ZERE’s patented Air Independent Internal Oxidation (AIIO) technology, fuel is oxidized internally using a solid stage oxygen carrier, capturing nearly 100% of the fuel energy and forming near zero emissions. ZERE technology accepts untreated biogas. Fuel oxidation and overall system efficiency are not affected by inert constituents such as carbon dioxide in the biogas and contaminants such as hydrogen sulfide and silanes are trapped in the reactor system bed. Carbon dioxide from the oxidation is not emitted to the atmosphere and can be liquefied for sale or sequestration. Regenerating the spent oxygen carrier in air takes place at a low temperature forming no NOx or other pollutants.

1.2 Project Goal This project produced fuel flexible biogas CHP through the design, construction, operation, and test of a ZERE Air Independent Internal Oxidation CHP system at the laboratory and prototype scale. The ZERE prototype demonstrated production of electric power and heat in a ZERE system fueled with untreated biogases (i.e. dairy digester, waste water treatment), natural gas, and mixtures thereof, while out performing CARB 2007 emissions standards.

In addition to the project goal, the following were achieved.

• Quantify and rank the technical and economic performance of AIIO systems utilizing multiple flex fuel configurations based on process models.

• Demonstrate the ability of ZERE CHP systems to operate on untreated biogas.

• Demonstrate that the CARB 2007 emissions standards are beat by ZERE CHP systems while operating on untreated biogas fuels, natural gas or mixtures thereof.

• Verify ZERE CHP system sustainability through full system life cycle analysis.

• Develop near term and long term commercialization path for ZERE fuel flexible CHP systems.

The project successfully completed a series of tasks designs to support the overall project goals and objectives.

Thermodynamic Process Screening and Analysis Using Simulation Models The thermodynamic modeling tasks were: 1) to investigate system yields including electricity, and heat, and project system emissions considering land fill gas, dairy gas, natural gas and combinations thereof as possible fuels 2) analyze each of the process models developed on an input/output basis 3) analyze candidate development paths which would enable adding bioliquids and biomass solids as fuels

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Economic and Life Cycle Analysis These tasks evaluated the system with the best technical performance, as determined in the thermodynamic analysis, and determine the best economic performance and life cycle benefits from that system.

Reactor Design Utilizing Computational Fluid Dynamics (CFD) The CFD tasks: 1) used modern methods of computational fluid dynamics (CFD) to predict the hydrodynamic performance of the two reactors comprising the AIIO reactor system 2) predicted reaction rates and heat loads in both reactors comprising the AIIO reactor system 3) validated the reactor models using experimental data 4) performed a CFD analysis of ZERE system with expanded flex fuel capability using gaseous and solid fuels.

Lab Scale Experimentation The team verified production of steam and CO2, verify near-zero emissions and validate the system models in a batch mode lab scale reactor operating alternately in fuel mode and air mode through the collection and analysis of lab scale reactor parameters such as inlet, outlet, and bed temperature as well as inlet and outlet gas compositions and flows.

Prototype Design, Construction and Testing The team 1) used the reactor system designed using CFD and developed a 100kW thermal input prototype system design 2) procured and constructed the prototype system 3) tested and operated the prototype system

Commercial System Plan This task directed the team to 1) integrate data and lessons learned from all previous tasks into a commercial scale ZERE plant design 2) develop a strategy for bringing ZERE technology to market 3) assess the commercial and rate payer benefits of being able to use a wider range of waste fuels and increase system fuel flexibility. Benefits considered included fuel price stability, fuel supply stability, environmental benefits (such as reduction of open burn emissions), additional sources of renewable fuel, and consistency of energy supply.

1.3 Project Results ZERE accomplished many of the project objectives including life cycle, techno-economic, thermodynamic, and computational fluid dynamic models created for various different implementations of ZERE technology. Those process models were validated using data collected with lab scale experiments and the data were used to design and build a 100kW thermal input gas fuels prototype system. Facility integration and initial functionality testing of the prototype was also completed.

Two areas of work were not completed as proposed. The solid fuel experiments conducted in the lab scale equipment did not produce usable results because the solid fuel particles exiting the fluidized bed before being fully reducted. Additionally, delays in the prototype construction made it impossible to finish the prototype testing and show full power generation capability. However, ZERE plans to continue the project and complete the goals after the conclusion of this grant.

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Overall, ZERE product development has been helped tremendously by the California Energy Commission investment. The building of a prototype with these funds should allow ZERE to attract investment through future grants or private funders to make the transition to a commercially available product that can benefit California rate payers.

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CHAPTER 2: Thermodynamic Analysis 2.1 Background In the thermodynamic analysis, thermodynamic data (e.g., heat capacity, enthalpy, entropy and Gibbs free energy) of all species are correlated as functions of temperature using coefficients provided by NASA Chemical Equilibrium with Applications (CEA) and/or HSC Chemistry database. Chemical equilibrium is assumed in both fuel reactors and air reactors. The equilibrium calculation is based on minimization of Gibbs free energy of the system as follows:

,

(1)

where is the amount of species j at equilibrium and is in a feed. A is atomic matrix with representing number of elements i in species j. N is total number of species (gas and solids). T is temperature of the reactor at equilibrium. P is pressure. H is enthalpy, S is entropy. G is Gibbs free energy. The superscript 0 on H, S, G and P represents standard condition (i.e., 1 bar). R is gas constant (8.314 J/mol/K). are coefficients of species j provided by thermodynamic databases.

With a given amount of feed, solving Equation (1) would provide equilibrium composition at the outlet of the reactor. The duty, rate of heat exchange, of the reactor is then calculated based on the difference between total enthalpies at inlet and outlet. Note that enthalpy is a function of temperature, therefore inlet temperature is used to evaluate enthalpy at the inlet while equilibrium reactor temperature is used to evaluate enthalpy at the outlet.

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Heat from reactors and heat recovery steam generators (HRSGs) downstream of reactors can be used for steam generation. In all cases, the feed is based on biogas where the flow rate is 6.45 kg/hr CH4 and 11.9 kg/hr CO2. The composition is 60% CH4 and 40% based on volume. The higher heating value (HHV) of the biogas is 100 kW. The feed is assumed to be at 25 °C and 1.2 bar. In a commercial process, the feed would be preheated (for example, by exhaust gases) however, the equilibrium composition would not be effected.

The steam system is based on Rankine cycle between 10 bar and 0.1 bar. Water is first pumped to 10 bar and sent to HRSGs and then reactors to be superheated at 200 °C. The superheated steam passes through a turbine (isentropic efficiency = 55%) to generate power. The exhaust gas is condensed at 0.1 bar and pumped back. A steam table is used to evaluate thermodynamic properties in various locations of the steam system (e.g., the inlet of evaporator, turbine, condenser and pump). The duty of each unit is then calculated as the difference in enthalpies at inlet and outlet. The steam flow rate is determined based on total heat available for steam generation (kW) and ΔH in steam generation (kJ/kg). It is assumed that the exhaust gases can be cooled down to 150 °C and heat curves (temperature vs. heat flow curves) are used to guarantee positive driving force in temperature.

2.2 Design Cases 2.2.1 Three Reactor Design (Biogas Diluted with H2O)

Table 1: Stream Properties of Three Reactor Design (Biogas diluted with H2O)

Air Reactor

Feed Air Reactor

Exhaust Fuel Reactor

Feed Fuel Reactor

Outlet Temperature (°C) 25 800 25 800

Pressure (bar) 1.2 1.2 1.2 1.2 Flow (kg/hr) 155 129.269 32.35 58.081

CH4 6.45 O2 36.1 10.366 N2 118.9 118.897

H2O 28.486 CO 0.0003 CO2 11.9 29.594 NO 0.0057

H2O(L) 14 Cu(cr) 76.878 76.878

Cu2O(cr) 57.011 57.011 CuO(cr) 159.619 159.619

Given that the biogas flow rate is very small as compared to the air flow rate, three or more reactors are necessary to match reactor exhaust flow rates if recycle is not used. In the three reactor design, two are air reactors and one is fuel reactor. The air flow is split evenly into two streams. The biogas has a very low flow rate, and it may be diluted with H2O or CO2. The feed

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flow rates are designed so that the exhaust gases have roughly the same volumetric flow rate among all three reactors. The results for the three reactor design (biogas is diluted with H2O) are shown in Table 1 and the duties of each unit in the steam system are shown in Table 2. The thermal to power efficiency is about 10.8% on HHV basis for this design.

Table 2: Heat Duties and Steam Flow Rate in Three Reactor Design (Biogas Diluted with H2O)

Air Reactor (kW) 28.7 Fuel Reactor (kW) 1.3 Air Reactor Exhaust (kW) 25.9 Fuel Reactor Exhaust (kW) 17.4 Evaporator (kW) 73.3 Turbine (kW) 10.8 Condenser (kW) 62.5 Pump (kW) 0.04 Steam flow rate (kg/hr) 100.1

2.2.2 Three Reactor Design (Biogas Diluted with CO2) The results for the three reactor design (biogas is diluted with CO2) are shown in Table 3 and the duties of each unit in the steam system are shown in Table 4. The thermal to power efficiency is about 12.2% on HHV basis for this design. The efficiency is a little higher because when biogas is diluted with H2O, the heat required to vaporize water may not be recovered.

Table 3: Stream Properties of Three Reactor Design (Biogas Diluted with CO2)

Air Reactor Feed

Air Reactor Exhaust

Fuel Reactor Feed

Fuel Reactor Outlet

Temperature (°C)* 25 800 25 800 Pressure (bar) 1.2 1.2 1.2 1.2 Flow (kg/hr) 155 129.269 40.65 66.381

CH4 6.45 O2 36.1 10.366 N2 118.9 118.897

H2O 14.486 CO 0.0005 CO2 34.2 51.894 NO 0.0057

Cu(cr) 76.877 76.877 Cu2O(cr) 57.012 57.012 CuO(cr) 159.619 159.619

*The temperature in the table is for gas only. The temperature of solids is 800 °C.

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Table 4 Heat Duties and Steam Flow Rate in Three Reactor Design (Biogas Diluted with CO2)


2.2.3 Effect of Al2O3 on System Performance Al2O3 is a commonly used support material for Cu particle. It has been reported that CuAl2O4 and Cu2Al2O4 may be formed during cycling. To study the effect of Al2O3, the feed rates and operating temperatures are maintained the same as before, however, the solid particles are based on a mixture copper on alumina (the percentage of Cu is 46% on weight basis in all cases). If the Cu flow rate is the same as the before (i.e. 128 kg/hr), it is shown that the fuel reactor becomes endothermic at 800 °C and the heat generated by the reactors are below 10 kW in both cases when the biogas is diluted by CO2 (Figure 1) or H2O (Figure 2).

It is observed that the methane is not fully oxidized as lots of CO is produced (Figure 3 and Figure 4). Gradually increasing the Cu flow rate (the Cu/Al2O3 is maintained at the same) would help reduce the amount of CO generated and improve heat generation. It is shown that when the Cu flow rate is tripled, the reactors using Cu/Al2O3 would be equivalent to those using Cu alone. Since the solid particles are in the reactors (they are based on kg) and the gases are flowing (they are based on kg/hr). Increasing the Cu flow rate by three times implies valve switch time interval is only 1/3 of the previous rate.

2.2.4 Two Reactor with Recycle Design To maintain roughly the same exhaust flow rate in fuel and air reactors in the two-reactor design, recycling part of the exhaust gas from the fuel reactor may be used. The stream properties and duties using a recycle ratio of 73.5% are shown in Tables 5-7. The thermal to power efficiency is 12.3%. This design is subject to availability of a high-temperature blower to recycle part of the gas back to the fuel reactor. The efficiency may be lowered due to power required to drive the blower.

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Figure 1: Affect of Catalyst Amount on Reactor Duties in the Presence of Al2O3 (Biogas Feed Diluted with CO2)

Figure 2: Affect of Catalyst Amount on Reactor Duties in the Presence of Al2O3 (Biogas Feed Diluted with H2O)

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Figure 3: Affect of Catalyst Amount on CO Formation in the Presence of Al2O3 (Biogas Feed Diluted With H2O)

Figure 4: Affect of Catalyst Amount on CO Formation in the Presence of Al2O3 (Fuel Diluted With

H2O)

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2.2.5 Sulfur Removal Biogas typically contains H2S which may become SO2 after combustion. Sulfur removal can be done before, during, and after combustion. For example, CaO can be mixed with oxygen carrier in order to capture sulfur during combustion.

Table 8 shows the equilibrium composition at the reactor exhaust when H2S is present in the biogas feed. Nearly all H2S is converted to CaSO4 while a trace amount exits the system when the CaO/H2S is set to 1/0.0685. CaSO4 is fairly stable and does not decompose to release SO2 based on the thermodynamic analysis. However, it is noted that some CaO is converted to CaCO3 in the air reactor, which implies CaO makeup may be required during operation.

Table 5: Stream Properties in Two Reactors with Recycle Design

Air

Reactor Feed

Air Reacto

r Exhau

st

Fuel Reactor Feed

Fuel Reacto

r Recycl

e

Fuel Reacto

r Outlet

Fuel Reacto

r Exhau

st

HRSG

Inlet

HRSG

Outlet

Temperature (°C)* 25 800 25 800 800 800 800 150

Pressure (bar) 1.2 1.2 1.2 1.2 1.2 1.2 1.2 1.2

Flow (kg/hr) 155 129.26

9 18.35 116.21

2 160.29

3 44.081 173.3

5 173.35

CH4 6.45

O2 36.1 10.366

4 10.36

6 10.366

N2 118.9 118.89

7 118.9 118.89

7

H2O 38.191

4 52.677

7 14.486

4 14.48

6 14.486

4

CO 0.0008 0.0011 0.0003 0.000

3 0.0003

CO2 11.9 78.020 107.61

4 29.593 29.59

4 29.594

NO 0.0057 0.005

7 0.0057 Cu(cr) 76.878 76.878

Cu2O(cr) 57.010 57.010

6

CuO(cr) 159.61

9 159.61

9 *The temperature in the table is for gas only. The temperature of solids is 800 °C.

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Table 6: Duties of Process Units in Two Reactors with Recycle Design


Table 7: Stream Properties of Steam System in Two Reactors with Recycle Design

Turbine Inlet Condenser Inlet Pump Inlet Evaporator Inlet Flow (kg/sec) 113.9 113.9 113.9 113.9

Temperature (°C) 200 45.8 45.7 45.8 Pressure (bar) 10 0.1 0.1 10

Quality 1 0.9394 0 0 Density (kg/m3) 4.8543 0.0726 989.886 990.2645

Table 8: Equilibrium Calculation of Biogas CLC with CaO at 800 ºC and 1.2 bar

Component Fuel Reactor Air Reactor In (kg/h) Out (kg/h) In (kg/h) Out (kg/h)

CH4 6.45 0 0 0 O2 0 0 126.4 100.5229 N2 0 0 418.6 418.5844

H2O 14 28.5226 0 0 CO2 11.79 28.7878 0 0.6965 NO 0 0 0 0.0334 H2S 0.0685 0 0 0 SO2 0 0.0003 0 0 Cu 0 77.9007 77.9007 0

Cu2O 0 55.859 55.859 0 CuO 159.619 0 0 159.619 CaO 1 0 0 0.8875

CaCO3 0 1.5841 1.5841 0 CaSO4 0 0.273 0.273 0.273

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2.2.6 Final Prototype Design For the final prototype design, it was determined that the system would be designed to allow for different gas flow rates for the fuel reactor and the air reactor. Each reactor will spend equal amounts of time in either the fuel mode or the air mode. The fuel reactor gas flow rate was set by the heat rate needed for the system, 100kW thermal input. The air flow rate was set to allow for enough oxygen to flow through the reactor and regenerate the bed while maintaining equal time in the air mode and the fuel mode.

Table 9: Equilibrium Calculation of Biogas CLC with CaO at 800 ºC and 1.2 bar

Air Reactor Feed

Air Reactor Exhaust

Fuel Reactor Feed

Fuel Reactor Outlet

Fuel Reactor Exhaust

HRSG Inlet

HRSG Outlet

Temperature (°C) 25 800 25 800 800 800 150

Pressure (bar) 1.2 1.2 1.2 1.2 1.2 1.2 1.2Gas flow (kg/hr) 138.41 112.6794 18.24 43.9706 43.9706 156.65 156.65CH4 0 0 6.45 0 0 0 0O2 32.25 6.5171 0 0 0 6.5171 6.5171N2 106.16 106.158 0 0 0 106.158 106.158H2O 0 0 0 14.4864 14.4864 14.4864 14.4864CO 0 0 0 0.0003 0.0003 0.0003 0.0003CO2 0 0 11.79 29.4839 29.4839 29.4839 29.4839NO 0 0.0043 0 0 0 0.0043 0.0043H2S 0 0 0 0 0 0 0SO2 0 0 0 0 0 0 0H2O(L) 0 0 0 0 0Cu(cr) 76.878 0 0 76.878 0Cu2O(cr) 57.0106 0 0 57.0106 0CuO(cr) 0 159.6192 159.6192 0 0CaO(cr) 0 0 0 0 0CaCO3(cr) 0 0 0 0 0CaSO4(I,II) 0 0 0 0 0Cu2S(a,b,c) 0 0 0 0 0CuS(cr) 0 0 0 0 0Al2O3(alpha) 0 0 0 0 0CuAl2O4(s) 0 0 0 0 0Cu2Al2O4(s) 0 0 0 0 0

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Table 10: Duties of Process Units in Final Two Reactor Design

Air Reactor (kW) 32.6 Fuel Reactor (kW) 17.1 Air Reactor Exhaust (kW) 22.6 Fuel Reactor Exhaust (kW) 11.8 Evaporator (kW) 84.0 Turbine (kW)* 12.4 Condenser (kW) 71.6 Pump (kW) 0.04 Steam flow rate (kg/hr) 114.8

*assuming 55% isentropic efficiency

2.3 Experimental Validation of Thermodynamic Design A thermodynamic analysis of experimental data was done based on the following assumptions:

• 160-gram bed of particles, 30% wt. is CuO. No formation of CuAl2O4 during cycles. • Fuel Reactor (FR) mode: 1 lpm CH4, 4 lpm Ar, 52.25 seconds (based on an average of all

cycles). • Air Reactor (AR) mode: 10 lpm Air (20% O2, 80% N2), 61.4 seconds (based on an average

of all cycles). • Reactor temperature is 730 °C.

It is shown that under the assumption of thermodynamic equilibrium, a cyclic steady-state can be reached, and the particles are 0.1556 mol Cu2O and 0.2923 mol CuO at the end of the FR mode and 0.6034 mol CuO at the end of the AR mode. Or, 0.078 mol O2 can be transferred between AR and FR modes. The fuel can be completely oxidized to CO2. The duties of AR and FR (after heating the feed to 730 °C) are calculated to be 10.9 kJ and 6.9 kJ. The heat will be absorbed by particles and the reactor wall etc. If the heat is just used to heat the particles (based on 160 grams of particles and a heat capacity of 800 J/gram/°C), the bed temperature increases by 86oC in AR mode and 54oC in FR mode. It is also found that the duties and gas compositions are the same even if the CuO in the particles is only 20wt%, i.e., as long as the particles are in excess, its amount would not change the results. The flow rates and temperatures of data collected on 2/22/2016 are shown in Figure 5.

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Figure 5: Flow Rates and Temperatures of Data Collected on 2/22/16

To calculate energy balance, the following data are used:

8CuO+CH4=4Cu2O+CO2+2H2O, ∆HRX = -2.71x105 J/mol at 700 oC 2Cu2O+O2=4CuO, ∆HRX = -2.65x105 J/mol at 700 oC CH4+2O2=CO2+2H2O, ∆HRX = -8.01x105 J/mol at 700 oC The heat required to heat gas from inlet temperature (assumed to be 50 °C) to reactor temperature (assumed to be 730 °C) are 38 kJ/mol for CH4, 14.1 kJ/mol for Argon, 21.1 kJ/mol for Air.

It is shown that the heat generated is more than enough to heat the gas from inlet to reactor temperature (see the following table). Note that the average net heat generation of 18 kJ per FR/AR cycle is consistent with the result obtained in the thermodynamic code. This net heat generation should lead to increases in bed temperature from cycle to cycle.

However, except the first cycle, the temperature increase is minimal (a few degrees C). In the first cycle, the heat tape might be on for part of the time.

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Table 11: Per Cycle Temperature Change and Heat Release

cycle1 cycle2 cycle3 cycle4 cycle5 cycle6 Average Argon cooling after Air mode (sec) 37.6 8 5 5 6 6 11.3 Fuel mode (sec) 73.2 57 83 51 35 34 55.5 Argon cooling after Fuel mode (sec) 16.3 14 8 8 7 6 9.9 Air mode (sec) 76.8 67 67 61 53 50 62.5 Heat released in the combined FR/AR cycle (based on CH4 + 2O2 = CO2 + 2H2O), J 4.36E4 3.39E4 4.94E4 3.04E4 2.08E4 2.02E4 3.31E4 Heating of gases from 50C to 730C, J 2.00E4 1.57E4 1.70E4 1.38E4 1.15E4 1.09E4 1.48E4 Net heat absorbed by bed or reactor wall, J 2.35E4 1.82E4 3.24E4 1.65E4 9.37E3 9.37E3 1.82E4

The CH4 conversion is close to 100% based on the emission measurement (=1-CxHy/(CxHy+CO2+CO) at the outlet). Therefore, the energy imbalance may be due to imperfect insulation of the reactor.

It is found that the temperature drop is roughly linear at the end of the AR step and also the argon purge step (Table).

The theoretical temperature drop rate due to air cooling near the end of the AR cycle (based on 10 lpm air, 125g particle and 800 J/g/K heat capacity) is dT/dt = -10/60/22.4*21.1e3/ (0.125*800) = -1.56 C/sec, very close to the experimental observations.

The theoretical temperature drop due to argon cooling (based on 4 lpm argon) is dT/dt = -4/60/22.4*14.1e3/ (0.125*800) = -0.42 C/sec. The dropping rate is more in measurement.

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Table 12: End of Cycle Temperature Drop

Cycle Temp Time

∆T/∆t (air cooling)

∆T/∆t (argon cooling)

1 Peak T in AR 753.3 980.2

End of AR Step 725.3 999 -1.49

End of Argon purge 718 1007

-0.91

2 Peak T in AR 750.2 1131

End of AR Step 730.5 1145 -1.41

End of Argon purge 724.9 1150

-1.12

3 Peak T in AR 749.8 1292

End of AR Step 732.2 1308 -1.10


-0.86

4 Peak T in AR 750 1420

End of AR Step 735.1 1433 -1.15


-1.15

5 Peak T in AR 743.8 1516

End of AR Step 717.4 1534 -1.47


-0.97

6 Peak T in AR 740.8 1612

End of AR Step 713.8 1630 -1.50

A possible explanation based on heat loss is as follows. When reaction produces heat, part of the heat is lost due to imperfect insulation. As a result, the temperature increase is lower than expected. Moreover, the net heat generation is not only absorbed by the bed, but also by the reactor wall. When air cools the reactor (near the end of the AR step), on the one hand, heat is lost due to imperfect insulation. On the other hand, the wall heats up the gas. The two effects roughly cancel each other. In Argon purging step, the heat loss due to imperfect insulation dominates, and therefore the temperature drop rate is higher than theoretical calculation (Figure 6).

