Thermodynamic Analysis of Coal to Synthetic Natural Gas Process

Thermodynamic Analysis of Coal to Synthetic Natural Gas Process Lei Chen 1, Rane Nolan 1, Shakeel Avadhany 2 Supervisor: Professor Ahmed F. Ghoniem 1 1. Mechanical Engineering, MIT 2. Materials Science and Engineering, MIT 77 Massachusetts Avenue, Room 3-335 Cambridge, MA 02139-4307 Tel: (617)-258-8740 leichen@mit.edu, wrnolan@mit.edu, avadhany@mit.edu Submitted: May 11 th, 2009 Abstract Natural gas is a clean energy source of the fossil fuels that dominates today s energy supply. The Coal-to-Synthetic Natural Gas (SNG) concept has been successfully demonstrated as a feasible energy production concept. As a final report for term project of Fundamentals of Advanced Energy Conversion, the scope of this research includes a state-of-the-art technologies review for Coal-to-SNG, the thermodynamic parametric study of main components in this process, and the efficiency assessment of the overall energy system implementing different gasification technologies, as well as the novel hydromethanation process. Results show that the optimized oxygen-steam-carbon constraints for Coal-to-SNG are in the range of ~0.25-0.3 for O 2 /C, and ~1.5-2 for H 2 O/C. High pressure is favorable to increase the methanation reaction, and increase the methane yield for hydromethanation. Efficiency analysis shows the moving-bed dry ash gasification achieves higher energy conversion efficiency (67%) than entrained flow gasification (57%) for the overall Coal-to-SNG process. Hydromethanation is a promising novel route with about 70% energy efficiency; however it is still under development because of the technique challenges on catalysts.

Table of Contents Table of Contents... i List of Tables and Figures...ii 1 INTRODUCTION... 1 1.1 Energy Landscape... 1 1.2 Motivation of Analysis... 1 2 OVERVIEW OF THE COAL-TO-SNG PLANTS IN OPERATION... 2 2.1 GREAT PLAINS SYNFUELS PLANT... 2 2.2 GREAT POINTS PILOT PLANT... 2 3 FUNDAMENTALS OF COAL-TO-SNG PROCESSES... 5 3.1 Gasification... 5 3.2 Water Gas Shift (WGS)... 5 3.3 Methanation... 6 3.4 Hydromethanation (Catalytic steam gasification)... 8 4 APPROACHES AND BASIC ASSUMPITIONS OF THE STUDY... 10 4.1 Thermodynamic analysis[12]... 10 4.2 Assumptions... 10 4.2.1