Overall, the thermodynamic code shows a good correlation with the experimental data.

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Figure 6: Experimental Gas Flows and Emissions

Time (sec)

400 600 800 1000 1200 1400 1600 1800

Flow

(lpm

)

-2

0

2

4

6

8

10

12

CH4

Ar

Air

400 600 800 1000 1200 1400 1600 1800

Out

let c

ompo

sitio

n,%

0

2

4

6

8

10

12

14

16

18

20

CxHy

CO2

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CHAPTER 3: Economic and Life Cycle Analysis 3.1 Commercial Scale Life Cycle Analysis 3.1.1 Life Cycle Benefits The primary benefit of ZERE is the ability to capture CO2 and thereby to enable natural gas combustion to achieve very low emissions, as measured by grams of CO2 per kWh of net electricity generation. This makes biogas combustion to be the energy conversion step in a system that has negative carbon emissions. The negative emission case arises in an energy system consisting of CO2 capture from the atmosphere by a living plant, conversion to biogas, ZERE performing conversion of the biogas to heat and steam, the CO2 out of the ZERE step being separated from the steam-plus-CO2 output stream, and the CO2 being sequestered, i.e., locked in the solid earth or deep ocean and thereby geologically isolated from the atmosphere.

Other benefits may also be achieved by ZERE: (1) better economics at small scale, compared to alternatives that capture CO2; (2) lower sulfur (SO2) and NOx emissions, again at small scale where air emission controls for SO2 and NOx from biogas combustion may be too expensive. When compared to conventional burning of natural gas, ZERE may not have any meaningful advantage over natural gas, assuming that the natural gas is already very low in sulfur. However, the capture of sulfur in the ZERE reactor particle bed can give ZERE an advantage over other biogas combustion systems. On NOx emissions, the air reactor temperature is low enough to keep NO and NO2 formation lower than in natural gas combustion. In lab testing, NOx emissions were consistently 0-1ppm with occasional measurements of 1-3ppm.

ZERE also considered that the fuel nitrogen (primarily as NH3) in biogas may form NOx. To address this concern, a thermodynamic analysis was performed considering 400ppm NH3 in the inlet fuel. At equilibrium, all of the NH3 is converted to N2 and H2O in the fuel reactor. These calculated results are promising for achieving very low NOx emissions. Very little data has been found on NH3 concentration in biogas and no lab testing was performed on NH3 in the fuel feed. Confirmation of the calculated zero NOx result will be a topic for future testing.

For comparison of benefits, the ZERE system performance at a scale that will generate on the order of 5 MW electric power will be compared to systems of about the same size that consume natural gas or biogas as the fuel. The performance of the parts of the total system upstream and downstream of the ZERE-with-CO2-capture subsystem will be derived from the GREET model, using the “California GREET” whenever possible. This is the CA-GREET model developed by and used by the California Air Resources Board (CARB or simply ARB) for use in the ARB’s implementation of low carbon fuel standards (LCFS).

The CA-GREET model is based on GREET, the model developed and revised over the past decade or more at the Argonne National Laboratory, a model specifically designed to do life cycle analysis of energy technologies and systems. Results of cases give energy, carbon, SO2, NOx, and several other pollutant and waste discharge flows and emission/effluents/wastes. . . .

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3.1.2 Final ZERE Case with Steam Power Generation Parameters and results for electricity generation and carbon dioxide emission from a steam power generation system that uses steam generated by heat transfer from the hot gases out of the ZERE unit are given below under Parameters. Key parameters as to size and efficiency are as:

• 1000 scfm (standard cubic feet per minute) input flow of the methane (CH4) in biogas that is 2/3 CH4 and 1/3 CO2, by volume

• 21.8% net thermal efficiency in converting the energy in this methane input (assumed 1000 Btu/scf) into net electrical output (after 15% of gross electric power assumed to be used for auxiliary power inside the ZERE-based power plant)

• 15,649 Btu/kWh net heat rate from biogas input to net electric output • 3.83 MW net electric output (4.51 MW electric gross)

This ZERE case uses these efficiency values that lead to the 21.8% overall net efficiency:

• 90% efficiency energy content of the biogas (CH4 fraction at 1000 Btu/scf) to hot gas output

• 30% gross thermal (1st Law) efficiency steam-to-electricity • 95% generator efficiency • 85% ratio of net to gross electric generation (15% auxiliary power demand) • Overall efficiency only 21.8% biogas-to-gross-electric-output (15, 651 Btu/kWh)

The size assumed here is appropriate for a very large dairy biogas power generation unit: 3.8 MW net electric from an input of 1000 scfm (standard cubic feet per minute) flow of biogas into the ZERE unit with the biogas being 67% by volume methane (CH4) and 33% carbon dioxide (CO2).

3.1.3 Emissions For estimating emissions other than CO2, i.e., sulfur and nitrogen gases and particulates, the precursors of these will be considered as minor additions to the gas streams assumed here, moving with these streams of major components methane, carbon dioxide and water (liquid and vapor).

The direct CO2 emission factor for this base case is 36.6 grams of CO2 per kWh (net). This assumes that 97% of the CO2 is captured in the C-capture unit downstream of the ZERE unit. The primary advantage of the ZERE technology is its capability to enable efficient and low-cost capture of CO2. When used with fuel containing biomass-derived carbon, ZERE makes possible negative carbon emission—the moving of carbon dioxide back from the atmosphere to some form of permanent storage underground or in the deep ocean, removing this carbon from the atmosphere-biosphere carbon cycle.

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3.1.4 Parameters of the Final ZERE Case For a small commercial plant of 1000 SCFM input flow of methane and an estimated net electric power of 3.83 MW and a gross electric power of 4.51 MW (after 95% generator eff.), the following calculation applies:

1. ZERE unit (biogas to hot gases in output):

90% heat yield

1,000 scfm of CH4 input

(2/3 of the input volume, and all of the energy)

1000 Btu/scf

60 MMBtu/h input

54 MMBtu/h out as hot gases

2. Power generation unit (hot gases into work)

30% gross efficiency

54 MMBtu/h input

16.2 MMBtu/h as power out

3.4124 MMBtu/MWh conversion factor.

4.75 MWh/h (MW) gross work output

3. Condensing out the H2O from the hot gases

70% waste heat fraction of "2" above

37.8 MMBtu/h into condenser

4. CO2 capture process and resulting CO2 emitted

1,000

scfm CO2 (1 to 1 with CH4 input, because CH4 + 4O CO2 + 2H2O)

for biogas input, add the CO2 in the biogas

51.89 grams/scf for CO2

3,113 kg/h CO2 from the CH4 in the biogas

add for the input biogas as 2/3 CH4

and only 1/3 CO2: hence add 50% more

1,557

kg/h CO2 from the

CO2 in input biogas

4,670 kg/h total CO2

out of ZERE process

% control =====> 97% fraction of CO2 captured by

condensing out the water

140 kg/h CO2 (3%)

4.75 MW gross output (work)

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95% assumed eff of generator

4.51

MW gross electric (MWh/hour)

31.06

kg/MWh emission factor (for gross MWh elec. output

This is 31 grams CO2 per kWh gross elec. But, if no capture the emission factor (gross) is

4,670 kg/h

4670kg/4.51MWh= 1035 kg/MWh (gross)

Based on net generation, and with 97% capture, this gives a carbon emission factor as follows:

15% of gross used for aux. power

0.68 MW auxiliary power

3.83 MW net elec. (3830 kW)

140 kg/h emission of CO2

36.55 grams CO2 per net kWh

3.1.5 Cases in the Preliminary LCA Report The preliminary report on Life Cycle Analysis (LCA) used the California GREET model of 2009 to obtain numbers for two parts of the analysis: (1) the indirect emission of fossil carbon CO2 due to processes that bring the biomass into and through the biogas production facility which is dairy with an anaerobic digestion (AD) system making what CA GREET calls “digester gas” and this project is referred to as biogas; and (2) a negative emission factor to credit the biogas plant for use of a non-fossil source, namely, the renewable biomass fuel derived from plant matter that took CO2 out of the atmosphere as it grew (and, for a dairy, before it became feed for the cows).

The second of these numbers, that is the negative emission factor, is used in GREET to give credit for the renewable biomass origin of what becomes biogas and then electric power. The 2009 CA GREET number for this is -52.265 kg CO2 per MMBtu of biogas produced in the AD process. In the ZERE cases, the biogas enters the ZERE unit where it is converted to hot gases (CO2 and H2O) that go out of the ZERE unit into a power and carbon capture system that generates electricity at 30% gross efficiency and then captures 97% of the CO2 (actually, 99% in the preliminary LCA). In the present analysis (final LCA), at least so far, no credit has been given for the non-fossil nature of the sources of the carbon in the biogas fuel going into the ZERE system. At the conclusion of this final LCA, the credit, that is the negative emission factor, is given only for the CO2 that is captured by the ZERE system. This limited use of the credit is because the comparison is between biogas systems only. If a natural gas system is added to the comparison, then some credit must be given to all biogas cases to credit the renewable, no-fossil-carbon property that biomass offers in contrast to natural gas which is a fossil fuel.

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The first of the two numbers taken out of the 2009 CA GREET result is an emission factor for the fossil carbon emissions that are due to processes before or during the AD process. This emission factor is 1109 grams CO2 per MMBtu of biogas. The energy input, from fossil sources, to gather and process the biomass to biogas is given by the 2009 CA GREET as 22,209 Btu per million Btu of biogas product.

On Page 4 of the preliminary report on LCA (January 29, 2014), the following text and table were presented:

The CA GREET model for this case gives these results for Digester Gas (DG) which is the same as biogas:

Process Upstream of

Energy CO2 equiv. CO2 equiv.

the Power Plant (Btu/MMBtu) per MJ per MMBtu

Digester Gas Recovery

22,209 1.17 1,109 Credit for Biomass Source -1,000,000 -55.14 -52,265

----------- ----------- -----------

Total (per MMBtu or MJ of DG) -977,791 -53.97 -51,156

The energy column in the above table shows the total energy input (direct use of fossil fuel and indirect) input required in two “processes” upstream of the power plant. The second line “Credit” is for the process of taking CO2 out of the atmosphere as the biomass grows as a living plant. The first line “Digester Gas Recovery” is for the process in the dairy that produces the biogas. (The biogas is called “digester gas” or “DG” above and in the GREET model analysis, and the process step is called “recovery” of the gas, a term adopted for a previous GREET analysis of landfill gas, also done for the ARB earlier in 2009.) The biogas (or DG) becomes the fuel input to the ZERE or Conventional plant that uses the 1,000,000 Btu of biogas fuel to produce useful energy.

The systems evaluated via that table in the preliminary LCA, namely ZERE and the “Conventional” alternative, are assumed to be systems that use no fossil fuel in the plant itself and generate their own electricity for internal use. Purchased fuel and electricity from the grid to start and restart the plant are assumed to be only very small amounts and are neglected here. Given these assumptions, the only fossil carbon emission is that given by the 2009 CA GREET model for “digester gas” (dairy AD) systems. This gives 1109 grams of CO2 per MMBtu of biogas produced and is the total of fossil CO2 via upstream and indirect use of fossil fuels.

3.1.5 Final ZERE Case For this final LCA and with the parameters shown above, the 1000 scfm methane input in a biogas stream gives 4670 kg/hour as the CO2 emission, if there is no CO2 capture. Note that 1/3 of this 4670 total is from the CO2 that is part of the input along with the CH4. With the assumed 97% capture the emission is cut to only 140 kg/h. For the 4.51 MW gross electric output the resulting emission factors are 1035 grams CO2 per kWh, if no CO2 capture, and 31 g/kWh with the 97% capture. For the assumed 15% demand for auxiliary power, the net electric output is

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3.83 MW and the emission factor with 97% control becomes 36.5 gCO2/kWh (net kWh), as shown.

In the approach taken here in this revised analysis, the credit for taking CO2 out of the atmosphere is to be taken at the exit of the power plant where CO2 is captured and then disposed of in a way that keeps it out of the atmosphere, i.e., “sequestered” as in CCS (carbon capture and sequestration). With biogas made from renewable biomass sources being the only fuel used in both the ZERE and the Conventional systems addressed here, the only fossil carbon emissions are the indirect emissions due to processes upstream of the energy plant. If external electric power or external fossil fuel such as natural gas, were to be used in any of these biogas-fueled plants, then some fossil carbon emissions would be associated with such external sources of electricity or fuel.

The Life Cycle Analysis (LCA) introduces consideration of emissions not at the power plant itself but, rather, emitted upstream of the power plant in the processes that were necessary to make the biogas and/or to bring the biomass material to the site where the biogas is produced via anaerobic digestion (AD). On Page 4 of the preliminary report on LCA, dated January 29, 2014, the following text and table were presented:

The CA GREET model for this case gives these results for Digester Gas (DG) which is the same as what this project refers to as biogas:

Process Upstream of

Energy CO2 equiv. CO2 equiv.

the Power Plant (Btu/MMBtu) per MJ per MMBtu

Digester Gas Recovery

22,209 1.17 1,109 Credit for Biomass Source -1,000,000 -55.14 -52,265

----------- ----------- -----------

Total (per MMBtu or MJ of DG) -977,791 -53.97 -51,156

The energy column in the above table shows the total energy input (direct use of fossil fuel and indirect) input required in two “processes” upstream of the power plant. The second line “Credit” is for the process of taking CO2 out of the atmosphere as the biomass grows as a living plant. The first line “Digester Gas Recovery” is for the process in the dairy that produces the biogas. (The biogas is called “digester gas” or “DG” above and in the GREET model analysis, and the process step is called “recovery” of the gas, a term adopted for a previous GREET analysis of landfill gas, also done for the ARB earlier in 2009.) The biogas (or DG) becomes the fuel input to the ZERE or Conventional plant that uses the 1,000,000 Btu of biogas fuel to produce useful energy.

The systems evaluated here, namely ZERE and the “Conventional” alternative, use no fossil fuel in the plant and generate their own electricity for internal use. Purchased fuel and electricity from the grid to start and restart the plant are assumed to be only very small amounts and are neglected here. Given these assumptions, the only fossil carbon emissions are those given by the 2009 CA GREET model for “digester gas” (dairy AD) systems: 1109 grams of CO2 per MMBtu of biogas produced.

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3.1.6 Three Biogas Cases For the three biogas cases discussed – namely, the ZERE case having the parameters detailed on above, and the ZERE and Conventional cases of the preliminary LCA report – the fossil CO2 emission factors are derived from the above 1109 grams fossil CO2 per MMBtu of biogas going out of the AD process at a dairy waste digester and into the ZERE or Conventional power system.

3.1.6.1 Results for Life Cycle Fossil Emissions. Fossil CO2 emission factors in the three biogas (“digester gas”) cases are as follows: (1) Preliminary Conventional Case, 11.3 grams/kWh; (2) Preliminary ZERE Case, 15.1 g/kWh; and (3) Final ZERE Case, 17.4 g/kWh. Note that the CO2 emissions here are the fossil CO2 only, not adding in the renewable biomass CO2. These fossil CO2 emissions are entirely the upstream and indirect as given in the 2009 CA GREET, 1109 g/MMBtu, with the MMBtu being the heat content in the biogas fuel entering the power system. Parameters of the three biogas-fueled power systems are summarized as:

Conventional Case:

5.1 MW net

11,331 Btu/kWh net HR 30.1% overall thermal eff.

1109 gCO2e/MMBtu

0.010201 MMBtu/kWh

11.3 gCO2/kWh (net)

ZERE Prelim. Case:

5.1 MW net


1109 gCO2e/MMBtu

0.013648 MMBtu/kWh

15.1 gCO2/kWh (net)

ZERE Final Case:

3.8 MW net


1109 gCO2e/MMBtu 90% biogas-to-hot-gases

0.015651 MMBtu/kWh 30% gross work output

17.4 gCO2/kWh (net) 95% generator eff.

Product of all 4 at right:

85% gross-to-net elec.

21.8% overall thermal eff

Therefore, the CA GREET life cycle analysis adds 17.4 grams of CO2 emission to the 36.5 g/kWh calculated for the 97% removal case (“Final ZERE”). However, the 36.5 g/kWh is CO2 from renewable biomass and the 17.4 g/kWh is fossil CO2 and the two types of CO2 are not the same for purposes of life cycle analysis of global warming gas emissions.

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3.1.6.2 Renewable biomass emission factors. The emission factor above is 36.55 gCO2/ kWh and this is for the “Final ZERE” case having the parameters previously discussed. This is for the direct emission at the plant and with 97% CO2 control via carbon capture by condensing the H2O vapor out of the CO2/H2O hot gas stream out of the ZERE unit. The equivalent emission factor for the ZERE and Conventional cases in the preliminary LCA were given there as 11.9 gCO2/kWh for ZERE and 1099 gCO2/kWh for Conventional. (Efficiency parameters involved are in the numbers shown immediately above and are taken from Page 1 of the preliminary LCA report.) The ZERE case in the preliminary LCA assumed 99% carbon capture and the ZERE case here assumes 97%. Therefore, the 11.9 g/kWh would be 35.7 g/kWh if the same 97% were assumed, and this is essentially the same as the Final ZERE case, both at 36 g/kWh. (All of the emission factors in this paragraph and the above paragraph net kWh basis, not gross kWh.)

3.1.7 Summary on CO2 Emissions The conventional and ZERE biogas-to-power plants analyzed in the preliminary report were both sized at approximately 5 MW net electric power in size and had CO2 emission factors of 1099 and 1190 grams of CO2 per kWh (net) if carbon capture is not taken into account. The Conventional plant has lower emission because it is more efficient – 11,331 Btu/kWh gross heat rate of the Conventional versus 13,649 for ZERE – due to the extra auxiliary power needed by the ZERE system. (See Pages 1 & 2 of the preliminary report.) With the 99% CO2 capture assumed for the ZERE plant, its emission factor was cut by a factor of 100 to only 11.9 gCO2/kWh. In both cases no credit (i.e., reduced emission) is taken into account to allow for fossil CO2 emission avoided by use of renewable biomass as the fuel source, which is a major item. Also not taken into account was the smaller correction for the natural gas combustion avoided by displacing natural gas use due to the heat produced in the CHP (combined heat and power) systems adopted for the preliminary report, both the “Conventional” and the ZERE systems.

In the case added here in the present final report, the ZERE system is of comparable size but less efficient. The gross thermal efficiency is 27% giving a gross heat rate of 13,304 Btu/kWh in this 4.5 MW (gross electric) biogas-to-power plant. To compare to the preliminary cases where the net efficiency, heat rate and emission factor were given, 15% auxiliary power use is assumed and the results are 22% net efficiency, 15,650 heat rate and 36 grams CO2 per kWh emission.

When the life cycle results for fossil carbon emissions, indirect and upstream of the ZERE system, are taken from the 2009 CA GREET “digester gas” case (dairy biogas), 15.1 grams of fossil CO2 per net kWh are added to the emission factor for the Preliminary ZERE case. However, only this 15.1 g/kWh is the emission of fossil CO2 as all of the rest, the 35.7 g/kWh are from renewable biomass, not fossil.

As just noted, when the zero-fossil-carbon property of the biomass source of dairy biogas is taken into account, only this 15.1 g/kWh remains as an addition of fossil CO2 to the atmosphere. This 15.1 g/kWh is for the ZERE case in the preliminary report. Due to the lower efficiency of the Final ZERE case added here, the 4.5 MW gross becomes 3.8 MW net (15% auxiliary power) and the 15.1 g/kWh emission for the preliminary ZERE case becomes 17.4 grams of fossil CO2

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per net kWh. The larger emission factor is due to the larger amount of energy (MMBtu) biogas input needed to make a net kWh of electricity. The relevant comparisons here are the net heat rates: 11,331 Btu/kWh “Conventional” and 13,649 for the preliminary ZERE biogas case, both of these at 5 MW net, and the 15,651 for this final ZERE biogas case at only 3.8 MW net.

3.1.8 Conclusion on CO2 Emissions The most important consideration for ZERE in terms of Life Cycle Analysis (LCA) is the potential for lower cost capture of CO2 due to the simple capture mechanism provided by condensing the water vapor out of the gas stream that exits from the ZERE unit. This becomes a negative emission of fossil carbon because, with subsequent disposal (sequestration) of the captured carbon, CO2 has been taken out of the atmosphere. In fact, the lower energy efficiency of the ZERE system means that more biomass is used to generate a net kWh and, therefore, more CO2 is removed from the atmosphere. However, less electricity generation hurts the economics as there is less revenue or savings to offset the cost of the carbon capture.

The estimate here for this LCA result on fossil carbon emissions, independent of the economics just mentioned, is as follows: The negative emission factor for a ZERE power or CHP plant is determined primarily by the amount of CO2 captured. For the case described here this is the 97% of the 4670 kg/h of CO2 that is captured: 4530 grams CO2 per hour. For the 3.83 MW net electric generation this gives 1183 g/kWh (4530 grams/h over 3830 kW net). This is for a plant that is sized at 1000 scfm (cubic feet per minute at standard conditions: 1 atm and 70 degrees F) of methane (CH4) input in a “biogas” input stream that is 2/3 by volume CH4 and 1/3 by volume CO2. In electric MW size this plant is 4.51 MW in gross electric power output and 3.83 MW net.

In life cycle analysis (LCA), the indirect fossil CO2 emissions from processes upstream of the ZERE system are taken into account and reduce the negative emission factor. Per net kWh of electricity this indirect upstream emission is 1109 grams CO2 per MMBtu of biogas produced and at the net heat rate of the ZERE power system this is 17.4 grams CO2 per net kWh (1109 g/MMBtu times 0.015651 MMBtu/ kWh) is 17.4 g/kWh). Therefore, the negative emission factor for the ZERE power system is -1166 grams CO2 per kWh: -1183 [direct, renewable biomass] + 17 [indirect fossil fuels] = -1166 g/kWh.

3.1.9 Criteria Pollutant Emissions (NOx, SO2, PM, CO)

Life cycle assessment is not needed for addressing the normal criteria emissions. This is because the upstream direct and indirect emissions for NOx, SO2, Particulate Matter and CO are the same for both the ZERE and the Conventional cases being compared. What matters is the comparison of direct emissions at the power plant. Data and analysis to address these emissions are being done in other tasks in this project. In general, it can be said that application of ZERE’s technology would generate electric power and heat, with near zero emissions that beat CARB 2007 standards. In ZERE’s technology, fuel is oxidized internally using a solid stage oxygen carrier, capturing nearly 100% of the fuel energy and forming near zero emissions. Fuel oxidation and overall system efficiency are not affected by inert constituents such as carbon dioxide in the biogas and contaminants such as hydrogen sulfide and silanes are trapped in the

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reactor system bed. Regenerating the spent oxygen carrier in air takes place at a low temperature forming no NOx or other pollutants.

3.2 Commercial System Economic Benefit Analysis Adding to the preliminary economics report, this final economic analysis has developed a steam cycle case at 5 MW for ZERE and a comparison case also at 5-MW size. The ZERE case differs from before in assuming an indirectly heated steam cycle rather than a cycle in which the hot gases are used directly at high pressure. The revised case is less efficient in electric power generation due to loss of exergy, i.e., capacity to do work, as high temperature heat (700-800 C) in the hot gases out of the reactors becomes lower temperature heat (250 C) in the steam out of the heat exchanger generating the steam that is the working fluid in the steam power cycle. The simplification of the ZERE system by using atmospheric rather than pressurized reactors is the benefit, and this type of system is expected to be the best path forward for early deployment, especially for use with dairy digester systems at sizes almost certain to be below 1 MW electric.

Key parameters of this new ZERE are shown on the left in Table 13 and the revised comparison Conventional case is on the right. Both are biogas cases at 5-MW size, and the cost of electricity is higher by 3.36 cents/kWh for the ZERE system with CHP compared to the Conventional case, which also is a CHP system. Both cases are for the power system downstream of a biogas production facility, such as a dairy with a lagoon digester, and in both cases the power system is not charged for fuel, which is biogas, nor is the power system charged for the capital expense of the digester system. (The costs of the digester system are considered as being paid out of a separate budget that the dairy pays as a cost of doing business, normal expenses of controlling odor, solid waste and liquid waste.) The right hand side of Table 13 shows performance parameters that would lead to the net and gross electric power efficiencies shown in the left hand side. Finally, at the bottom of Table 13 resulting cost of electricity and its components are given. Appendix A shows the line-by-line details of the calculation.

The bottom line differences between ZERE and Conventional in Table 13 are as follows: (1) ZERE electricity alone, no heat byproduct credit, is 8.69 cents/kWh, then, taking credit for heat recovery for other uses at $5.00/ MMBtu, a 1.64 cent credit, gives a net cost of electricity as 7.04 cents/kWh; (2) Conventional has electricity alone at 6.63 cents/kWh, then the heat credit at 2.95 leads to a net cost of electricity of 3.68 cents/kWh. Hence, the extra cost of the ZERE case is 3.36 cents/kWh (7.04 – 3.68). Given that the avoided carbon dioxide via the capture of CO2 by a ZERE system gives a net avoided emission advantage versus Conventional of 955 grams CO2 per kWh (0.955 metric ton of CO2 per MWh). This gives a cost of 0.003513 cents/gram or $35.13 per metric ton CO2

This $35/ton is a reasonable price for avoiding fossil CO2, reasonable compared to estimates of the cost of various options investigated in studies of economics of climate change and fossil emissions. However, it is high compared to the current price of fossil carbon offsets and allowances in the “cap-and-trade” system that exists in California under the AB32 (2006) Climate Change Solutions Act, where the current price is $12 or $13 per metric ton CO2. If the two major items of higher cost for the ZERE system were reduced to the same as Conventional,

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i.e., capital cost of $2500/kW cut to $2000/kW and O&M cost of 3.5 cents/kWh cut to 2.5 cents/kWh, then the extra cost of ZERE would be 1.3 cents/kWh instead of 3.3 cents/kWh and the cost of CO2 emission avoided would be close to the current cap-and-trade price range. Therefore, a target for the ZERE system is $2000/kW installed capital cost and 2.5 cents/kWh total of fixed plus variable operating cost.

Table 13: ZERE vs. Conventional Biogas-to-Power CHP

Input assumptions: Units ZERE Conventional Size in kW (net electric) kW elec 5108 5108 Aux. Power (fraction of gross elec.) 0.20 0.10 Gross efficiency (generator out. / fuel energy in.) % 30.0% 28.5% Net eff. (elec. net over fuel energy input)

24.0% 25.7%

Capital cost in $/kW (net kW elec.) $/kW 2500 2000 Percent/year assumed for capital recovery %/y 15% 15% Cents/kWh for operating costs cents/kWh 3.5 2.5 Annual capacity factor % 0.85 0.85 Percent of "waste heat" that displaces nat. gas % 33.0% 62.0% Percent net CHP efficiency (net elec. + net heat) % 47.1% 70.0% Dollars/MMBtu value of heat sales/use $/MMBtu $5.00 $5.00 Energy density in fuel (~65% CH4 by vol.) Btu/scf 650 650

Cost of electricity results: Units ZERE Conventional MWh/year net electricity generation MWh/year 38,034 38,034 Capital Recovery (15%/y per above) cents/kWh 5.04 4.03 O&M (labor, maint., chemicals, etc.) " 3.65 2.60 Fuel (biogas from digester) " 0 0 Credit for CHP heat (% and $ per above) " -1.64 -2.95

--------------- ---------------

TOTAL net cost of net electricity cents/kWh 7.04 3.68

Efficiency values in the above: Units ZERE Conventional Generator: electric / shaft work 0.95 0.95 Gross elec first law (thermal) eff. % 30.0% 28.5% Net elec. / gross elec. % 80.0% 90.0% Net electric efficiency % 24.0% 25.7% Auxiliary Power: Net electric output - kW elec. kW elec. 5,108 5,108 Gross electric kW elec. kW elec. 6,385 5,676 Auxiliary electric - kW kW elec. 1,277 568

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Breakdown of aux power: Compression power - kW kW elec. 709 0 Balance of plant - kW kW elec. 568 568

Net fossil CO2 emission & cost/ton:

Percent CO2 captured by ZERE % 97% None Net CO2/kWh from upstream per GREET*** gCO2/kWh -698 -580

LCA grams CO2 net emission g/kWh gCO2/kWh -661 294 Avoided CO2 via ZERE vs. conventional " 955 n.a.

Extra cost of ZERE system cents/kWh 3.36 n.a. Cost of CO2 avoided by ZERE vs. conv. cents/gCO2 0.003513 n.a.

Cost above converted to $/metric ton CO2 $/MTco2 $35.13 n.a.

*** The credit for organic rather than fossil CO2 is in this number from the GREET model.

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CHAPTER 4: Computational Fluid Dynamics 4.1 CFD Hydrodynamic Model of Semi Continuous Reactors 4.1.1 Introduction Fluidized beds are widely used in the process industries including chemical-looping combustion (CLC) process where the oxygen carrier serves as the bed medium. A consensus has been established to use a bubbling fluidized bed for the fuel reactor and a circulating fluidized bed for the air reactor (Kolbitsch et al. 2009; Kolbitsch, Proll, and Hofbauer 2009). Despite their widespread application in the process industries the design of fluidized bed reactors is still very challenging. Complex gas–solid hydrodynamics inherent to these reactors is closely coupled to heat transfer and reaction kinetics. Because of this intimate coupling leading to a highly non-linear system, the use of empirical models for scale up is challenging.

First-principles based computational fluid dynamics (CFD) has become an emerging and effective tool to explore the complex hydrodynamics behavior in gas-solid fluidized bed. CFD offers the advantage that it can provide more insight into the physical origin underlying the various phenomena transpiring in fluidized beds, and can be used for scale-up, design, or optimization (Gidaspow 2004; Gidaspow, Bezburuah, and Ding 1991; Kuipers 1998). Different approaches have been taken in early attempts to apply CFD methods including direct numerical simulation (DNS), discrete particle method (DPM) and two-fluid method (TFM) to explore the phenomena prevailing in gas-fluidized beds (van der Hoef 2008). DNS and DPM are limited to a relatively small scale application due to their high computational costs. Amongst these methods the TFM based on adaptations of the kinetic theory of gases is computationally less expensive. The general idea in formulating the TFM model is to treat each phase as an interpenetrating continuum, and therefore to construct integral balances of continuity, momentum and energy for both phases, with appropriate boundary conditions and leap conditions for phase interfaces. TFM apply averaging techniques and assumptions to obtain momentum balance for the solids phase(s) since the resultant continuum approximation for the solid phase has no equation of state and lacks variables such as viscosity and normal stress (Pain 2001). The TFM equations are coupled with constitutive relations derived from data or analysis of nearly homogeneous systems (Gidaspow 1994). The interphase drag force between gas and solid phases is modeled by various empirical correlations reported in the literature, including those of Syamlal and O’ Brien (1989) (Syamlal 1989), Gidaspow (1994) (Gidaspow 1994), and Wen and Yu (1966) (Wen 1966).

Traditionally, gas-solids friction coefficients have been expressed using semi-empirical correlations such as the well-known Wen-Yu and Ergun correlations (Wen 1966), but recently, similar expressions were fitted to simulation data obtained using physical models based on first principles. These recently derived drag models were obtained by finely resolving the fluid flow around the particles, and the friction or drag can be obtained by integrating the fluid viscous stress acting on the particles according to Newton’s law of viscosity. Such models are known as filtered drag correlations (Andrews IV, Loezos, and Sundaresan 2005; Benyahia 2009; Igci et al. 2008; Milioli et al. 2013).

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In this study the TFM was used to simulate the prototype fluidized bed of air and fuel reactor of ZERE with the filtered-drag model proposed by (Milioli et al. 2013).

4.1.2 ZERE Prototype Reactors Configuration Two fluidized bed reactors will operate in parallel. At any moment in time, one operates in air mode, and one operates in fuel mode, and after 10 minutes in operation the reactors switch their operation mode. So, the operational cycle time in each mode (define cycle time) is 10 minutes. In either case, the reactor temperature is nominally 800 °C. Each reactor will be filled with 228 kg of particles with diameter 300 μm, made with 30% copper loaded on porous alumina support. Both reactors have identical configuration, and are fabricated from 18-inch (ID = 0.42 m) Schedule 40 steel pipe.

These reactors are nominally designed to operate in a bubbling regime, with 𝑈𝑈0𝑢𝑢𝑚𝑚𝑚𝑚

≈ 36 when

the reactor is in air mode, and 𝑈𝑈𝑢𝑢𝑚𝑚𝑚𝑚

≈ 9 when the reactor is in fuel mode. Here U0 and umf are the

inlet and the minimum fluidization gas velocities respectively. The inside of the reactor will be baffled to break-up bubbles, thereby ensuring a high conversion of methane in the fuel during fuel mode, and reducing particle entrainment during oxidation mode. Two baffles are inserted in each bed to break apart bubbles, which will increase the gas-solid contact by promoting mass transfer, and, thereby increase fuel conversion. The baffle is made from a 3/8-inch steel plate with numerous 2.0-cm holes through it, on a square pattern and center-to-center spacing is 3.1 cm. The baffle is designed to keep about 30% of the baffle cross sectional area open to gas flow.

At the distributor plate, there are 12 tuyeres for flow of fuel gas during fuel mode and air during air mode. In this novel design, gas flows through the four tuyeres nearest the center of the distributor during fuel mode, and is distributed to all 12 tuyeres in air mode.

4.1.3 Simulated Reactors Configuration The majority of numerical simulations of pilot or large scale fluidized beds are carried out with a two-dimensional flow assumption in which a cut-plane along the axis of the cylindrical column is used. Combinations of the aforementioned simplifications can be found in many simulations, for example 2D cold flow simulations of riser flow in FCC process (Li et al. 2014). In our 2-D model, the volumetric flow rate of inlet gas is taken to be a constant value equal to the average of inlet and outlet volumetric flow rate. For example, the volumetric flow rate of inlet fuel gas for the 100 kW fuel reactors is 0.016𝑚𝑚3

𝑠𝑠 but after oxidation of the fuel, the volume

flow rate increases to 0.036𝑚𝑚3

𝑠𝑠. In simulation the average volumetric flow rate of 0.026𝑚𝑚3

𝑠𝑠 was

used as a constant flow rate throughout the reactor during fuel mode. Similarly, the average volumetric flow rate for air reactor of 0.097𝑚𝑚3

𝑠𝑠 is used.

The baffles were modeled as impermeable surfaces where gas and solids can flow through 2-cm openings from the holes. About 33% of the baffle cross sectional area is opened to gas flow. There are seven 2-cm openings on each baffle.

ZERE gas distributors are specially designed for the fluidized bed. However, in the fuel mode of prototype design, fuel gas flows only through the center 4 tuyeres. Gas flowing through the

38

tuyeres will be at a high velocity, prior to decelerating in the bed. To accommodate that in the simulation, a single central jet is introduced in the fuel reactor by maintaining inlet velocity with the aforementioned jet velocity emitting from the tuyeres. For instance, gas flows through either the four fuel tuyeres having 32 jets of 5.0 𝑚𝑚𝑚𝑚 size. All of these jets are altogether considered as a single central jet for the fuel reactor simulation. The (single) jet opening in the simulation is calculated as 32 × 5 𝑚𝑚𝑚𝑚 = 160 𝑚𝑚𝑚𝑚 for fuel mode.

4.1.3.1 Simulation Setup Geometry and Discretization The dimensions of the simulated beds are shown in Table 14. Two dimensional (2D) Cartesian coordinates were used with uniform structured grid cells. There were a 63 × 487 (𝑎𝑎𝑎𝑎𝑎𝑎𝑎𝑎𝑎𝑎 ×𝑣𝑣𝑣𝑣𝑣𝑣𝑣𝑣𝑎𝑎𝑣𝑣𝑎𝑎𝑎𝑎) mesh resolution, corresponding to a computational mesh with 𝜕𝜕𝑎𝑎 = 𝜕𝜕𝜕𝜕 = 6.7 𝑚𝑚𝑚𝑚 in the 2-D geometry. From our past experience and literature studies, the research team believes this grid resolution will give grid independent results.

Numerical Model The interpenetrating two-fluid model (TFM) (Ishii and Hibiki 2010; Ishii and Mishima 1984; Boure, Bergles, and Tong 1973; Drew 1983; Gidaspow 1994) based on the Eulerian-Eulerian flow field was applied to simulate the gas-solid hydrodynamics of fluidized bed. This approach has been confirmed to give adequate representations of the hydrodynamics of fluidized bed units (Lun et al. 1984; Chalermsinsuwan, Piumsomboon, and Gidaspow 2009; Gidaspow, Bezburuah, and Ding 1991). To consider the effect of unresolved sub-grid scale heterogeneous structures on the inter-phase drag force, the filtered drag correlation proposed by Milioli et al. was used which is obtained from very fine grid simulation (Milioli et al. 2013).

Flow Solver and Solver Settings The National Energy Technology Laboratory’s (NETL’s) computational fluid dynamics (CFD) open-source code MFIX (Syamlal 1993) was used as flow solver. The modified phase coupled SIMPLE scheme, which uses a solids volume fraction correction equation instead of a solids pressure correction equation, was used for pressure–velocity coupling. The second-order higher accuracy SuperBee schemes were used for the spatial discretization of all remaining equations. A combination of point successive under relaxation and biconjugate gradient stabilized method (BiCGSTAB) method were used for the linear equation solver. A maximum residual at convergence of 10−03 was used to improve the accuracy of the continuity and momentum equations solution. First order implicit temporal discretization was used to ensure stable and accurate solutions. It has been shown that 2nd order time discretization is necessary for accurate solution of fast-moving riser flows with the TFM (Cloete, Amini, and Johansen 2011), but this is not the case for dense bubbling beds where the vast majority of the bed moves very slowly. An automatic time-step adjustment with a maximum and minimum time-step of 4 × 10−03 s and 10−07 s respectively was specified.

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Table 14: Physical Properties of Simulation Parameters

Designed Prototype Simulated reactors

Air reactor

Fuel reactor

Particle diameter, µm 300~500 300 300

Particle density, kg/m3 1964 1964 1964

Gas density, kg/m3 0.31~0.34 0.34 0.31

Gas viscosity, kg/m/s (37 ~44) x 10-6 44 x 10-6 37 x 10-6

Gas velocity, cm/s 11~87 67.5 18

Minimum fluidization velocity, cm/s 2.40 ~ 2.80 2.40 2.80

Maximum packing limit 0.58 0.58

Particle–particle restitution 0.90 0.90

Angle of internal friction 30° 30°

Bed diameter, m 0.42 0.42 0.42

Static bed height, m 1.60 1.60 1.60

1st baffle position, m 0.70 0.70 0.70

2nd baffle position, m 1.40 1.40 1.40

Reactor height, m 3.26 3.26 3.26

Temperature, K 1073 1073 1073

Initial and Boundary Conditions The standard initial conditions were used to describe the 2D simulations. Initially, bed height was 1.6 𝑚𝑚 with 50% void fraction. Initial bed pressure drop was 15 𝑘𝑘𝑘𝑘𝑎𝑎, which is the hydrostatic bed pressure. It was assumed that initially bed was under the minimum fluidization condition with the minimum fluidization velocities of 2.4 𝑐𝑐𝑚𝑚

𝑠𝑠 for air and 2.8 𝑐𝑐𝑚𝑚

𝑠𝑠 for fuel reactors,

respectively.

Boundary conditions (BC) were specified over flow planes/2D surfaces that are normal to one of the coordinate directions and coincide with a face of the scalar control-volume. A constant gas flow rate was specified at the distributor of the fluidized bed as inlet boundary and a constant pressure at the top of the domain was used as outflow boundary. The wall boundaries were specified as partial slip with no-slip for gas and free-slip for solid phase. As indicated above, baffles were specified as internal surfaces, and specified as impermeable for gas and solids. Internal surfaces acted as free-slip walls in stress computations.

Simulation Summary A summary of the physical properties and simulation parameters are given in Table 14.

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4.1.4 Results and discussions 4.1.4.1 Bubble Size and Frequency The hydrodynamics of a fluidized bed have a primary influence on bed characteristics such as solid and gas mixing, heat transfer to immersed surfaces and elutriation of particles from the bed. For beds operating in the bubbling regime, the bed hydrodynamics are largely governed by the number, size and motion of bubbles passing through the bed and erupting on the surface. Contours of the void fraction observed in the 2D fluidized bed at one instant are shown in Figures 7-9. The shapes of the bubbles are far from the spherical or ellipsoidal forms observed in small particle beds. However, similar bubble shapes were observed in the two-dimensional bed (Glicksman, Lord, and Sakagami 1987).

The influence of complex flow structures on reactor performance is complicated. Internal baffles may be introduced to modify the gas-solid flow structures, in an effort to form a more uniform and active gas-solid flow to enhance heat and mass transfer so as to improve the overall performance of fluidized bed reactors, especially to facilitate scale-up. Figure 7 and Figure 8 illustrate the effect of internal baffles on bed hydrodynamics, especially on bubbles, clusters, and non-uniform flow structures.

Figure 7: Effect of Baffles on Bubble Break-Up in the Air Reactor Contour of Gas Volume Fraction at T= 6.0 S with Uniform Inlet Gas Velocity = 0.675 M/S

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When a bed does not contain internals of any sort, the movement of bubbles in the bed is unrestricted. As bubbles rise, they gradually increase in size and tend to move horizontally toward the center of the bed. Much of the gas flow is ‘short-circuited’ through the bubbles, which greatly limits interaction between particles and gaseous reactants, and thus impacts the conversion and selectivity of a chemical reaction, especially for the Group B particles at high superficial gas velocities. It shows that the maximum bubble diameter appears to be as wide as the bed diameter, a phenomenon called “slugging”. Such large bubbles would violently shake the unit as tons of oxygen carrier splashed when they come out of the bed surface. In addition, mass transfer from such an enormous bubble would be so poor that it would significantly reduce the rate of reaction (Jin, Wei, and Wang 2003).

Figure 8: Effect of Baffles on Bubble Break-Up in the Fuel Reactor Contour of Gas Volume Fraction at T= 6.0 S with Uniform Inlet Gas Velocity = 0.18 M/S

42

The performance of a fluidized bed reactor can be improved by decreasing the bubble size and renewing the bubbles surface for interchanging the gas between bubbles and the interstitial gas in the emulsion phase. Table 15 shows quantitatively the effect of baffle insertion in the fluidized bed on bubble break-up (sample calculation has shown in the appendix). In all cases shown, baffles are effectively breaking large bubbles and reducing the size except in the case of jet flow effect. The number of bubbles nearly doubles in the baffled bed case. However, when a high velocity central jet flow of gas is considered for the fuel reactor instead of a uniform inlet gas velocity, bubble size is even larger than for the unbaffled bed (Figure 9). It is noted that the average bubble size in the air reactor also appears to be doubled as compared to the fuel reactor. This is due to the high inlet gas velocity of the air reactor.

Table 15: Summary of Bubbles

Air reactor Fuel reactor

W/O baffle

W/ baffle

W/O baffle

W/ baffle

Jet W/ baffle

Number of bubbles 23.0 41.0 16.0 34.0 23.0

Total area of bubbles, cm2 2820.42 3352.16 509.96 797.46 691.53

Average diameter of bubbles, cm 8.59 7.90 4.41 4.25 5.07

Number of bubbles having area > 0.5 cm2

19 40 13 32 21

Total area of bubbles (area > 0.5 cm2), cm2

2820.16 3352.09 509.70 797.16 690.79

Average diameter of bubbles (area > 0.5 cm2), cm

10.35 8.09 5.37 4.50 5.49

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Figure 9: Effect of Inlet Gas Velocity Condition in the Fuel Reactor. Contour of Gas Volume Fraction at T= 6.0 S. In Both Cases, 33% Open Area in the Baffle Hole

The net effect of incorporating baffles is to reduce average bubble size by an average of about 22% in the case of the air reactor. Furthermore, there is a 16% reduction of bubble size in case of fuel reactor with uniform fuel gas velocity. For the fuel reactor, comparing the scenario of uniform gas distribution to the scenario of gas distribution in a single jet, both baffled, it was found that the effect of a single jet is to cause an increase in bubble size. This is not surprising, since bubble size in the bed depends significantly on the size of bubbles formed at the distributor, which in turn is determined by the gas velocity at the distributor.

4.1.4.2 Bed Expansion Understanding the bed expansion characteristics of a bubbling fluidized bed is crucial for several reasons and most importantly; bed expansion is used to design a reactor. In reactive

44

fluidized-bed reactor systems, the information on mass of solids per unit bed volume (the bed density of fluidized bed) is important because this influences the chemical conversion calculation (Geldart 2004). When the heat transfer is to be calculated, the bed expansion gives the bed voidage, which is necessary to predict the heat transfer coefficient, and the bed height, which defines the heat transfer surface.

As the demarcation of bed surface is nearly impossible to identify in a vigorously bubbling fluidized bed, either in the case of a simulation or an actual bubbling fluidized bed. Thus, reporting the bed height at any moment in time is not straightforward. The gas volume fraction distribution must be post-processed in some manner to determine the bed height. A method suggested in the literature was adopted, in which the bed height is said to be the height below which 90% of the bed weight is found (Syamlal and O'Brien 2003).

The instantaneous area-averaged axial solid volume fraction, 𝜺𝜺𝒔𝒔� (𝒉𝒉, 𝒕𝒕), is calculated as a function of the bed height.

𝜺𝜺𝒔𝒔� (𝒉𝒉, 𝒕𝒕) = 𝟏𝟏 − ∫ ∫ 𝜺𝜺𝒈𝒈(𝒙𝒙,𝒚𝒚,𝒕𝒕)𝒅𝒅𝒚𝒚 𝒅𝒅𝒙𝒙𝒉𝒉

𝟎𝟎 𝑾𝑾𝟎𝟎

∫ ∫ 𝒅𝒅𝒚𝒚 𝒅𝒅𝒙𝒙𝑯𝑯𝟎𝟎 𝑾𝑾

𝟎𝟎 (2)

Where, W and H are the reactor dimension in x and y direction respectively. h is the vertical coordinate above the gas distributor. The instantaneous bed mass can be expressed as:

𝝌𝝌(𝒉𝒉, 𝒕𝒕) = 𝑨𝑨 ⋅ 𝝆𝝆𝒔𝒔 ∫ 𝜺𝜺𝒔𝒔� (𝒉𝒉, 𝒕𝒕)𝒅𝒅𝒉𝒉𝒉𝒉𝟎𝟎 (3)

Where A is the cross-sectional area of the bed and 𝜌𝜌𝑠𝑠 is the particle density.

In all cases, initially the bed expands with time until it levels off at a quasi-steady bed height. After the initial expansion and collapse, the bed surface fluctuation is more orderly over time. It has been noticed that baffled fluidized bed expands more than that of unbaffled bed. The baffles cause larger bubbles to break apart and produce smaller bubbles which rise through the bed more slowly. Thus small bubbles reside in the bed a longer time, and cause the bed to expand relative to the effect of larger bubbles.

Comparing Figures 10 and 11, air reactor expansion is much higher than fuel reactor. This large expansion results from the higher superficial gas velocity at the gas distributor. The ratio of the inlet to the minimum fluidization gas velocity ( 𝑈𝑈0

𝑢𝑢𝑚𝑚𝑚𝑚≈ 28) for air reactor falls in the vigorously

bubbling regime however in the fuel reactor case ( 𝑈𝑈0𝑢𝑢𝑚𝑚𝑚𝑚

≈ 7) it falls in the gently bubbling regime

(Daizo and Levenspiel 1991). Another fact to note from Figure 13 is that the initial time to reach a quasi-steady state for an unbaffled bed is twice as long as the baffled bed.

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Figure 10: Baffle Effect on Bed Expansion in the Air Reactor (Uniform Inlet Gas Velocity = 0.675 M/S)

Figure 11: Baffle Effect on Bed Expansion in the Fuel Reactor (Uniform Inlet Gas Velocity = 0.18 M/S)

It has been expected that introducing a central jet instead of uniform gas distribution at the fuel reactor will cause the bed to expand, however Figure 12 shows that both models for the gas distribution produce similar results for bed height. This implies that the uniform gas distribution assumption for fuel is safe enough to apply in the simulation for the purpose of finding bed height.

0.0

1.0

2.0

3.0

4.0

0 1 2 3 4 5 6 7 8 9 10

Expa

nded

hei

ght,

m

time, s

Air reactor with baffle Air reactor without baffle

0.0

0.4

0.8

1.2

1.6

2.0

0.0 1.0 2.0 3.0 4.0 5.0 6.0 7.0 8.0 9.0 10.0

Expa

nded

bed

hei

ght,

m

time, s

With baffle Without baffle

46

Figure 12: Baffle Effect on Bed Expansion in the Fuel Reactor (Uniform Inlet Gas Velocity = 0.18 M/S and Jet Velocity = 0.475 M/S)

4.1.4.3 Average bed expansion After a quasi-steady state is achieved, the solids distribution in the bed is time-averaged to get the simulated bed height. The height determined this way is normalized by comparison to the value corresponding to 90% of the initial bed height (0.90*1.60 m). Table 16 shows the average bed expansion for air and fuel reactors with and without baffles. The data is averaged from 3.0 -10.0 s of simulation time since beds reach, in all cases, quasi-steady state approximately after 3.s. As discussed in the previous section, a baffled bed creates more bubbles and those bubbles are small in sizes which causes the beds to expand more than the unbaffled bed condition.

In the case of operation in air mode with baffles, bed expansion is significant; the bed height is predicted to be 2.45 m, and it is expected that additional solids are present above this nominal bed height. The real value of the baffles comes from the reduced bubble size, which will cause better interaction between air and the particles, promoting faster particle oxidation. Also, because the bubbles erupting on the bed surface are relatively smaller, it is expected that the rate of particle entrainment from the bed to the cyclones is somewhat reduced (discussed below). Finally, the reduced pressure fluctuations (discussed below) mean that the system is likely to require less maintenance, since large pressure fluctuations can be very hard on both upstream and downstream gas processing equipment.

0.0

0.4

0.8

1.2

1.6

2.0

0.0 1.0 2.0 3.0 4.0 5.0 6.0 7.0 8.0 9.0 10.0

Expa

nded

bed

hei

ght,

m

time, s

Single jet at the center Uniform inlet gas velocity

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Table 16: Summary of Average Bed Expansion

Air reactor Fuel reactor

W/O baffle W/ baffle W/O baffle W/ baffle Jet

Expanded height, m 2.10 2.45 1.47 1.53 1.50

% of expansion 46.16 70.45 2.28 5.97 3.94

Expansion ratio, H/H0 1.46 1.70 1.02 1.06 1.04

4.1.4.4 Bed Pressure Drop Pressure fluctuation data obtained from fluidized beds are a rich source of information on the hydrodynamic states of these systems (Brown and Brue 2001) (Brue and Brown 2001)). The resulting time series data can be analyzed by a number of different methods, including standard deviation, probability density functions, autocorrelation analysis, and power spectral density (PSD) analysis (Yates and Simons 1994). One of the most common pressure fluctuation analyses is standard deviation. It has often been used to identify different regimes in fluidized beds, where a maximum value with respect to inlet gas velocity is associated with the transition from a bubbling to turbulent fluidization regime. Standard deviation has also been used to determine minimum fluidization velocity (Sobrino et al. 2008) and to detect the onset of defluidization in operating fluidized beds (van Ommen, de Korte, and van den Bleek 2004).

Figures 13-15 show the pressure fluctuation for air and fuel reactors with and without baffles. The initial hydrostatic bed pressure drop is 15 𝑘𝑘𝑘𝑘𝑎𝑎 (∆𝑘𝑘 = (1 − 𝜀𝜀0) 𝐻𝐻0 𝜌𝜌𝑠𝑠 𝑔𝑔). The average pressure drop shown in all cases here closely represents the bed hydrostatic pressure. All the pressure drop data considered here for analysis is 1.0 ~10.0 s of simulation period; the simulation at times before 1.0 s is excluded, since no bubbles have erupted yet on the bed surface, creating a relatively smooth pressure trace. As previously mentioned, standard deviation has often been used to identify different fluidization regimes in fluidized beds. Comparing the standard deviation of air and fuel reactors when the pressure drops, it is understandable that air reactors shows vigorously bubbling characteristics and fuel reactors are moderately bubbling.

The difference in standard deviation between baffled and unbaffled reactors highlights the effect of baffles, especially in air mode. Bubbles grow without restriction in an open bed, causing higher pressure drop fluctuation (Figure 13). However, such deviation is not present in the case of fuel reactors (Figure 14).

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Figure 13: Baffle Effect on Bed Pressure Drop Fluctuation in the Air Reactor (Uniform Inlet Gas Velocity = 0.675 M/S)

49

Figure 14: Baffle Effect on Bed Pressure Drop Fluctuation in the Fuel Reactor (Uniform Inlet Gas Velocity = 0.18 M/S)

50

Figure 15: Baffle Effect on Bed Pressure Drop Fluctuation in the Fuel Reactor (Uniform inlet gas velocity = 0.18 m/s and jet velocity = 0.475 m/s)

As the fuel reactor with central jet flow of gas operates at a velocity higher than uniform gas velocity condition, the standard deviation also depicts that effect with slightly higher values than uniform one. As previously shown, bubble size is a little bigger in jet flow than uniform flow.

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4.1.4.5 Particle Entrainment In the bubbling zone of a fluidized bed, bubbles grow by coalescence and rise to the surface of the bed where they break. As bubbles break at the surface of the bed, particles are thrown up in the freeboard zone and are entrained by the upward flowing gas stream. In this zone, some particles are carried far above the bed surface and are elutriated while others fall back to the bed. The freeboard zone usually affords an opportunity for the disengagement of particles and for the lean phase reactions. During the operation of a fluidized bed, a large amount of fine particles could be elutriated continuously.

In order to examine the elutriation from the air reactor, solid flux is calculated at the gas outlet located at the top of the domain. Solid volume fraction, 𝜺𝜺𝒔𝒔(𝒙𝒙, 𝒕𝒕), at the exit of the reactor height, H is defined as

𝜺𝜺𝒔𝒔(𝒙𝒙, 𝒕𝒕) = 𝟏𝟏 − 𝜺𝜺𝒈𝒈(𝒙𝒙, 𝒕𝒕) (4)

Solid particle velocity has two components, but the x-component of the velocity has no effect on particle entrainment from the reactor. Y-component solid particle velocity, 𝑽𝑽𝒔𝒔(𝒙𝒙,𝑯𝑯, 𝒕𝒕) at the exit of the reactor height, H can easily be obtained from simulation. Average solid flux at the reactor exit is defined as

𝝓𝝓𝒔𝒔(𝒕𝒕) = 𝝆𝝆𝒔𝒔 ∫ 𝜺𝜺𝒔𝒔(𝒙𝒙,𝑯𝑯,𝒕𝒕) 𝑽𝑽𝒔𝒔(𝒙𝒙,𝑯𝑯,𝒕𝒕)𝒅𝒅𝒙𝒙 𝑾𝑾𝟎𝟎

∫ 𝒅𝒅𝒙𝒙 𝑾𝑾𝟎𝟎

(5)

Figure 16 shows the solid flux time profile for air reactor with and without baffles. During the initial unsteady period, solid elutriation is much higher in unbaffled air reactor than baffled one, even though the bed surface is lower. This is due to the unrestricted growth of bubbles in unbaffled reactor. However, excluding the unsteady period, bed shows no elutriation in either of the cases. No elutriation is predicted for operation during fuel mode, with or without baffles.

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Figure 16: Baffle Effect on Solid Flux in the Air Reactor (uniform inlet gas velocity = 0.675 m/s)

4.1.5 Conclusions CFD simulations of bubbling fluidized bed of air and fuel reactors with and without internal horizontal baffles have been presented in this study. It shows that internal baffles are effectively breaking large bubbles. This will enhance the interchange of gas between the bubbles and the emulsion phase. Chemical reactions and mass transfer can be improved when bubbles

53

are small and evenly distributed throughout the bed volume. However, experimental data is required for further validation of the simulation results.

4.2 Model Reaction Kinetics in the Semi Continuous Reactors The principal aim of the current study is to develop a reactor model which best fits the experimental data and could be reliable for scale up and development. The existing models on bubbling fluidized bed reactors suggest that the hydrodynamic model has to include the effect of bubbles and associate mixing. This requires appropriate selection of the hydrodynamic equations to describe the bed behavior. The developed model describes the behavior of a bubbling fluidized bed to process a fuel gas with semi-continuous chemical looping combustion system. The hypotheses considered for the model are: steady state, isotherm bed at macroscopic level, no existence of particle fragmentation or attrition, and no elutriation. The model is one dimensional, based on empirical and semi-empirical expressions and takes into account lateral exchange of gas in the bottom bed between bubbles and emulsion. The detail of the modeling of the fluidized-bed reactor is described in the following sections. The experimental setup is as described in Chapter 5.

4.2.1 Fluid Dynamics of the Model In bubbling fluidized bed, fluid-dynamics is largely governed by bubble rise velocity. An understanding of the bubble rise velocity requires the knowledge about the bubble size. The bubble size changes along the length of the solid bed due to the increased gas volume fraction of gas after fuel reaction with oxygen carrier particles. To include this volume expansion, simultaneous solution of hydrodynamics and material balances are important. The changes in the gas velocity, U, and corresponding bubble size, dB, are incorporated through the following equations.

𝒅𝒅𝑩𝑩(𝒛𝒛) = 𝟎𝟎.𝟓𝟓𝟓𝟓𝒈𝒈𝟎𝟎.𝟐𝟐 �𝑼𝑼(𝒛𝒛)− 𝒖𝒖𝒎𝒎𝒎𝒎�

𝟎𝟎.𝟓𝟓�𝒛𝒛 + 𝟓𝟓

√𝑵𝑵�𝟎𝟎.𝟖𝟖

(6)

Here, 𝑢𝑢𝑚𝑚𝑚𝑚 is the minimum fluidization velocity which is estimated from Wen and Yu correlation (Daizo and Levenspiel 1991); N is the number of orifice for the gas distributor; and z is the height of the reactor.

The gas expansion changes the superficial gas velocity by:

𝒅𝒅𝑼𝑼𝒅𝒅𝒛𝒛

= 𝟏𝟏𝑨𝑨

𝒅𝒅𝒅𝒅𝒛𝒛�∑ 𝝂𝝂�̇�𝒋

𝑹𝑹𝑹𝑹𝑷𝑷𝒕𝒕𝒋𝒋 � (7)

Here, 𝜗𝜗 is the reaction stoichiometry of methane with CuO; R and T are the ideal gas constant and temperature; Pt is the total pressure; and A is the cross-sectional area of the reactor.

Since the reactor will operate under isothermal condition, the temperature change is neglected. The solid pressure drop variation across the reactor height is expressed by:

𝒅𝒅𝑷𝑷𝒅𝒅𝒛𝒛

= −𝝆𝝆𝑶𝑶𝑶𝑶𝒈𝒈 (𝟏𝟏 − 𝜺𝜺𝒃𝒃(𝒛𝒛))(𝟏𝟏 − 𝜺𝜺𝒎𝒎𝒎𝒎) (8)

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Here, 𝜌𝜌𝑂𝑂𝑂𝑂 is the density of the oxygen carrier; The porosity at the minimum fluidization conditions, 𝜀𝜀𝑚𝑚𝑚𝑚, was obtained using the equation proposed by Grace (Grace 1986).

𝜺𝜺𝒎𝒎𝒎𝒎 = 𝟎𝟎.𝟓𝟓𝟖𝟖𝟔𝟔 𝑨𝑨𝒓𝒓−𝟎𝟎.𝟎𝟎𝟐𝟐𝟎𝟎 � 𝝆𝝆𝒎𝒎𝝆𝝆𝑶𝑶𝑶𝑶

� (9)

Where 𝐴𝐴𝑣𝑣 = 𝑑𝑑𝑝𝑝3𝜌𝜌𝑚𝑚(𝜌𝜌𝑂𝑂𝑂𝑂 −𝜌𝜌𝑚𝑚)𝑔𝑔𝜇𝜇2

. Here, dp is the mean particle diameter; 𝜌𝜌𝑚𝑚and 𝜇𝜇 are methane density

and viscosity. The bubble fraction, 𝜀𝜀𝑏𝑏, in the solid bed is

𝜺𝜺𝒃𝒃(𝒛𝒛) = 𝑼𝑼(𝒛𝒛)−𝒖𝒖𝒎𝒎𝒎𝒎𝒖𝒖𝑩𝑩(𝒛𝒛)

(10)

With the consideration of the fluidized bed column diameter, the bubble rise velocity, 𝑢𝑢𝐵𝐵, was calculated using the equation proposed by Werther (Werther 1984).

𝒖𝒖𝑩𝑩(𝒛𝒛) = 𝝓𝝓𝑩𝑩�𝒈𝒈 𝒅𝒅𝑩𝑩(𝒛𝒛) (11)

Where Φ𝐵𝐵 is taken as 0.64 based on the reactor diameter studied in this model.

4.2.2 Material balances into the reactor Material balances for the reactant and products were developed for the two phases considered in the gas-solid fluidized bed reactor. Plug flow for the gas was considered in all cases. The reactor model was solved from the distributor plate towards the upper part of the reactor, following the differential equations showed below. The pathway for methane reaction with the CuO/Al2O3 oxygen-carrier was considered to happen in:

𝐶𝐶𝐻𝐻4 + 4𝐶𝐶𝑢𝑢𝐶𝐶−→ 𝐶𝐶𝐶𝐶2 + 2𝐻𝐻2𝐶𝐶 + 4𝐶𝐶𝑢𝑢 (12)

Or 𝐶𝐶𝐻𝐻4 + 8𝐶𝐶𝑢𝑢𝐶𝐶−→ 𝐶𝐶𝐶𝐶2 + 2𝐻𝐻2𝐶𝐶 + 4𝐶𝐶𝑢𝑢2𝐶𝐶

In the dense bed, a gas exchange between bubbles and emulsion is considered allowing the exchange of products and reactants between these phases by diffusive and/or bulk flow mechanism. Indeed, as the gas suffers a volumetric expansion during methane reaction, some of the gas in the emulsion must move to the bubble phase to maintain the minimum fluidization condition in the emulsion phase. Considering all the above assumptions, the mass balances were given by the following differential equations for each gas i (CH4, CO2, and H2O) in the emulsion and bubble phases, respectively:

𝒅𝒅𝑵𝑵𝒃𝒃𝒊𝒊̇

𝒅𝒅𝑽𝑽= 𝒅𝒅��𝑼𝑼−𝒖𝒖𝒎𝒎𝒎𝒎�𝑶𝑶𝒃𝒃𝒊𝒊�

𝒅𝒅𝒛𝒛= −𝑲𝑲𝒃𝒃𝒃𝒃𝜺𝜺𝒃𝒃 �𝑶𝑶𝒃𝒃𝒊𝒊 − 𝑶𝑶𝒃𝒃𝒊𝒊� (13)

𝒅𝒅𝑵𝑵𝒃𝒃𝒊𝒊̇

𝒅𝒅𝑽𝑽= 𝒅𝒅{(𝟏𝟏−𝜺𝜺𝒃𝒃)𝒖𝒖𝒎𝒎𝒎𝒎𝑶𝑶𝒃𝒃𝒊𝒊}

𝒅𝒅𝒛𝒛= 𝑲𝑲𝒃𝒃𝒃𝒃𝜺𝜺𝒃𝒃 �𝑶𝑶𝒃𝒃𝒊𝒊 − 𝑶𝑶𝒃𝒃𝒊𝒊� − 𝒌𝒌′′′𝑶𝑶𝒃𝒃𝒊𝒊

𝒏𝒏 𝜺𝜺𝒎𝒎𝒎𝒎(𝟏𝟏 − 𝜺𝜺𝒃𝒃) (14)

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Here, �̇�𝑁 represents the molar flow rates of gas i in the differential volume element dV. The subscripts b and e represent the bubble and emulsion phases, respectively. 𝑘𝑘′′′ is the volumetric rate constants. 𝐶𝐶𝑏𝑏 and 𝐶𝐶𝑒𝑒 are the gas concentrations in the bubble and emulsion phases. Since no reaction is assumed in the bubble phase, gas will be exchanged between the bubble and emulsion phases through the interchange coefficient Kbe. In the emulsion phase, methane gas will be reacting with solid oxygen carrier particles and product gases will be flowing from the emulsion to the bubble. These equations allow determining the concentration of gas i in both phases: emulsion and bubbles.

4.2.3 Kinetic Model for Oxygen Carrier For the heterogeneous gas-solid reaction, a shrinking unreacted-core model is used to describe the reduction of CuO/Al2O3 particles with methane under reactor conditions. In this model, chemical reaction is considered as the main resistance to the global reaction. Early studies showed that external and internal mass-transfer resistances as well as the particle size have no or minimal effect on reaction rate of CuO/Al2O3 particles with methane fuel (Abad et al. 2010; Garcia-Labiano et al. 2004). A plate-like geometry for CuO in the porous surface of the Al2O3 particle with unchanging size and also with no ash layer formation was used for the kinetic model. Considering all of these assumptions, the equations that describe the conversion time, t, for plates is the following (Levenspiel 1999)

𝒕𝒕 =𝝆𝝆𝑶𝑶𝒖𝒖𝑶𝑶

𝑴𝑴𝑾𝑾𝑶𝑶𝒖𝒖𝑶𝑶

𝝑𝝑𝒌𝒌"𝑶𝑶𝒑𝒑𝒏𝒏 �𝟏𝟏 − 𝜹𝜹

𝑳𝑳� (15)

Here, 𝜌𝜌𝑂𝑂𝑢𝑢𝑂𝑂 and 𝑀𝑀𝑊𝑊𝑂𝑂𝑢𝑢𝑂𝑂are the density and molecular weight of copper oxide; L and 𝛿𝛿 are the initial and unreacted thickness of CuO layer in the oxygen carrier particles. 𝐶𝐶𝑝𝑝 is methane concentration around the particles which is equal to the emulsion phase concentration of methane; n is the reaction order, and 𝑘𝑘" is the surface rate constant in terms of Arrhenius equation as

𝒌𝒌" = 𝒌𝒌𝟎𝟎" 𝒃𝒃𝒙𝒙𝒑𝒑 �− 𝑬𝑬𝒂𝒂𝑹𝑹𝑹𝑹� (16)

Here, 𝐸𝐸𝑎𝑎 is the activation energy. The time 𝜏𝜏 required for complete conversion is given when

𝛿𝛿 = 0, or 𝜏𝜏 =𝜌𝜌𝑂𝑂𝐶𝐶𝑂𝑂

𝑀𝑀𝑊𝑊𝑂𝑂𝐶𝐶𝑂𝑂

𝜗𝜗𝑘𝑘"𝑂𝑂𝑝𝑝𝑛𝑛 . At any moment, the unreacted CuO layer thickness, 𝜂𝜂, can be found from

𝜼𝜼 = 𝟏𝟏 −𝝑𝝑𝒌𝒌"𝑶𝑶𝒑𝒑𝒊𝒊

𝒏𝒏

𝝆𝝆𝑶𝑶𝒖𝒖𝑶𝑶𝑴𝑴𝑾𝑾𝑶𝑶𝒖𝒖𝑶𝑶

𝑳𝑳𝒕𝒕 (17)

The thickness of the layer, L, over the Al2O3 support was determined considering the surface area reported by the vendor and the weight fraction of active CuO in the oxygen carrier particles. Using the specific surface area and the density of the oxygen carrier particles,

volumetric rate constant was estimated as 𝑘𝑘′′′ = 𝜗𝜗𝑘𝑘"

𝜀𝜀𝑚𝑚𝑚𝑚�1 − 𝜀𝜀𝑚𝑚𝑚𝑚� 𝐴𝐴𝑆𝑆𝑆𝑆 𝜌𝜌𝑂𝑂𝑂𝑂.

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4.2.4 Break Through Time for Fuel Break-through is defined as when the reactant gas is first appearing at the fluidized bed reactor exit. At the start of a non-circulating fluidized bed, all the solid oxygen carriers will be available for gas-solid reactions. Soon after the reactant gas flow starts, solid particles at the entrance reduces quickly, and as the time progress the conversion of the remaining active solid particles also move forward. This solid reduction process continues and at some point there will not be any active solid particles to react with fuel gas. When most of the solid particles will be in reduced state, part of the reactant gas will leave the reactor as unreacted. The time when a trace amount of reactant gas is noticed at the reactor exit will be referred as the break-through time. For instances, break-through time for methane is the time when a trace amount of methane is detected by the model calculations. The subsequent description provides the details about break-through time calculation.

As described in the material balance, a part of the inlet gas will be in the emulsion phase at minimum fluidization, and any gas excess to minimum fluidization has to move through the bubble phase. Since there is no reaction in the bubble phase, all the solids will be in the emulsion phase and react with methane. For the reaction, solid particles have been considered in a series of interconnected CSTRs. A multi stage model is applied where the whole bed was divided into 10 stages of the same solid mass. A simplified schematic of the used model is shown in Figure 17.

Figure 17: Schematic of the Overall Model Used in this Study

The concept of stages refers to a split of the bed length according to the amount of the solid particles in the bed. A looping algorithm is developed for this staging concept as: in the first

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loop methane will react with oxygen carrier particles assigned in all stages based on the average methane concentration in the respective stage and this will continue till the first stage particles are fully expired; at the start of the second loop inlet methane has no reaction in the first stage but it will possess new hydrodynamics which will be used to find the expiration time of the second stage as well as the unconverted state of the other stages. This looping will continue until all the particles in all stages are fully expired. At the end of each stage, mass and fluid-dynamic equations are solved simultaneously to update gas concentrations, Cb and Ce, the superficial gas velocity, U, the bubble rise velocity, uB, the bubble volume fraction, ε_b, the bubble diameter, dB, and the unconverted oxygen carrier particles throughout the length of the reactor. The variables aforementioned need to be updated at the beginning of each stage for the following reasons: the increase in bubble size determines an increase in the superficial gas velocity and thus in uB and ε_b; the bubble growth along the bed height leads to a decrease in methane concentration in the emulsion phase. All these variables influence the mass transfer term between the two phases.

4.3 Validate Reactor Models 4.3.1 Results of Laboratory Experiments Figure 18 shows the inlet gas flows used for reduction of the CuO/Al2O3 particles with methane gas. The time between subsequent reductions is for oxidation. It has been observed that the inlet gas flow time is reduced after 3rd reduction, which may refer to the degradation of the oxygen transferring capacity of CuO/Al2O3 particles. It may arise from the particle attrition and subsequent dust formation of oxygen carrier particles during the oxidation period where a high velocity of air has been used.

Figure 18: Inlet Gas Flow Rates During Successive Redox

Reaction of Cuo/Al2O3 Particles With Methane

0

2

4

6

8

10

12

8000 9000 10000 11000 12000 13000 14000 15000 16000 17000 18000

Flow

rate

s, L

/min

Time, s/10

Methane ArgonAir

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Figure 19 shows the temperature and bed pressure drop variation during the redox reactions. Bed temperature remained nearly constant during reduction periods but varied during the oxidation due to the lack of quick heat removal from the system. In the oxidation state, reactions are very fast and highly exothermic, which results this temperature fluctuation.

The differential pressure profile shows the bubbling characteristics of the fluidized bed. Large pressure drop is observed during the oxidation compared to the reduction (𝑈𝑈 ≈ 2.5 𝑢𝑢𝑚𝑚𝑚𝑚) periods due to the high inlet gas velocities (𝑈𝑈 ≈ 6 𝑢𝑢𝑚𝑚𝑚𝑚).

Figure 20 shows the outlet gas concentrations after condensation of water as a function of time for the reducing cycles when methane was used as fuel. A sample of 126.6 g of CuO/Al2O3 was used at a temperature of ~720 ◦C. In Figure 22the CH4 is turned on but the residence time in the system delays the response with 10–15 s before the CO2 rapidly increases. CO2 reaches a maximum a few second later and remains constant before the inlet CH4 flow turned off. One the other hand, CH4 increases through the whole cycle. However, methane flow has turned off after a certain volume of CH4 is detected in the exit gas analyzer. CO and H2 concentration remains nearly zero for all the reduction cycles. Worth noting is that all of the CH4 reacts to form CO2 and H2O before any detectable CH4 at the reactor exit. However, below 100 ppm of CO and H2 were detected by the gas analyzer when there was CH4 at the reactor exit. This indicates that a small fraction of the inlet CH4 goes through partial oxidation by producing CO and H2.

Figure 19: Bed Temperature and Pressure Drop during Successive Oxidation and Reduction of Cuo/Al2O3 Particles

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Figure 20: Exit Gas Composition During Reduction of Cuo/Al2O3 Particles with Methane

4.3.2 Comparison of model with experimental results 4.3.2.1 A Different Model of Reactor Staging A number of authors applied the concept of dividing axially the bubbling bed in stages and considering a number of stages of the same length (Jafari, Sotudeh-Gharebagh, and Mostoufi 2004; Porrazzo, White, and Ocone 2014; Hashemi Sohi et al. 2012). A large number of bubbles with its minimum size present at the bottom and the incoming gas, which encounters the solid particles, create turbulent motions that result in a low solid volume fraction at the bottom. As shown in Figure 21, the length of the reactor is divided into 10 stages based on the equal mass percentage in each stage but different length. This subdivision method also provides a validation of the total mass balance of the solid particles used in the experiment. It shows that there is a 25% (height was 88 mm and 110 mm at minimum fluidization and reactor conditions, respectively) bed expansion which is expected for this bubbling fluidized bed.

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Figure 21: Staging of the Solid Bed Used in the Model Based on the Mass of Solid Oxygen Carrier Used in the Experiment

Total mass of oxygen carrier was 126.6g.

Figure 22 shows the concentration of methane leaving from different stages of the bed. From left to right, different curves represent the unreacted methane concentration. It indicates that all the methane is consumed until the 6th stage and after that methane starts to leave the reactor as unreacted. At 10th stage, all the methane leaves the reactor as unreacted.

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Figure 22: Axial Profile of CH4 Leaving from Different Stages Considered in the Model of the Fluidized Bed Reactor

4.3.2.2 Model Validation The lab-scale fuel-reactor described in the experimental section using a Cu-based oxygen-carrier was simulated according to the model developed in this study. The main design parameters are shown in Table 17, and the operating conditions and experimental results previously described.

Predictions from the model were obtained at the operating conditions used for the experiments. The main outputs of the model consist of the fluid dynamical structure of the reactor – solids concentration profiles and gas flow distribution between bubbles and emulsion phases –, axial profiles of flow and gas composition (CH4, CO2 and H2O), gas composition at the reactor exit, and conversion of the carrier in the reactor. For instance, Figure 23 shows the concentration of CH4, CO2, and H2O in the reactor. The solid concentration is shown by the unreacted amount of CuO in the oxygen carrier during the first looping step described in the model section. Experimental results showed that formation of CO and H2 were negligible. So, the partial oxidation of methane was not considered for model validation. As it can be seen, methane is fully consumed within the first 30 mm of bed height and all the CuO is converted to either Cu or Cu2O.

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Table 17: Parameters Used in Model Predictions

Parameters Values Reference

Concentration of CH4 at STP (kg/m3) 0.66 (Plawsky 2014; Kelly 1973)

Viscosity of CH4 at 720 °C (Pa s) 3.75 E-5

Diffusivity of CH4 through Argon (m2/s) 1.6 E-4

Molecular weight of CuO (g/mol) 79.54

Density of CuO (kg/m3) 6.31 E 3

BET surface area of oxygen carrier (m2/g) 140 Clariant Corporation

Minimum fluidization velocity (cm/s) 6.0 (Daizo and Levenspiel 1991)

Activation energy(kJ/mol) 60 (Garcia-Labiano et al. 2004)

Pre-exponential factor (mol1-n m3n-2 s-1) 4.5 E-4

Reaction order, n 0.4

Orifice density, N (1/m2) 350000 (Daizo and Levenspiel 1991)

Figure 23: Axial Profiles of Gas and Solids Concentration During First Looping Condition

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The main goal of this model is to predict the methane break-through time for CuO/Al2O3 carrier particles. This will enable to choose a best switch time for the 100 kW pilot scale reactor, which will help in overall fuel conversion and carbon capture efficiency of the pilot plant. Figure 24 shows the predicted methane break-through for the first two reduction cycles of the experiment. As discussed, cycle time decreases for reduction with methane due to the particle attrition and/or dust formation, and consequently loss of those particles through elutriation during high air flow rates, the model cannot capture these. Thus, the model is validated only for the first two reduction cycles. It can be seen experimental results falls within the two reduced state of CuO/Al2O3 particles. In both states, the slope of the experimental curve and the model show similar trend. When most of the oxygen carrier particles are in reduced state, the reactor behaves like a plug flow and methane bleeds through the solid bed without any reaction. The model with the assumption of Cu2O reduced state of particles shows better predictability with the experiment.

Figure 24: Concentration Profile - Reduction Period with CH4 as Reducing Gas as Temperature of ~720°C

The first two reduction cycles are shown here for model prediction. State1 and State2 represent the reduced state of CuO as Cu and Cu2O.

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The stoichiometric calculation of methane reaction with CuO/Al2O3 reveals that about 1.92 g or 0.96 g of CH4 is required for the reduced state of Cu or Cu2O for 126.6 g particles. From Figure 20, it is seen methane was fed for about 1.5 min during the first two reduction cycles at ~1.0 L/min rate, which is approximately 1.10 g of CH4. This indicates that the exact reduced state of the oxygen carrier remains with the two reduced state of the CuO under experimental conditions. However, the emission spectroscopy of the oxygen carrier at the oxidized state shows there is copper aluminum oxide in addition to CuO (Figure 25). Further analysis of the oxygen carrier after successive oxidation and reduction phases will help better understanding of the reduction reactions and accordingly the model prediction can be improved for methane break-through time.

Figure 25: Presence of Different Oxides in Oxygen Carrier Supplied by Clariant

Image courtesy: Clariant.

4.3.3 Conclusions A multi-stage model has been developed to simulate behavior of a lab-scale fluidized bed reactor. The model considers all the processes affecting the reaction of fuel gas with the oxygen-carrier, such as reactor fluid dynamics, reactivity of the oxygen-carrier and the reaction pathway. The main outputs of the model consist of the fluid dynamical structure of the reactor, e.g. solids concentration profiles and gas flow distribution between bubbles and emulsion phases, axial profiles of flow gases and composition (CH4, CO2 and H2O), gas composition at the

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reactor exit, and solids conversion in the reactor. The model shows better correlation with the experiment when reduced state of CuO particles is assumed. Understanding the oxide and reduced states of the oxygen carrier in successive cycles will improve the reaction rate calculation, and thereby the prediction of fuel break-through time from the reactor will be better predicted.

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CHAPTER 5: Lab Scale Experimentation 5.1 Introduction and Summary of Experiments For the lab scale experimentation, ZERE designed a small reactor that would allow for cold fluidization, hot fluidization, and reaction experiments using various kinds of metal oxide particles in a fluidized bed. Lab scale experiments were performed to validate the thermodynamic and computational fluid dynamic analysis and gain data on a variety of different oxygen carrier particles. The ability to perform experiments at a small scale allowed ZERE to evaluate a number of different oxygen carrier particles that could be obtained easily and at reasonably low cost in small quantities.

5.1.1 Experimental Setup and Apparatus Figure 26 shows a simplified schematic of the ZERE lab scale fluidized bed reactor. The reactor is made of quartz glass to allow for reaction temperatures in the range of 800-900 deg. Celsius. The reactor has a total length of 500mm and an inner diameter of 47mm. The oxygen carrier particles were placed on a quartz frit (porous plate) at the bottom of the reactor creating a bed of material 80-100mm in height.

Insertion points for thermocouples and a differential pressure sensor were built into the glass to measure temperatures and pressures as shown in Figure 27. Mass flowmeters measured inlet gas flows of argon, carbon dioxide, methane, and air. Heat required to preheat the reactor was provided via an electric heat tape controlled by a digital thermocouple controller or a variable autotransformer (Variac). The flow rate of the outlet gas was not measured but a portion of the gas from the reactor was led to an electric cooler where the water was removed, and then to a gas analyzer (Testo 350). The Testo® 350 unit provides real time monitoring of O2, CO, CO2, NO, NO2, and CxHy (total hydrocarbons). For some experiments the Testo® 350 also provided data on H2. For the sulfur capture experiments, the Testo® 350 was modified to measure O2, CO, CO2, H2, H2S, SO2, and CxHy. Data was recorded using a National Instruments DAQ with LabVIEW and also through the Testo® software.

The reactor is operated by flowing fluidizing gasses and fuel through the inlet in the bottom of the reactor. The bed is preheated to a temperature of 600-700 deg. C, usually with air flowing through the bed to assure that the bed is in a fully oxidized state when the cycling process begins. In a few cases, the bed was kept still during preheat or was fluidized with argon during the preheat phase. Once preheated, gas is flowed up through the frit and the bed is fluidized allowing for the reaction of the bed material alternately with fuel or air as needed for reduction or regeneration in a cycle between metal and metal oxide. The exothermic nature of the combustion reaction means that there will be a release of heat and therefore a subsequent temperature rise. To limit this temperature increase, methane was diluted with argon before being fed into the reactor. Thus, large temperature increases were avoided since there was no possibility to cool the reactor in the present setup. Figure 28 shows a close up of the fluidized bed of hot copper particles in the reactor.

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Figure 26: ZERE Lab Scale Reactor Schematic

68

Figure 27: Lab Scale Reactor During Cold Fluidization Test

Figure 28: Lab Scale Reactor With Red Hot (800°C) Copper Particles

Thermocouple

Differential Pressure Taps

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5.1.2 Oxygen Carrier Evaluation Prior to this work, ZERE worked in partnership with Professor Hongfei Lin at the University of Nevada, Reno to synthesize 250g each of inert supported copper oxide particles using wet impregnation method to attain 15%, 30%, and 45% copper loading on the inert γ-Al2O3 supports with a final particle size is between 50 mesh and 35 mesh. As part of the current project, ZERE worked with a number of suppliers and sources to obtain seven addition bed particles for evaluation. All of the particles were evaluated on the basis of cold and hot fluidization behavior, activation temperature, reactivity, agglomeration behavior, and particle uniformity. Table 18 lists all the particles evaluated in the laboratory experiments. Based on the analysis of particles and the ability to get a large quantity of oxygen carrier particle, approximately 500kg, manufactured, ZERE selected the Clariant 30% particles for the prototype system. Performance of the Euro Support particles was similar to that of the Clariant particles. After the Clariant particles were selected for the prototype system, the majority of lab scale experiments were performed with those particles. Relative behavior of the particles is discussed further in the sections on oxygen carrier cycling.

Table 18: Bed Particles Evaluated

Particle Designation/Description Manufacturer Size (microns)

Custom or Stock

15% CuO on gamma-Al2O3 UNR 300-500 Custom



15% CuO on macroporous TiO2 UNR 300-500 Custom

20% Cu on gamma-Al2O3 Johnson Matthey 100-500 Stock

20% Cu on zirconia Johnson Matthey 100-500 Stock

30% Cu on gamma-Al2O3 (1) Euro Support 300-500 Custom

30% Cu on gamma-Al2O3 (2) Euro Support 300-500 Custom

30% Cu on gamma alumina Clariant 300-500 Custom

Copper powder Alfa Aesar 100-420 Stock

CuO/Fe2O3 on gamma alumina NREL* 100-200 Custom

*The CuO/Fe2O3 particles were manufactured by a subcontractor for NREL who provided the particles to ZERE.

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Figure 29 and Figure 30 show examples of the bed particles used.

Figure 29: Thirty Percent Copper/Copper Oxides on Alumina Support

Figure 30: Thirty Percent Copper/Copper Oxides on Alumina Support (30X Magnification)

5.1.3 Fluidization Experiments Each set of particles was subjected to fluidization tests. These tests are used to determine the minimum fluidization rates of the particles and determine optimal flow rates for high temperature operation. Minimum fluidization flow rates change significantly with temperature so experiments are performed over a range of temperatures. The data can then be used to determine optimal flow rates and can also be used in fluid dynamic calculation to predict bed behavior. The figures show typical fluidization tests of a bed at room temperature and then the same bed at 400°C. Figure 31 shows a cold fluidization test. The results show a minimum fluidization rate slightly less than nine liters per minute. Figure 32 shows a particle fluidization test for the same bed of particles at a nominal bed temperature of 400°C. The results show a

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minimum fluidization rate of slightly more than 2 liters per minute. Optimal fluidization rates can be 2 to 10 times the minimum fluidization rate.

Figure 31: Cold Particle Fluidization Test

Figure 32: Hot Particle Fluidization Test

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5.2 Oxygen Carrier Cycling with Production of Steam and CO2 from Biogas and Natural Gas Each different set of particles previously shown was subjected to a group of cycling tests. For each of these test, the reactor and bed was heated with an external electric heater. The preheat temperature was typically in the range of 600-650°C. Once the bed reached the desired preheats temperature, cycling of the reactor could begin. Since preheat was typically conducted with air flowing through the reactor the cycling would begin with the introduction of methane diluted with argon into the reactor. Fuel would continue to be added to the reactor until emissions data indicated a reduction in CO2 output and breakthrough of CO or CH4 or both. Some experiments were conducted allowing large amounts of methane to break through while other experiments were conducted in a manner to try and avoid methane breakthrough. Once the emissions reached the limit set point for the experiment, fuel was turned off and air was introduced into the reactor. Air flows were typically continued until the O2% in the exhaust increased above 0%. In some cases, air was allowed to flow until O2% came close to atmospheric values (approx. 21%). Each set of particles was studied to verify its ability to reduce methane and to regenerate by reacting with oxygen in the air. For all experiments, the temperature profile was monitored. In some cases, the goal of the experiment was to maintain a given temperature set point within a variation band. In others, the temperature was allowed to vary widely as long as the reaction was maintained.

Figure 33 shows the flow data for an experiment where the copper bed was cycled with alternating flows of air for oxidation and methane diluted with argon for reduction. You can see that it is possible to maintain the overall bed temperature within a 10-15°C band over multiple cycles using the same fuel flow rate. The bed temperature can be increased and then maintained in a higher temperature range by increasing the fuel flow rate. The air flow rate is also increased when the fuel flow rate is increased in order to oxidize the same amount of bed material that was reduced during the fuel phase. This is done so that parallel reactors can be used simultaneously, one each for fuel and air. When the bed material is reduced by fuel in one stack, the other is oxidized. Flow of fuel and air are then reversed, yielding a continuous process.

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Figure 33: Flow Profile of Fuel and Air Cycles (approximately 750°C)

Figure 34 shows the flow profile of three cycles of an experiment that contained 37 cycles. This zoomed in view is useful for observing the flow and emissions behavior. Figure 35 shows the emissions behavior associated with the fuel and air flows show in Figure 36. During these three cycles, the temperature is maintained in a 30°C band between 730° and 760°C. These emissions results show the exact behavior that is desired. After the introduction of fuel, the system produces only CO2 and steam (which is not measured) for a significant period of time. When the CO2 production begins to drop, CO begins to rise and afterwards begin to see methane breakthrough. This staged emissions behavior allows for the system to be controlled to prevent methane emissions. Once CO is detected, fuel flow can be turned off and air mode can begin. You can see that during air mode, no O2 is detected in the emissions. This means that the copper bed is removing all of the oxygen in the air and using it to regenerate the bed.

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Figure 34: Flow Profile of Three Fuel and Air Cycles (approximately 750°C)

Figure 35: Emissions Profile of 3 Fuel and Air Cycles at approx. 750°C

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Figure 38 shows a typical NOx emissions profile. As mentioned in Chapter 3, NOx emissions are generally between 0-1 ppm, with occasional values in the 1-3ppm range.

Figure 36: NOx Emissions During Cycling

5.3 Oxygen Carrier Cycling with Production of Steam and CO2 from Biomass Solid Fuel Oxidation A limited number of experiments were conducted with solid fuels. Although the original scope included expanding the number of solid fuels investigated, it was discovered during experimentation that the lab scale equipment did not function well enough to provide reliable data on the differences between fuels.

When working with solid fuels, the fuel particles must be small enough that they can react in the fluid bed and they also must be allowed a long enough residence time for the char to be reduced to ash.

The lab scale reactor was modified and a solid fuel inlet port was added in the side wall near the bottom of the fluid bed (Figure 37). Solid fuel was placed in the inlet port and attempts were made to feed the fuel into the reactor during operation. Although some fuel did enter the reactor, it could not be continuously fed without the solid fuel sticking in the inlet. In addition, some of the solid fuel tended to devolatilize in the inlet tube. Multiple variations were tried to improve the fuel feed but none were successful. After many attempts, it was decided that it would be more effective to test the solid fuel in more of a batch mode. Solid fuel was added to the oxygen carrier bed a small distance above the frit and then the remainder of the bed was added on top of the fuel. The bed was preheated either without gas flow or with a small flow of argon.

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Figure 37: Lab scale Reactor with Solid Fuel Feed

During preheat, at around 300-400°C water was released from the solid fuel and could be seen dripping out of the bottom of the reactor. At 600-700°C argon flow in the bed was increased to fluidize the bed and promote reduction of the solid fuel. Production of CO2 was observed but the char particles became very small and were entrained in the gas flow. The reactor was not tall enough to allow for disengagement of the particles from the gas stream and much of the char and ash exited the reactor or was caught on the exit screen. Full oxidation of the fuel was not achieved and when air was introduced into the reactor the char particles auto ignited and could be seen as small sparks flying around in the reactor. This same behavior was observed for multiple experiments.

Figure 38 and Figure 39 show the data collected during two experiments with corn stover. In both cases, the data show very different behavior from the gas fuel experiments. Significant amounts of CO2 production are mainly seen during the preheat phase of the experiment when the fuel is sitting in the stationary bed and begins to devolatilize. When argon is introduced to fluidize the bed, the CO levels rise almost immediately and no significant heat is generated. When air is introduced into the system, CO2 and CO levels rise again as the fine char particles are exposed to hot air and begin to auto-ignite. When argon is reintroduced, emissions levels go down again due to lack of oxygen in the gas space above the bed where the char particles are

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entrained. A small temperature rise is seen when air is first introduced as the small amount of oxygen that was removed from the bed is replaced.

Figure 38: Solid Fuel Experiment with Corn Stover – Test 1

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Figure 39: Solid Fuel Experiment with Corn Stover - Test 2

5.4 Sulfur Capture in a Fluid Bed One major advantage of ZERE technology is the ability to accept untreated biogas. To demonstrate this ability, tests were performed with simulated biogas consisting of 60% methane, 39.92% CO2, and 800 ppm H2S. As described in section 6.1.4 multiple H2S removal techniques were considered with the final selection using lime, CaO in the reactor bed. Lime is affordable and is effective at trapping sulfur.

Prior to running the tests with the lime and biogas, a baseline reading of the gas was taken with the Testo 350. The biogas was diluted with argon at a 1:1 ratio to stay within the allowable maximum concentration of 50% for CO2 detector. As shown in Figure 40, the reading match the expected percentages of close to 30% CH4, close to 20% CO2, and close to 400ppm H2S.

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Figure 40: Baseline Gas Composition of Test Biogas

For this set of tests, the biogas was first tested with a bed of 148 grams of alumina supported CuO with no lime. Figure 41 shows that the expected result was achieved. The H2S was converted to SO2 during the fuel reaction and there was still good conversion of the CH4 to CO2 and H2O until the CuO in the reactor bed was depleted and unreacted gas began to break through to the exhaust.

Per thermodynamic analysis, a ratio of approximately 160:1 CuO to CaO is more than sufficient to capture the fuel bound sulfur. To assure that the proper ratio was maintained over multiple cycles as the CaO was converted to CaSO4, 10 grams of lime were added to the 148-gram bed. The bed was then cycled as in previous experiments between air and fuel mode. As shown in Figure 42, the lime performed very well. During fuel mode, methane conversion is maintained, CO2 production is high, and H2S and SO2 levels are at 0ppm until the bed is depleted and methane break through occurs. In addition, during air mode, the CaSO4 remains stable and does not decompose and release SO2. This shows that with the addition of a small makeup system and cyclone filtration of properly sized lime particles, in bed H2S removal can be an effective solution.

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Figure 41: Biogas Test without Lime

81

Figure 42: Biogas Test with Lime in the Bed

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CHAPTER 6: Prototype Flex Fuel CHP Plant Design & Construction 6.1 Reactor Design 6.1.1 Introduction Various reactor diameters were evaluated to determine the final dimensions of the ZERE prototype reactors. The final version of the reactor design uses a smaller pipe size (18-inch vs 24 inch) relative to previous designs. This is motivated by the expected benefit of reducing the cost of fabricating the ZERE prototype. The smaller reactor diameter is achieved by reducing the design cycle time to 10 minutes from a previous design target of 20 minutes. Various minor changes associated with these changes are included. The gas distributor was changed significantly through the design process.

6.1.2 Descriptions Two chemical reactors will operate in parallel. At any moment in time, one operates in air mode, and one operates in fuel mode. In either case, the reactor temperature is nominally 800 °C. Each reactor will be filled with 228 kg of particles. The particles are 300 µm, made with 30% copper loaded on porous alumina. A process of wet impregnation is recommended to produce the particles. The small scale high-temperature batch experiments performed were essential to establish the oxygen capacity of these particles at 800 °C.

Both reactors have identical configuration, and are fabricated from schedule 40 pipe, made of a high-temperature steel alloy. Table 19 summarizes three designs. The first column shown is the previous design, while the next two columns show two new designs. The selected design is shown in the column labeled New 1, using 18-inch Schedule 40 steel pipe. The high temperature environment, featuring cyclic oxidation, is especially challenging for steel corrosion. One suitable austenitic steel alloy contains 11%Ni, 21%Cr, and 2% Si. It is sold as 253 MA (Europe) or as UNS S30815 (US), with a maximum operating temperature of 1150 °C. It is said to be less expensive and more easily welded than SS 309 / SS 310. Another option to improve reactor life is the use of ceramic coatings or insulation on the inside of the reactor to protect the steel wall.

Heat exchange within the reactors is done with ½ -inch schedule 80 pipe made from high-temperature steel. The flow of water into the tubes will regulate the production of steam and removal of heat from the reactors. Water flow is first to the reactor in air mode, where most heat is extracted. The flow of water leaving the air reactor passes through the reactor which is operating in fuel mode, to extract further heat, to heat the steam to 200 °C. The exchanger is designed to remove 29 kW, with water entering at 180 °C, and leaving as superheated steam at P = 10 bar, T = 200 °C, at a rate of 114 kg/h (1.4 lb/min). Minimum tube length is 2.0 m in the bed. The heat exchanger is tentatively designed as a “W” shaped tube, extending into the reactor from above, and making use of the baffles for structural support. The actual in-bed length is about 3.6 m as designed, greater than the minimum required (Figure 43). Alternative configurations may be considered.

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Table 1: Summary of Design Comparisons

Old New1* New2

Cycle time (min) 20 10 10

Nominal Sch. 40 pipe size (inch)

24 18 16

Particles needed (kg) (per reactor)

455 228 228

Number of reactors 2 2 2

Weight of steel (kg/m) 254.48 156.26 116.08

Expanded bed height (m) 2.14 2.26 2.96

Freeboard (m) 1.00 1.00 1.50

Height of Plenum (m) 0.1 0.1 0.1

Total reactor height (m) 3.24 3.36 4.56

Weight of reactor wall (kg) 825 525 529

Mass of distributor and baffles (kg) (assume ¼ inch

thick)

38.71 21.54 17.00

Mass of steel per reactor, (kg)

863 547 546

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Figure 43: Reactor Schematic

Column height = 3.26 m

Bed height ≈ 1.6 m

Hot water in (180 °C) Steam out (200 °C)

Baffles (see detailed drawing)

Gas distributor (see detailed design drawing)

Plenum box h = 0.1 m

Gas outlet

0.7 m

0.7 m

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During operation, it was anticipated that fines will be produced from the particles, and it would be preferable to return those fines back to the reactor. Although the reactor is configured to operate as a bubbling fluidized bed during fuel and air modes, the air velocity is high enough, especially during air mode, that many particles will be entrained. To prevent a large amount of particles from leaving the reactor, an internal cyclone is required. Figure 44 shows a schematic of a reverse-flow gas cyclone, the dimensions were selected according to a common industrial design known as the “Stairmand high-efficiency” design. The cyclone has internal diameter D = 12.8 cm, and height 51 cm. It fits easily inside the reactor above the bed. Particles are returned to the bottom of the fluidized bed, promoting particle circulation in the reactor. Gas is exhausted through the roof of the reactor, through 2 ½ inch pipe. The design pressure drop for the cyclone is 0.5 psi, and the cyclone is designed to capture particles 1 µm and larger.

These reactors are nominally designed to operate in a bubbling regime, with U0 / umf ≈ 36 when the reactor is in air mode, and U0 / umf ≈ 9 when the reactor is in fuel mode. The inside of the reactor is baffled to break-up bubbles, thereby ensuring a high conversion of methane in the fuel reactor. There are two horizontal baffles inserted in each bed, which increases the gas-solid contact and, thereby increases fuel conversion. (Breaking apart the bubbles also results in fewer particles being entrained from the bed.) Figure 45 gives the hole arrangement on each baffle. Each baffle is constructed from a high-temperature steel plate, at a thickness suitable for ease of production, but at a minimum thickness of 9.5 mm (3/8 inch) to provide structural integrity. It has been assumed to keep about 30% of the baffle cross sectional area open to gas flow.

The gas distributor is designed with the following goals: (a) pressure drop large enough to ensure uniform distribution across the reactor cross section; (b) minimize jet velocity to reduce attrition of particles; (c) inherent design to minimize back flow of particles through distributor into chamber, when the gas flow is turned off; (d) ease of fabrication. The selected design uses "tuyeres", a common configuration in industrial fluidized-bed reactors. Because the volume of gas flow entering the reactor during air mode is nearly seven times the volume of gas entering during fuel mode, the velocity of gas flowing through the distributor is quite different in the two cases. The pressure drop through the distributor is proportional to the square of that velocity, so that the pressure drop of air flow through the distributor is nearly 50 times that of fuel gas flowing through the same distributor. The pressure drop is a key design parameter, and an alternative approach is required. The detail is provided in Figure 46 and Figure 47. The distributor has two components; eight tuyeres are located near the pipe wall, designated “A”. Four additional tuyeres are found closer in to the center, labeled “F”. When in fuel mode, fuel gas is provided through a two-tier plenum to the central four tuyeres only. When in air mode, all 12 tuyeres are utilized.

This distributor design is unusual, and in both modes gas distribution would seem to be imperfect. An effort was made to balance the requirement for uniform gas distribution with practicality of fabrication and operation.

The plenum chamber accommodates this two-part gas distributor (Figure 48). The fuel gas is fed to an internal secondary plenum chamber, with distribution only to the central four tuyeres. Air flow will be to both the internal (secondary) and external (primary) plenum chambers,

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making use of all 12 tuyeres. Thus, air is provided to the plenum through two different pipes, as indicated. The fuel inlet line has a “Tee” to allow for air feed during air mode.

Figure 44: Cyclone Design

Following dimensions are given relative to the

diameter D:

A 4.0

B 2.5

C 1.5

E 0.375

J 0.5

K 0.5

L 0.2

N 0.5

The design calls for D = 0.128 m (5.0”), and has height A = (4 x D) = 0.510 m, and is connected to a standpipe of length sufficient to reach to a height 0.4 m above the distributor, a length of approximately 2.35 m. The standpipe has ID = E = 4.8 cm, similar to 2- inch Schedule 40 (high temperature) steel pipe. Gas exhaust out the top is through a pipe of diameter N = 6.4 cm, approximately compatible with 2.5-inch Schedule 40 (high temperature) steel pipe.

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Figure 45: Baffle Design

The baffle is made from a 3/8-inch steel plate with numerous 2.0-cm holes through it, on a square pattern. Center-to-center spacing is 3.1 cm. The reactors are baffled. Note: the heat-exchange tubes will occupy four of these holes in both reactors, and they have OD = 2.14 cm, slightly larger than the hole size given.

Figure 46: Distributor Plate Layout, with 12 Tuyeres

Viewed from above, each tuyere has a diameter of 54 mm. Tuyeres annotated with “A” are for air flow, while the tuyeres marked “F” are for flow of both fuel gas during fuel mode and air during air mode. In this design, gas flows through either the four “F” tuyeres in fuel mode, or is distributed to all 12 tuyeres in air mode as summarized in Table 20. The eight tuyeres designated “A” have only five of the eight holes open to gas flow. It is important to position

42.9 cm

Center-to-center spacing is 3.1 cm

3.5 cm

42.9 cm

A

A

A A

A

A A

A

F

F

F

F

θ

19.4 cm

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those tuyeres with the blocked holes facing the nearby wall to prevent erosion. The angle θ = 22.5˚, is chosen so that the four “F” tuyeres are located between alternate “A” tuyeres.

Figure 47: Tuyere Design

Each tuyere has eight radial holes, as indicated, pointing out from the center, at 45° spacing. Each hole is drilled to the center of the shaft, with diameter J. The hole is drilled upward at an angle of 10° toward the center, to prevent particles from raining through when gas flow is shut off. Note that three of the eight horizontal holes are blocked for the eight tuyeres designated “A”, as described above. The shaft is 34 mm diameter, and has a single concentric hole of diameter S drilled up from the base, to a height of 20 mm, allowing for distribution of air to each of the eight radial holes. A cap is located above the shaft, with diameter 54 mm and thickness 3 mm.

Table 20: Summary of Distributor Design

Total number of tuyeres

Number of identical tuyeres

Fuel mode 4 4 “F”

Air mode 12 4 “F”, 8 “A”

22 mm

3 mm

34.0 mm

54 mm

12 mm

10°

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Figure 48: Plenum Chamber

(Not to scale; dimensions as indicated)

The chamber is made from either Schedule 40 16-inch steel pipe, or from ¼ inch steel. Its internal diameter is 42.9 cm, and is 10 cm tall. There is an internal chamber, diameter = 24.6 cm, internal height = 5.0 cm, which can be fabricated from either 10-inch schedule 40 pipe, or ¼-inch steel plate. Tuyeres sit on top. Air is provided to the plenum through 3-inch Schedule 40 pipe, and fuel gas is fed through a 1 ¼ inch Schedule 40 pipe. During air mode, air is fed both to the outside plenum and to the inside plenum. Note that the fuel gas does not mix in the (outside) plenum, but instead, is fed directly to the internal plenum chamber in the center of the distributor plate.

6.1.3 Instrumentation and Control Pressure: A pressure tap is required below the plenum, and one in each of the three vertical sections defined by the baffles, and finally in the gas outlet. A pressure gauge should be installed somewhere in the water line.

Temperature: A thermocouple should be installed in each of the three vertical sections defined by the baffles as well as the water inlet and outlet line.

6.1.4 H2S or SO2 Removal for Gaseous Fuels 6.1.4.1 Effects of H2S on Chemical-Looping Combustion (CLC) Literature Survey 1 Ksepko, Ewelina, Ranjani V Siriwardane, Hanjing Tian, Thomas Simonyi, and Marek Sciazko. 2012. "Effect of H2S on chemical looping combustion of coal-derived synthesis gas over Fe–Mn oxides supported on sepiolite, ZrO2, and Al2O3." Review of. Energy & Fuels 26 (4):2461-72.

Experiments were done in a TGA (Ksepko et al. 2012).

Temperature: 800 °C

Oxygen carrier (OC) Particle size: <200 μm

Reduction: 1. CO 36%, H2 27%, CO2 12%, He 25

2. CO 38%, H2 30.8%, CO2 13%, He 17.8%, H2S 4000 ppm

10 cm

tuyeres

3 in Sch 40 pipe

A A A A F

1¼ in Sch 40 pipe

F Air

Fuel gas

24.6 cm

5 cm

90

Oxidation: Air

Purge: Nitrogen

Reaction Gas Flow Rates: 45 cc/min

Duration:

- Reduction: 120 minutes

- Oxidation: 90 minutes

- Purge: 15 minutes

Findings Table 21 summarizes the documented effects of H2S performance and reaction rates of Fe-Mn oxide chemical looping combustion for coal-derived synthesis gas.

Table 21: Findings and Recommendations about H2S Effects on CLC

Sample Oxygen transport capacity

Performance

during 5 cycle test w/o & w H2S

Reaction Rate

with syn-gas w/o & w H2S

Cost Reactor test

60% Fe2O3,

20% MnO2,

20% Sep.

Great Very good,

stable

Fast, No effect

due to H2S

Very low --

60% Fe2O3,

20% MnO, 20% ZrO2

Great

~19 wt %

Very good No effect due to H2S

Very low Stable H2 & CO fully

combusted

Literature Survey 2 Dueso, Cristina, María T Izquierdo, Francisco García-Labiano, F Luis, Alberto Abad, Pilar Gayán, and Juan Adánez. 2012. "Effect of H 2 S on the behaviour of an impregnated NiO-based oxygen-carrier for chemical-looping combustion (CLC)." Review of. Applied Catalysis B: Environmental 126:186-99.

Experiments were done in a bubbling fluidized bed (500Wth CLC) (Dueso et al. 2012).

Oxygen carrier: 18% NiO, 82% α-Al2O3

Temperature: 870 °C (reduction), 950 °C (oxidation)

Oxygen carrier Particle size: 100-300 μm

Reduction gas: CH4 30%, N2 70%, H2S 500 ppm

Oxidation: Air, 0.45 m/s

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Reaction Gas velocity: 0.1 m/s

Findings After H2S feeding, CO2 concentration decreased and some un-burnt CH4 appeared, while the CO and H2 concentrations increased slowly up to 1.5 and 3 vol.%, respectively. H2S and SO2 were found at very low concentrations (<10 vppm) in the fuel reactor outlet stream. SO2 was mainly released from the air-reactor and its concentration increased with time. A sulfur mass balance in the system and the reactivity loss of the oxygen-carrier were seen after introduction of H2S which explained as part of the sulfur was being accumulated in the oxygen-carrier. Dueso et al claimed that these results were completely reproducible

Air atmosphere possible reactions

Ni3S2 + 3.5 O2 = 3 NiO + 2 SO2

Ni3S2 + 4.5 O2 = 3 NiSO4 + 2 NiO

Nitrogen atmosphere

NiSO4 = NiO + SO2 + ½ O2

Ni3S2 + 7 NiSO4 = 10 NiO + 9 SO2

In any case, the subsequent sulfur balance showed that part of the sulfur present in the samples remained inside the particles despite of the successive regeneration steps with nitrogen or air.

The amount of SO2 released at temperatures below 1173 K was very low. Independently of the oxygen concentration and SO2 partial pressure, decomposition started being noticeable at 900 K.

A maximum of 12 wt.% of the sulfur could be emitted as SO2 at the oxidation conditions in air atmosphere.

The SO2 release decreased as the O2 content in the gas fed to the reactor raised, since the nickel sulfide oxidation to form nickel sulfate was favored when the oxygen concentration was high.

The complete regeneration of the oxygen-carriers during oxidation in the air-reactor was not possible.

Literature Survey 3 Forero, CR, P Gayán, Francisco García-Labiano, LF De Diego, A Abad, and J Adánez. 2010. "Effect of gas composition in chemical-looping combustion with copper-based oxygen carriers: fate of sulphur." Review of. International Journal of Greenhouse Gas Control 4 (5):762-70.

Experiments were done in a bubbling fluidized bed (500Wth CLC) (Forero et al. 2010).

Oxygen carrier: 14% CuO, 86% ϒ-Al2O3

Temperature: 800 °C (reduction), 900 °C (oxidation)

Oxygen carrier Particle size: 300-500 μm

Reduction gas: CH4 20-30%, N2 70-80%, H2S 800-1300 ppm

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Findings • Complete conversion of CH4 was not affected by the presence of H2S when there was

higher amount of excess O2 (90% excess OC). All the sulfur introduced as H2S in the fuel reactor (FR) was transformed into SO2. A very small amount of SO2 was detected at the outlet of air reactor, which can be absorbed into the OC as there was no leakage observed for CO2 and CO.

• Complete conversion of CH4 was not attributed when there was lower amount of excess O2 (30% excess OC), and SO2 concentration was high. This behavior was attributed to the formation of copper sulfide, Cu2S, in the FR which produced a deactivation of the oxygen carrier and a clear decrease in the combustion efficiency. Accumulation of Cu2S in OC was observed in successive recycle and no steady state was found.

• Cu2S was formed in the external and internal parts of OC.

• After complete regeneration of OC, the reactivity of the OC was restored as a fresh sample.

• At OC: Fuel ≥1.5, the H2S did not affect the combustion efficiency because the great majority of sulfur was released in the FR as SO2 and no copper sulfide was formed.

- Reactions produced in the FR

4CuO + CH4→ 4Cu + 2H2O + CO2 (1)

CuO + CH4→ Cu + CO + 2H2 (2)

2CuO + H2S → 2Cu + SO2 + H2 (3)

2CuO + 2 H2S → Cu2S + 2 H2O + S (4)

CuO + H2 → Cu + H2O (5)

CuO + CO → Cu + CO2 (6)

2Cu + H2S → Cu2S + H2 (7)

CO + H2 O ↔ CO2 + H2 (8)

2 H2O + H2S → SO2 +3 H2 (9)

2CO2 + H2 S → SO2 +2CO + H2 (10)

CO2 + H2 S → COS + H2 O (11)

CO + H2 S → COS + H2 (12)

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Reactions produced in the Air Reactor (AR)

Cu + 1/2 O2 → CuO (13)

Cu2S + 2 O2 → CuSO4 +Cu (14)

Cu2S + 2 O2 → 2CuO + SO2 (15)

6.1.4.2 Comparison of Different H2S Removal Processes from Biogas Table 22 compares the advantages and disadvantages of various available technologies to remove H2S or SO2 from biogas.

Table 22: Advantages and Limitations of Different Technologies to Remove H2S from Biogas

Process Advantages Limitations

Iron sponge (Abatzoglou and Boivin 2009)

1. Iron-sponge-based H2S removal is operated in batch mode with separate regeneration, or with a small flow of air in the gas stream for continuous, at least partial, regeneration.

2. Operate at low temperature (~50 oC)

3. Commercially available

4. Unit could last up to two years before change-out would be necessary.

5. works at ~85% efficiency

1. Due to S build-up and loss of hydration water, iron-sponge activity is reduced by about one-third after each regeneration cycle.

2. Regeneration is only practical once or twice before a new iron sponge is needed.

Alkaline adsorption (Abatzoglou and Boivin 2009)

1. Alkaline solutions rapidly react with diffused H2S

1. The low interest is attributed to the high reactivity of CO2 with alkaline solutions. Thus, this method is much less selective for H2S, and the captured CO2 consumes relatively expensive alkaline reactants (i.e. NaOH or CaO).

Lime stone (Fenouil, Towler,

1. About 97 wt % CaCO3, can reduce H2S to a level of

1. Sulfidation is usually conducted by lowering the reactor

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and Lynn 1994; Brooks and Lynn 1997)

100-200 ppm, well below that mandated by the Clean Air Act, with conversion of 90+% of the calcium to CaS.

2. Cheap and commercially available.

temperature to 800 °C. 2. Sintering of CaCO3 might inhibit

the sulfidation of limestone.

3. CO2 seems to act as a catalyst for the sintering process; the surface area loss is still significant even if the fraction of CO2 is down to 5% with 95% N2.

4. Borgwardt and Roache studied the sulfidation of limestone over the temperature range 570-850 C for particles of size 1.6-100 μm. They found the rate of reaction to be inversely proportional to particle size and to be high initially, but to fall off rapidly above 11% conversion.

Activated carbon (AC)(Abatzoglou and Boivin 2009)

1. Commercially available 2. Works well under a wide

range of temperature and humidity conditions

1. Large volume of water required and acidic stream produced during regeneration.

2. Biogas contains water is that this gas-borne water reacts with CO2 forming carbonates and contributes to the formation of sulfurous acid, which deactivates the basic catalytic sites, resulting in decrease in capacity

Scrubbing (Abatzoglou and Boivin 2009)

1. Although there are several solvent-based gas-scrubbing technologies for scavenging H2S from gaseous streams in large-scale industrial operations (mainly refineries), applications

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of such technologies in biogas are not known.

2. During alkaline scrubbing, CO2 is also retained intensively, thus increasing the cost of alkaline chemical (i.e. NaOH or Na2CO3) consumption and cost.

6.1.4.3 Conclusion The reaction of the sulfur compounds with the oxygen carrier (OC) may form metal sulfides/sulfate that can deactivate the OC in a bubbling fluidized bed, decreasing their reactivity and therefore, the combustion efficiency of the process. However, TGA experiment of iron oxygen carrier shows no impact on conversion or efficiency.

Among all the available technologies to remove H2S from biogas, iron sponge poses as a potential removal technology for H2S for ZERE project. However, this is a wet process and is not suitable for the integrated semi-batch bubbling fluidized bed reactor systems. Thus, the use of limestone in fluidized bed will be more convenient and efficient in this case.

6.2 Prototype Plant Design Based on the reactor design provided, a prototype plant was designed to incorporate the design features of the reactor system (Figure 49). The 100kW thermal input size of the reactor system imposed certain limitations on the overall system design. The typical maximum electrical efficiency of a power generation system is approx. 30-35%. With a 100kW thermal input, the maximum electrical output of the system would be near 30kW. This led to find a commercially available turbine rated for 30kW or less. After an extensive search, a 15kW steam turbine (Figure 50) was found. The steam needs and isentropic efficiency of the steam turbine were used as design criteria for the balance of the system design. The decision to base the design on an atmospheric pressure reactor system meant that energy in the reactor gas flows had to be used indirectly through heat transfer into a closed loop steam system.

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Figure 49: Prototype Reactor System

As shown in the thermodynamic calculations in Chapter 3, these design criteria and the isentropic efficiency of the turbine led to a calculated system output of approx. 12kW electrical. Elements of the system design are:

Reactor System The reactor system as described in the design material includes two bubbling bed reactors with internal gas distributor, baffles and cyclone, and fuel and air inlet to allow for mode switching.

Exhaust Gas Heat Exchanger The exhaust gas heat exchanger shown in Figure 51 was designed to accept an inlet gas temperature of 850°C and have an outlet gas temperature of 140°C or less. This heat transfer provides enough heat to assure that all of the water at the required flow rate will be converted to saturated steam before entering the reactor beds to be superheated.

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15kW Steam Turbine The steam turbine is a high speed turbine that generates 15kW electrical output with an inlet steam flow of 144kg/hr @ 200-220°C. The high speed turbine spins at 26,000rpm and generates three phase AC power at 1000Hz. After rectification the DC power produced is at 500V.

Figure 50: 15kW Steam Turbine

Closed Loop Steam System The steam piping and valving system shown in part (Figure 51) was designed to allow for the control feature mentioned in Chapter 4 where the water/steam flow is always directed first through the air reactor. The steam loop is designed to take heat from both the reactor exhaust and the reactor beds (Figure 52). Water from the supply tank is pumped into the exhaust gas heat exchanger where it is turned into steam. The steam then flows through the air reactor and then the fuel reactor to transition from saturated steam to dry steam. If the steam does not reach the required temperature, it will bypass the turbine. Once the steam has reached the required dry state, it flows into the steam turbine to generate power. The steam is pulled from the turbine with a vacuum pump and fed through a condenser and then returned to the supply tank.

Instrumentation and PLC/HMI System The prototype instrumentation was designed to give proper control of the prototype as well as give insight into the operation of the system that can be applied to future development. Thermocouples were placed to allow monitoring of all inlet temperature, four different positions within the reactor and at the HRSG inlet. Pressure readings are taken throughout the system with the differential pressure measured across the fluidized bed. As discussed in Chapter 5, the differential pressure is used to monitor the fluidization state of the bed. All of the instrumentation data is fed into the system PLC where it can be recorded for data analysis and used for system control. The integrated control panel shown in Figure 53 contains all of the instrumentation I/O modules as well as the relays, contactors, and power supplies needed to operate the prototype. HMI screens, some samples are shown in Figure 54, have been designed

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to allow for real time monitoring of the system parameters. System valves can also be manually controlled from the HMI. For remote operations, the HMI screen can be mirrored on a PC using a VPN connection. A full control philosophy for the system, contained in Appendix D, was developed and used for the control system programming.

Figure 51: Steam System

Figure 52: Exhaust Gas Heat Exchanger

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Figure 53: Control Panel

Figure 54: HMI Screen Examples

Preheat System To function, the metal oxide bed material must be preheated until it reaches a reactive temperature. In lab scale experiments the copper oxide particles show some reactivity at temperatures as low as 300°C. However, the reactivity is significantly increased at higher temperatures and lab experiments performed much better when preheat temperatures were closer to 600°C. The prototype beds and reactors also have a significant thermal mass that needs to be heated. For the preheat system, a Sylvania 18kW electric heater was selected (Figure 55). The heater has a maximum operating temperature of 760°C. The heater is only run during pre-heat mode and is controlled with a dedicated control panel.

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Figure 55: Prototype Preheat System

The bill of materials for the system is included on the system process and instrumentation diagram included in Appendix C.

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CHAPTER 7: ZERE Prototype Flex Fuel CHP Plant Testing and Operation The prototype system, shown in Figure 56, was shipped from the manufacturer in a partially disassembled state, to fit in the shipping container. After arrival at the Prospect Silicon Valley Demonstration Center in San Jose, the prototype system was placed in the facility high bay space, reassembled, and fitted with an exhaust duct to direct system exhaust gases outside of the facility as shown in Figure 57.

Figure 56: ZERE Prototype in Shipping Configuration

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Figure 57: ZERE Prototype with Exhaust Duct and Insulation

Once reassembly and facility integration was completed, prototype testing began. Unfortunately, when prototype testing was started a ground fault shorted out the control system and the HMI had to be replaced. Once the HMI was replaced, subsystem testing was able to resume. Basic functionality of all of the system valves and instrumentation has been confirmed. The preheat system has been tested and has functioned properly.

Work is continuing on tuning the control system and performing calibration of the inlet flow measurements and other instrument readings. Although the system is not yet fully operational the ZERE team will continue to work on the system in order to operate the prototype and demonstrate electric power production.

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CHAPTER 8: Commercial System Plan 8.1 Design of ZERE Commercial Scale System 8.1.1 Introduction To expand the usability and market for the ZERE system it is desirable to have a truly fuel flexible system that can accept solid biomass fuel as well as gaseous fuel. A design that still has many of the benefits, such as limited particle mass transfer, of the ZERE prototype design, but incorporates the features needed for solid biomass fuels was developed.

Biomass combustion simply means burning organic material. Since biomass fuels are primarily composed of carbon, hydrogen and oxygen, the main products from burning biomass are carbon dioxide and water. Flame temperatures can exceed 2000°C, in traditional systems, depending on the heating value and moisture content of the fuel, the amount of air used to burn the fuel and the construction of the reactor. While wood is the most common feedstock for biomass combustion, almost any plant material can be used as a combustion feedstock.

As in gaseous fuel systems, oxygen has to be provided to sustain the fuel reaction. Char is the remaining material after all the volatiles have reacted with the oxygen in the bed. At or about 800 ºC, the char oxidizes. Again oxygen is required, both at the bed for the oxidation of the carbon and, secondly, above the bed where it mixes with carbon monoxide to form carbon dioxide. Long residence time for fuel in a reactor allows the fuel to be completely consumed. Considering the complete reduction of char, the cycle time is increased to 20 minutes for solid fuel, which requires twice the amount of oxygen carrier used during gaseous fuel case. This increased amount of oxygen carrier needs a tall reactor if 18-inch pipe is used as reactor. Consequently, this solid fuel version the reactor design uses a bigger pipe size (24-inch vs 18 inch) relative to the previous designs for gaseous fuel Other changes associated with these modifications are included.

8.1.2 Commercial System Design The general description of the reactor system is similar to what is proposed for the gaseous fuel design. All of the features previously described including the baffles and inlet distributor are used with this design. Two identical fluidized bed reactors will operate in parallel. At any moment in time, one operates in air mode, and one operates in fuel mode. In either case, the reactor temperature is nominally 800 °C. During fuel mode, white pine sawdust with a particle size of 500 𝜇𝜇𝑚𝑚 and a wood density of 710 𝑘𝑘𝑔𝑔/𝑚𝑚3and heat of combustion value of 20 MJ/kg (assuming 20% moisture in dry sawdust) is considered as solid fuel. With that consideration, a minimum of 0.30 kg/min dry sawdust particles flow is required to get the 100 kW thermal powers.

An approximate wood chemical formula is used to calculate the amount of oxygen carrier particles (CuO on Al2O3).

𝟓𝟓.𝟏𝟏𝟐𝟐 𝑶𝑶𝒖𝒖𝑶𝑶+ 𝑶𝑶𝑯𝑯𝟏𝟏.𝟓𝟓𝟓𝟓𝑶𝑶𝟎𝟎.𝟔𝟔𝟔𝟔−→ 𝟐𝟐.𝟎𝟎𝟔𝟔 𝑶𝑶𝒖𝒖𝟐𝟐𝑶𝑶 + 𝑶𝑶𝑶𝑶𝟐𝟐 + 𝟎𝟎.𝟕𝟕𝟐𝟐 𝑯𝑯𝟐𝟐𝑶𝑶 (18)

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During the fuel mode, saturated steam at 300 ℃ will be used to feed the dry sawdust through a J-type valve into the reactor. Additionally, saturated steam at 300 ℃ will be used to fluidize the bed materials. In air mode, 20% excess oxygen flow is assumed for full oxidation of the oxygen carrier particles.

Each reactor will be filled with 530 kg of particles. The particles are 300 µm, made with 30% copper loaded on porous alumina. A process of wet impregnation is recommended to produce the particles.

Both reactors have identical configuration, and are fabricated from schedule 40 pipe, made of a high-temperature steel alloy. The Table 23 summarizes the designs. The first column shown is the design for solid fuel, while the next column shows the previous design (Gaseous FUEL). Figure 60 shows a sketch of the simplified reactor design. Recommended construction materials are the same as for the gaseous fuel design.

Table 23: Summary of Solid vs. Gas Only Design Comparisons

Solid fuel Gaseous fuel

Cycle time (min) 20 10

Nominal Sch. 40 pipe size (inch) 24 18

Particles needed (kg) (per reactor) 530 228

Number of reactors 2 2

Weight of steel (kg/m) 254.48 156.26

Expanded bed height (m) 2.60 2.26

Freeboard (m) 1.00 1.00

Height of Plenum (m) 0.10 0.10

Total reactor height (m) 3.70 3.36

Weight of reactor wall (kg) 957 525

Mass of distributor and baffles (kg) (assume ¼ inch thick)

43.8 21.54

Mass of steel per reactor, (kg) 1000 547

Heat exchange within the reactors is done with ½ -inch schedule 80 pipe made from high-temperature steel. The flow of the water into the tubes will regulate the production of steam and removal of heat from the reactors. Water flow is first to the reactor in air mode, where most heat is extracted. The flow of water leaving the air reactor passes through the reactor which is operating in fuel mode, to extract further heat, to heat the steam to 200 °C. The exchanger is designed to remove 29 kW, with water entering at 180 °C, and leaving as superheated steam at

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P = 10 bar, T = 200 °C, at a rate of 114 kg/h (1.4 lb/min). Minimum length is 2.0 m in the bed. The heat exchanger is tentatively designed as a “W” shaped tube, extending into the reactor from above, and making use of the baffles for structural support. The actual in-bed length is about 3.6 m, greater than the minimum required (Figure 58). Alternative configurations may be considered.

Figure 58: Simplified Reactor Design

These reactors are nominally designed to operate in a bubbling regime, with U0 / umf ≈ 36 when the reactor is in air mode, and U0 / umf ≈ 9 when the reactor is in fuel mode. The inside of the reactor will be baffled to break-up bubbles, thereby ensuring a high conversion of methane in the fuel reactor. There will be two baffles inserted in each bed to break apart bubbles, which

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will increase the gas-solid contact and, thereby increase fuel conversion. (Breaking apart the bubbles will also result in fewer particles being entrained from the bed.) Again, the baffles design, mirrors that of the gaseous fuel design. It has been assumed to keep about 30% of the baffle cross sectional area open to gas flow.

During operation, it is anticipated that fines will be produced from the ash (post combustion of sawdust) particles, as well as, oxygen carrier particles, and it would be preferable to return those oxygen carrier particles back to the reactor and separate out the ash particles from the exhaust hot gas stream before it goes through the boiler. Although the reactor is configured to operate as a bubbling fluidized bed during fuel and air modes, the air velocity is high enough, especially during air mode, that many fine oxygen carrier particles will be entrained with the ash particles.

Saturated steam at 300 ℃ is suggested to use for solid sawdust feed into the reactor through a J-type valve (Figure 59) arrangement. The minimum steam flow rate is 6.2 m/s for the 100 kW thermal power design. The pressure drop inside the pipe is 0.20 psi. Inside the reactor, feed pipe length is 0.4 m. Four opening points as shown in figure will be used for solid discharge. Number 1-3 will be directed towards periphery and number 4 will be in downward direction. Due to the large volume change from feed line to reactor inside, the change in steam pressure inside reactor is assumed to have minimal effect on fluidization.

To prevent a large amount of particles from leaving the reactor and to separate the ash, two external cyclones are required. First cyclone, the “Stairmand high-flow rate”, will be used to separate particles of 10 𝜇𝜇𝑚𝑚 and higher from the exhaust gas and returns those particles back in the reactor. Figure 60 shows a schematic of a reverse-flow gas cyclone, and the dimensions are selected according to a common industrial design known as the “Stairmand high-flow rate” design. The cyclone has internal diameter D = 10.8 inch, and height 43.2 inch. The design pressure drop for the cyclone is 0.2 psi.

It has been assumed that ash particles will be smaller than 10 𝜇𝜇𝑚𝑚 and they will be separated out by the “Stairmand high-efficiency” cyclone. The exit gas from cyclone 1 will be used as the feed for the second cyclone to separate the ash particles. Figure 61 shows a schematic of a reverse-flow gas cyclone, and the dimensions are selected according to a common industrial design known as the “Stairmand high-efficiency” design. The cyclone has internal diameter D = 3.4 inch, and height 13.6 inch. The design pressure drop for the cyclone is 0.2 psi, and the cyclone is designed to capture particles 1 µm and larger.

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Figure 59: J-Valve Design

Cyclone 1: High Flow Rate The design calls for D = 0.271 m (10.70”), and has height A = (4 x D) = 1.08 m, and is connected to a standpipe of length approximately 2.0 m. The standpipe has ID = E = 6 inch, similar to 6- inch Schedule 40 (high temperature) steel pipe. Gas exhaust out the top is through a pipe of diameter N = 8 inch, approximately compatible with 8-inch Schedule 40 (high temperature) steel pipe.

Following dimensions are given relative to the diameter D:

A 4.00

B 2.50

C 1.50

E 0.575

J 0.875

K 0.750

L 0.375

N 0.750

Figure 60: Cyclone 1 Design

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Cyclone 2: High Efficiency Cyclone The design calls for D = 0.086 m (3.40”), and has height A = (4 x D) = 0.346 m, and is connected to a standpipe of length approximately 1.0 m. The standpipe has ID = E = 1.3 inch, similar to 1 1/4- inch Schedule 40 (high temperature) steel pipe. Gas exhaust out the top is through a pipe of diameter N = 1.7 inch, approximately compatible with 2-inch Schedule 40 (high temperature)

steel pipe.

Following dimensions are given relative to the diameter D:

A 4.00

B 2.50

C 1.50

E 0.375

J 0.50

K 0.50

L 0.20

N 0.50

8.1.3 Scalability There are two possible configurations for the ZERE fuel flexible reactor system. One configuration embodies a number of identical reactors, each operating in sequential batch mode. At any moment in time, at least one reactor is in oxidation (air) mode while at least one other reactor is in reduction (fuel) mode. This is intrinsically a sequential batch design, and has been designed, fabricated, and tested for this project. The ZERE two reactor designs could be scaled up for higher fuel flows and greater electrical output by increasing the number of reactors in the system from two to 3, 4 or more as needed. Another embodiment of the fuel-flexible reactor has three distinct reactors, with solids (oxygen carrier particles) flowing

Figure 61: Cyclone 2 Design

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continuously from one to the other. In this continuous design, each zone or reactor is designed for one task; oxidation, reduction, or char burn off.

In process design, batch reactors are deemed suitable for small scale, while large scale economics require continuous reactors. Batch reactors can be scaled up, as required, but the cost of numerous parallel reactors, at some point, will overcome any benefits realized from complexities associated with continuous operations.

It is expected that the largest scale power unit, approximately 50MW or higher, will require a continuous design. Small scale distributed power perhaps below 100 kW will benefit most from the batch reactor design. A transition from batch to continuous is at some intermediate scale between the two. Tradeoffs of capital cost, particle manufacturing, and operational complexity will all be considered in evaluating which design is most appropriate for a given project.

8.1.4 Fully Fluidized Continuous System As previously stated, at larger scale, a fully fluidized system becomes much more feasible. The added cost and complexity of this system is outweighed by the ability to reduce the size of the reactor vessels or number of reactor vessels as compared to the size or number required for the previous design. In addition, with continuous oxygen carrier recycling the initial oxygen carrier load can be decreased with a continuous system. ZERE and UNR have patented a design for a three stage reactor system with would allow for the use of solid and gaseous fuels (Figure 62). The three stage design uses an initial fuel reactor where solid fuels are devolatilized. The partially used oxygen carrier and the solid fuel char are then transferred to a second burn off reactor where the char is fully reduced to ash. The spent oxygen carrier and the ash are then transferred to the air reactor where the oxygen carrier is regenerated. As the oxygen carrier is transferred back to the fuel reactor, the ash, oxygen carrier and air are separated via cyclones in a similar manner to the solid fuel design described above. This fully fluidized system could also operate as a pressurized system that would allow for the direct use of the system exhaust gases in gas and steam turbines, making more efficient use of the energy generated.

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Figure 62: ZERE Patented 3-Stage Reactor Design

Fuel In

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8.2 Rate Payer Benefits 8.2.1 Ratepayer Benefits of ZERE Gas Only Systems The patented ZERE process can help to establish California as a leader in controlling CO2 emissions. The research conducted for this project has demonstrated that ZERE technology has the potential to be used as a cost effective carbon capture technology while utilizing renewable resources within California. ZERE’s gas fueled system can accept untreated biogas, natural gas, and combinations thereof and produce combined heat and power with near zero emissions and a net negative CO2 impact when the captured CO2 is sequestered.

Biogas potential from animal manures, landfill gas, anaerobic digestion of food, leaves and grass from the current MSW disposal stream, and from waste water treatment plants is estimated to be about 93 billion cubic feet of methane per year. (California Biomass Collaborative 2015) This amount of methane has the potential to be used to install more than 1,000 MW of renewable energy capacity. For each MW of capacity installed using ZERE technology with CO2 sequestration, the yearly negative emissions factor would be at least -8,682,036 kg CO2/year. The very low NOx, SOx and particulate emissions from ZERE systems would help in meeting EPA and CARB emissions targets, especially in area like the central valley that have large amounts of biomethane capacity and high rates of severe to extreme ratings for air quality non-attainment. In addition, biogas power is a renewable source of base load power which can contribute to electrical grid stabilization.

Distributed Generation Goals The Governor has set a goal of 12,000 MW of renewable distributed generation (DG) (20 MW or less) by 2020. As of October 31, 2015, more than 7,200 MW of distributed generation capacity was operating or installed in California, with an additional 900 MW pending. This still leaves almost 4,000 MW to be installed by 2020 with more needed after that to meet aggressive greenhouse gas reduction targets. Having additional baseload distributed generation options like ZERE technology can help to meet this goal.

Energy Usability and Grid Reliability As more intermittent sources of renewable energy are installed and added to the electrical grid, reliability of the grid and the ability to meet real time demand becomes more difficult. Integration of solar and wind energy into the grid is already having an effect on load curves and the need for different policies for base load running reserve. As shown in Figure 63 the hourly average net load curve is strongly affected by solar energy which peaks well before the demand peak later in the day. Figure 64 shows the breakdown of renewable distributed generation in California. Use of biomass accounts for only 560 MW of the 7,210 MW installed, with the largest contributor being 5,100 MW of solar energy. ZERE technology could help to increase the overall percentage of base load biomass sources.

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Figure 63: Cal ISO Hourly Average Net Load 2-22-16

Source: (California Independent System Operator 2016)

Figure 64: Renewable Distributed Generation in California (20 MW or Smaller, Includes Self-Generation) by Fuel Type

Source: Renewable Portfolio Standard Tracking Progress (California Energy Commission 2015)

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Dairies Each of California’s 1.8 million cows produces 140 pounds of waste per year. The total waste production from California’s dairy cows is equivalent to 3.4 million bone-dry tons per year. When that waste is collected and contained in an oxygen free environment, such as an enclosed lagoon, the decomposing waste produces methane, which can be converted to electricity, transportation fuels and/or pipeline gas. Currently, California dairies convert less than one percent of the total dairy waste to energy. If all dairy waste were collected and converted to energy, it would be enough to generate: 11.8 billion cubic feet of biomethane per year. Energy from dairy waste can provide base load renewable electricity, which is available 24/7 and can complement wind and solar power. (Bioenergy Association of California 2014)

Bioenergy Potential in the Solid Waste Sector California sends more than 20 million tons per year of organic waste to landfills — that is food and food processing waste, yard and other green waste, lumber and construction debris, soiled paper and other organic materials. The organic waste that is already in landfills produces enough biogas as it decomposes to generate - about 1,900 MW of renewable electricity. (Bioenergy Association of California 2014)

Waste Water Treatment Plants California’s Waste Water Treatment Plants (WWTP) generate more than seven billion cubic feet of biomethane per year according to the Fuels Potential from Organic Residues in California compiled by Rob Williams of UC Davis (May 2014.). If WWTPs co-digest food and/or other organic waste, they could generate 450 megawatts of renewable electricity, enough to power a quarter of a million homes. WWTPs could generate even more megawatts of power…if all the facilities have anaerobic digestion and use the biogas that they generate to produce electricity and transportation fuels. Of California’s 557 WWTPs, only 97 do not have anaerobic digestion onsite. Of the remaining 460 plants that have anaerobic digestion onsite, 42 do not beneficially use the biogas that they generate.

Landfills According to the U.S. EPA LMOP Database – there are 186 candidate or potential landfills in Californi, with 30 sites rated as possible landfill projects in the state (Figure 65) (https://www3.epa.gov/lmop/projects-candidates/candidates.html)

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Figure 65: U.S. EPA Map of Operating and Candidate Landfill Projects

Work Force Bioenergy creates two to six times as many jobs as natural gas and can provide jobs and income in rural and disadvantaged communities. Statewide, bioenergy employs more than 5,000 people and could employ several times that many. Bioenergy can also provide revenue (and energy) for local governments and public agencies that generate energy from their solid waste and wastewater treatment facilities. (UC Berkeley’s Green Job Calculator, http://rael.berkeley.edu/greenjobs)

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8.2.2 Added Benefits of Solid Fuel Capability Incorporating the ability to accept solid fuels into the ZERE offering would expand the total market for ZERE technology and extend all the benefits mentioned above. California biomass considered to be available on a technically sustainable basis is estimated at 35 million BDT/y. The current technical potential includes more than 12 million BDT/y in agriculture, 14 million BDT/y in forestry, and nine million BDT/y in municipal wastes. Dedicated crops are being grown mostly on an experimental basis at present and are not included in the total. Gross electrical generation potential from biomass is currently near 9,900 MWe with more than 2,300 MWe from agriculture, 3,500 MWe from forestry, and 3,900 MWe from municipal wastes including landfill and sewage digester gas. (California Biomass Collaborative 2015)

Expanding into this larger market would possibly allow for greater economies of scale and higher system efficiencies which would in turn bring down the cost of ZERE systems and reduce the cost of carbon capture further.

8.3 Commercialization Plan Renewable energy and carbon capture technologies continue to be a growing area of interest in the United States and the world. ZERE technology has the potential to be a valuable addition to the efforts for a low carbon energy future. A staged plan has been developed to bring ZERE technology from prototype development to commercial introduction. Each phase has two major areas, technical development and business development.

Phase 1 Technical development goals for phase one were to get the prototype unit fully operational and perform a series of tests to determine lifetime behavior of bed material and to inform decisions on changes needed for a commercial unit with the same design basis as the prototype; a unit capable of using a variety of different gaseous fuel inputs. Specifically, the prototype design needs to be converted into a design that is grid independent with all auxiliary power provided by the system. Startup power can be provided by a battery system and preheat can be provided with a burner system instead of the electrical system used on the prototype. In addition, the control system needs to be modified to provide for autonomous operation with push to start and push to stop functionality.

Business development goals for Phase 1 were to seek and procure investment from the business community to build out the business side of the company. Investment should be sufficient to add staff that can perform sales and marketing tasks such as customer research, market segmentation analysis, and development of future customer relationships. Seek out partnerships with businesses aligned with ZERE. For instance, develop a partnership with an anaerobic digester company in order to provide a full system offering. Business development should also investigate pathways to licensing of ZERE patents as a source of revenue. In addition, the investment procured should allow for additional technical staff to speed up the development of a commercially deployable product.

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Phase 2 Technical development goals for Phase 2: Build and deploy first commercial units utilizing the design completed in Phase 1. Take lessons learned from first deployments and use them to make any needed improvements to the commercial design. Continue development of alternative system designs that allow for the use of solid fuels.

Business development goals for Phase 2: Seek and procure funding through loans, grants, project finance or PPAs for deployment of first commercial unit. Expand the market and sales of ZERE first generation system. Develop a customer base for the ZERE solid fuel offering and find first solid fuel system customer. Procure funding for ZERE operations to continue through Phase 2.

Successful completion of Phase 2 of the commercialization plan, would make ZERE an attractive target for acquisition by a larger firm interested in ZERE technology. Alternately, enough sales and revenue may be generated from patent licensing and new business for the company to be self-sustaining.

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GLOSSARY

Term Definition

EPIC Electric Program Investment Charge

AIIO Air Independent Internal Oxidation

Al2O3 Alumina

Biogas Gas consisting mainly of methane and carbon dioxide formed from the anaerobic digestion of organic waste.

CaO Calcium oxide, lime, quicklime

CCHP Combined cooling, heat, and power

CH4 Methane

CFD Computational Fluid Dynamics

CHP Combine Heat and Power

CLC Chemical Looping Combustion

CO Carbon Monoxide

CO2 Carbon Dioxide

Cu Copper

CuO Copper oxide, cupric oxide

Cu2O Copper oxide, cuprous oxide

Emissions Gases exiting to the atmosphere

Flex Fuel A system capable of accepting many different fuels such as multiple gases or gases and solids.

GREET Greenhouse Gases, Regulated Emissions, and Energy Use in Transportation Model – A greenhouse gas inventory system developed by The Argonne National Laboratory. The California GREET model was developed by the California Air Resources Board as a California specific version based on the original GREET model.

H2O Water

H2S Hydrogen Sulfide

OC Oxygen carrier: A metal oxide compound capable of transferring oxygen in a reactor system.

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Term Definition

PIER Public Interest Energy Research

Silanes Saturated chemical compounds containing one or more silicon atoms i.e. SinH2n+2

Siloxanes A functional group in organosilicon chemistry with the Si–O–Si linkage

SO2 Sulfur Dioxide

Symbology Definitions

𝐴𝐴 Cross sectional area of the fluidized bed 𝑎𝑎 Stoichiometric coefficient

𝐶𝐶𝑏𝑏 Concentration of gas species in the bubble phase, mole/m3

𝜕𝜕 Stoichiometric coefficient

𝐶𝐶𝑒𝑒 Concentration of gas species in the emulsion phase, mole/m3

Greek symbols

𝑑𝑑𝐵𝐵 Bubble diameter, m 𝜖𝜖𝑚𝑚𝑚𝑚

𝐷𝐷 Diameter of the fluidized bed, m 𝜌𝜌𝑚𝑚 Gas density, kg/m3

𝐷𝐷𝑎𝑎𝑏𝑏 CH4 diffusivity in CO2, m2/s 𝜌𝜌𝑝𝑝 Particles density, kg/m3

𝑔𝑔 Acceleration of gravity, m/s2 𝜇𝜇 Gas viscosity, Pa s

𝐻𝐻 Solid bed height, m Φ𝐵𝐵 parameter for bubble diameter

𝐻𝐻𝑚𝑚𝑚𝑚 Bed height at minimum fluidization, m

𝑘𝑘 Rate constant of reaction, 𝑚𝑚𝑚𝑚𝑎𝑎𝑣𝑣1−𝑛𝑛𝑚𝑚3𝑛𝑛−2𝑠𝑠−1 Subscripts

𝐾𝐾𝑏𝑏𝑒𝑒 Mass transfer coefficient between bubble and emulsion phases, s-1

𝑏𝑏 Bubble phase

𝐾𝐾𝑐𝑐𝑒𝑒 Mass transfer coefficient between cloud and emulsion phases, s-1

𝐵𝐵 Bubble

𝐾𝐾𝑏𝑏𝑐𝑐 Mass transfer coefficient between bubble and cloud phases, s-1

𝑣𝑣 Cloud phase

𝑛𝑛 Order of reaction 𝑣𝑣 Emulsion phase 𝑁𝑁𝑑𝑑 Number of orifice per unit area of the bed, m-2 𝑚𝑚𝑚𝑚 Minimum fluidization

𝑢𝑢𝐵𝐵 Bubble velocity, m/s 0 Initial

𝑢𝑢𝑚𝑚𝑚𝑚 Gas velocity at minimum fluidization, m/s 𝑏𝑏𝑣𝑣 Bubble rise

𝑈𝑈𝑏𝑏𝑏𝑏 Single bubble rise velocity, m/s

𝑈𝑈0 Inlet superficial gas velocity, m/s

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REFERENCES

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A-1

APPENDIX A: Details of Economics Calculation Table A-1: Line-by-Line Detail of Economics Calculation (1-MW Biogas Fueled)

Biogas Biogas

Economics & CO2 (both use biogas):

1 MW 1 MW

Description of Line Item Units ZERE Conventional

1

Size in kW (net electric) kW elec 950 950 2

Net/Shaft (= GenEff x NetElec/GrossElec) % of gross 78.40% 88.20%

3

Size in kW (gross electric) kWgross 1212 1077 4

gross eff. (work to shaft, before generator) % 21.63% 25.48%

5

gross heat rate from this Btu/kWh 15,775 13,392 6

MMBtu/h energy input MMBtu/h 19.1 14.4

7

Line 6 in kW thermal input to plant kW thermal 5,602 4,227 8

net heat rate (Line 6 / Line 1) Btu/kWh 20,121 15,183

% net efficiency from above line % 16.96% 22.47%

9

HHV of gaseous fuel (biogas: CH4 + CO2) Btu/scf 650 650 10

Size in scf/minute gas input SCFM 490 370

11

Capital Cost in $/kW(net kW) $/kW 3500 3000 12

Capital cost of the plant $ million 3.33 2.85

13

%/year assumed for capital recovery %/y 15% 15% 14

$M/y allowed for capital recovery $M/y 0.50 0.43

15

cents/kWh for operating costs cents/kWh 4.0 3.0 16

Size: net electric kW (Line 1) kW net 950 950

17

$M/y operation if 24/7 (8760 h/y) $M/y 0.33 0.25

18

annual capacity factor % % 85% 85% 19

result: hours/year in operation hours/y 7446 7446

20

MWh/y net generation MWh/y 7,074 7,074 21

fuel cost in $/MMBtu $/MMBtu 0 0

22

fuel input/year in MMBtu MMBtu/y 142,332 107,401 23

fuel cost in $M/y $M/y 0 0

24

fuel cost in $/MWh (net MWh) $/MWh 0 0

25

$M/y operating cost (Line 17 x L18) $M/y 0.28 0.21 26

$M/y capital recovery (Line 14) $M/y 0.50 0.43

27

$M/y total if no fuel cost (L25+L26) $M/y 0.78 0.64 28

Line 27 in $/MWh (for the MWh/y in Line 20) $/MWh 110.5 90.4

29

cents/kWh if no fuel cost cents/kWh 11.1 9.0 30

cents/kWh fuel cost (from Line 24) cents/kWh 0 0

31

TOTAL: c/kWh, with fuel cost included cents/kWh 11.05 9.04

(continued next page - CHP calculations)

A-2

Table A-1 (continued): Line-by-Line Detail of Economics . . .

(Table A-1 continued - CHP calculations)

Biogas Biogas


1 MW 1 MW


Combined Heat & Power (CHP) Credit for Heat:

32

gross electric gen in kW (=kWh/h) kWgross 1,212 1,077 33

gross electric gen in MMBtu/h MMBtu/h 4.1 3.7

34

energy input in fuel (from Line 6) " 19.1 14.4 35

"waste heat" potential (Line 39 - L38) " 15.0 10.7

36

% of Line 40 for sale or use that % 33% 66%

displaces natural gas

37

MMBtu/h nat gas displaced (L40*L41) MMBtu/h 4.9 7.1 38

hours/year of heat sales or valuable use hours/y 7446 7446

39

$/MMBtu value of heat sales/use $/MMBtu 5.00 5.00 40

$M/y credit for heat sales $M/y 0.18 0.26

41

MWh/y of elec net gen MWh/year 7,074 7,074 42

credit for heat sales per unit elec gen cents/kWh 2.60 3.73

43

net TOTAL cost of elec (L31 - L42) " 8.45 5.31

Carbon Emissions:

44

net heat rate (from Line 8) Btu/kWh 20,121 15,183 45

energy density in fuel (from Line 10) Btu/scf 650 650

46

biogas standard cubic ft. (scf) / kWh (net) scf/kWh(net) 30.96 23.36 47

% by volume CH4 in biogas input % (volume) 65% 65%

48

vol. % CO2 in biogas ==> scf CO2 / kWh is this -->

CO2 scf/kWh 10.83 8.18

49

vol. % CH4 in biogas, 1 mol CO2 per mol CH4 ==> CO2 scf/kWh 20.12 15.18

50

CH4 + 4O ==> CO2 + 2H2O, so scf H2O = 2*above

H2O scf/kWh 40.24 30.37

51

Total CO2 per net kWh (L48 + L49) CO2 scf/kWh 30.96 23.36 52

% of CO2 captured and not emitted % capture 97% 0%

53

CO2 emission factor in scf/kWh (net kWh) CO2 scf/kWh 0.929 23.359 54

density of CO2 in kg per scf (based on g CO2 / scf 55.9 55.9

0.1234 lbs/scf and 453 grams/lb.)

55

Emission Factor: grams CO2 / kWh gCO2/kWh 51.9 1305.7

1 MW 1 MW

Combined Heat & Power (CHP) and net fossil CO2 avoided: ZERE Conventional

56

gross electric gen in kW (=kWh/h) kWgross 1,212 1,077 57

gross electric gen in MMBtu/h MMBtu/h 4.1 3.7

58

energy input in fuel (from Line 6) " 19.1 14.4 59

"waste heat" potential (Line 58 - L57) " 15.0 10.7

60

% of Line 59 for sale or use that % 33% 66%

displaces natural gas

61

MMBtu/h nat gas displaced (L59 x L60) MMBtu/h 4.94 7.09 62

heating value of nat gas displaced Btu/scf 983 983

63

scf/h of nat gas use displaced scfCH4/h 5,029 7,217

(continued next page - CHP & CO2 avoided)

A-3

Table A-1 (continued): Line-by-Line Detail of Economics . . .

(Table A-1 cont. - CHP, CO2, summaries)

Biogas Biogas


1 MW 1 MW


Continuation of combined heat & power (CHP) and net fossil CO2 avoided:

64

CO2 emission fctr nat gas (grams/scf) gCO2/scfCH4 55.9 55.9 65

kg CO2 per hour emission avoided kgCO2/h 281 403

66

net electric gen (kW = kWh/h, Line 1) kWe net 950 950 67

avoided gCO2/kWh (kWh net elec.) gCO2/kWh 295.9 424.7

68

corrected emission fctr. (Line55-Line67) gCO2/kWh -244.0 881.1

Cost of CO2 Avoided (no LCA credits):

69

cost/kWh net (w/ heat credit) cents/kWh 8.45 5.31 70

Conv. gen cost/kWh (w/ heat credit) " 5.31 n. a.

71

avoided CO2 via ZERE vs. Conv. gCO2/kWh 1125.1 n. a. 72

extra cost of ZERE system cents/kWh 3.14 n. a.

73

cost of CO2 avoided by ZERE vs. Conv. cents/gCO2 0.002790 n. a. 74

cost above converted to $/metric ton CO2 $/MTco2 $27.90 n. a.

75

ZERE cost per net kWh generated cents/kWh 8.45 5.31 76

conventional gen cost / net kWh gen. " 5.31 n. a.

77

net CO2/kWh from upstream per GREET gCO2/kWh -807 -685 78

LCA grams CO2 net emission g/kWh gCO2/kWh -1051.0 196.1

79

avoided CO2 via ZERE vs. conventional " 1247.1 n. a. 80

extra cost of ZERE system cents/kWh 3.14 n. a.

81

cost of CO2 avoided by ZERE vs. conv. cents/gCO2 0.002517 n. a. 82

cost above converted to $/metric ton CO2 $/MTco2 $25.17 n. a.

$/MMBtu used as the value of CHP heat

1 MW 1 MW

ZERE Conventional

83

Breakdown of Cost of Electricity: MWh(net)/y = 7,074 7,074 84

Capital Recovery (15%/y per Line 13) cents/kWh 7.05 6.04

85

O&M (labor, maint., chemicals, etc.) " 4.00 3.00 86

Fuel " 0.00 0.00

87

Credit for CHP heat " -2.60 -3.73

---------- ----------

88

TOTAL net cost of net electricity cents/kWh 8.45 5.31

ZERE cost is 3.14 cents/kWh more than Conv.

89

total gas flow out of fuel reactor: CO2 + H2O scf/kWh 71.20 53.72 90

CO2 flow out of fuel reactor: CO2 fraction " 30.96 23.36

91

kW aux power used for CO2 compression kW elec. 135 0 92

kW of other aux power (balance of plant) " 127 127

93

(net electric) / (shaft gross power before gen.) " 78.40% 88.20% 94

gross kW electric " (gross) 1212 1077

95

net kW electric " (net) 950 950

1 MW 1 MW

ZERE Conventional

B-1

APPENDIX B: Bubble Size and Frequency Calculations Assumptions

1. A contour that contains 90% void area will be treated as a bubble area 2. The area of an irregular shape bubble will be treated as an equivalent area of a circle to

calculate the area equivalent diameter of bubble. 3. Contours having area smaller than 0.5 cm2 will not be treated as bubble area. This

condition is employed based on the grid size (6.7 𝑚𝑚𝑚𝑚 × 6.7 𝑚𝑚𝑚𝑚 = 0.449 𝑣𝑣𝑚𝑚2) used in the simulation.

Figure 68: Bubble Area in the Fuel Reactor without Baffle at t = 6.0 s.

B-2

Based on assumption 1, individual bubble areas were calculated. Figure 68 and Figure 69 show

the individual bubble areas of the bubbles shown on the left side of Figure 10.

Figure 69: Bubble Areas in the Fuel Reactor without Baffle at t = 6.0 s

B-3

Figures 70-71 are generated based on assumption 2 where the nonuniform shaped bubble

equivalent area is calculated and the bubble is show as a circle of with the equivalent diameter.

Figure 70 bubble sizes were calculated by considering the individual bubbles areas of Figure 68

as a circle. Figure 71 bubble sizes were calculated by considering the individual bubbles areas

of Figure 69 as a circle.

Figure 70: Circular Bubble sizes in the Fuel Reactor without Baffle at t= 6.0 s

B-4

Figure 71: Circular Bubble sizes in the Fuel Reactor without Baffle at t= 6.0 s.

B-5

Based on the assumption 3 above, the redefined bubble frequency is shown in Figures 72-73

Bubble sizes in Figure 72 are calculated by considering the individual areas of Figure 70 which

are > 0.5 𝑣𝑣𝑚𝑚2. Bubble sizes in Figure 73 are calculated by considering the individual areas of

Figure 71 which are > 0.5 𝑣𝑣𝑚𝑚2.

Figure 72: Circular Bubble sizes > 0.5cm2 in the Fuel Reactor without Baffle at t= 6.0 s.

B-6

Figure 73: Circular Bubble sizes > 0.5cm2 in the Fuel Reactor without Baffle at t= 6.0 s.

C-1

APPENDIX C: Prototype Design Details

Figure 74: System P&ID

Due to the size of the process diagram. A full image is shown here with indications of section markings. Sections are shown individually on the pages below.

A1

A2

B1

A3

B2 B3

C

C-2

Figure 75: System P&ID Section A1

ΤΚ−2

C-3


C-4


C-5

Figure 78: System P&ID Section B1

C-6


C-7


C-8

Figure 81: System P&ID Section C

D-1

APPENDIX D: Control Philosophy The following document outlines the basic function of the ZERE Prototype, with respect to the mechanical portions of the system. The device numbers listed refer to marking in the system P&ID shown above as Figure 74.

The intention of this document is to outline the basic parameters of the prototype water / steam systems, and reactor with heat recovery systems. Reference the prototype P&ID in conjunction with this functional description.

Controls around each system input to the PLC will include set-points for normal operating, start-up, shutdown, high limit, low limit, high-high, low-low, and at least two alarm points each. An HMI with GUI interface is required.

It will be required that the commissioning operator be able to adjust the variable (value) of each set-point. Example, setting the desired steady state temperature and pressure of the steam at TI-5 and PI-5.

Reactor Operation

The system has four stages of operation; start-up, operation, switching, and shut-down.

Mode 1. Start-up. • During startup mode, the reactors F-001 and F-002 must be heated until the reach the

temperature required to start the fuel and air reactions. • Heating is provided by the hot air blower B-101. • The valves CV-101 and CV-102 will both be opened and B-101 turned on. • Temperature in the reactor beds will be monitored. When bed temperature reaches

400 °C in F-001, valve CV-101 will be closed and CV-103 (proportional control) will be opened to start the flow of fuel. Bed temperature will be monitored to verify that the fuel reaction starts and bed temperature continues to rise.

• If bed temperature does not begin to rise, CV-103 (proportional control) is closed and CV-101 is opened to continue pre-heat of the reactor.

• If bed temperature rises, CV-103 will remain open. • When TT-103 indicates 800 °C, the fuel flow will be adjusted to maintain this

temperature. Maximum temperature is 850 °C. • Gas emissions at the outlet of F-001 will be monitored. • When levels of CO begin to rise in the reactor (COT-101 and/or COT-102, and

monitoring with emissions equipment), after reaching operational temperature, CV-103 (proportional control) will be closed, CV-105 (proportional control) will be opened, CV-102 will be closed, CV-104 (proportional control) will be opened, and B-101 will be shut off.

D-2

• TT-103 and TT-107 will be monitored to verify maintenance of the required bed temperature.

• TT-101, TT-102, TT-103, TT-104, TT-105, TT-106, TT-107, and TT-108 will be monitored.

• Note: The steam system will be operating in startup bypass mode during the system startup phase.

Mode 2. Operation

• During operation bed temperature (TT-101 to TT-108) and output gas emissions (TT-109) will be monitored. If bed temperature begins to fall or CO levels begin to increase (COT-101 and/or COT-102, and monitoring with emissions equipment), the system will change to switching mode. Adjustments in air, fuel, and steam flow can be used to adjust bed temperatures slightly.

Mode 3. Switching

• During switching mode, the valve trains will switch the inlets to reactors F-001 and F-002. There are two switching scenarios:

o 1: F-001 is fuel reactor and F-002 is air reactor. Switching will occur by closing CV-103 (proportional control) and CV-106 (proportional control) and opening CV-105 (proportional control) and CV-104 (proportional control).

o 2: F-001 is air reactor and F-002 is fuel reactor. Switching will occur by closing CV-104 (proportional control) and CV-105 (proportional control) and opening CV-103 (proportional control) and CV-106 (proportional control).

• Switching will involve the TT-101 to TT-108, PDT-101 and PDT-102 for air / fuel operation to switch in function control.

• Note: The steam system will also change direction of flow at this time to always maintain the air reactor as the first reactor to receive the steam flow. See mechanical description below

Mode 4: Shut down.

• The cycle may also be initiated by an E-stop command. • During shut down mode, gas flow through the reactors will be shut off by closing all

open valves on the supply side, CV-103, CV-104, CV-105, and CV-106 (all proportional control).

• Cooling may be initiated via the steam system by opening V-2, to allow cooling water to enter the steam system and thereby provide cooling to the reactors.

• The blower B-101 would remain on, but w/o heat, to force un-heated air through the reactors, as desired.

• Once the gas flow through the system has been stopped, reactions in the bed will also stop.

D-3

• Flow of the water/steam will continue until bed temperature reduce to approximately 300-400 °C.

E-HRSG1 • The E-HRSG1 (heat exchanger) is a heat recovery device. This is a sub-system that integrates

with the overall system. • The E-HRSG1 takes the gaseous flow from the reactors at 800 °C to pre-boil the steam that is

to enter the reactors. Water enters the E-HRSG1 and is boiled to saturated steam at 200degC, 10Bar.

• The capacity of the heat exchanger shall be such that for a fixed water flow input, the saturated steam output pressure/temperature can be controlled via reactor fuel input, and control of the by-pass valve V-13. This is a 3-way valve that can throttle the flow of the hot gas to the E-HRSG1.

• The E-HRSG1 is fixed heat exchange device that has both a water flow and a gaseous waste heat flow.

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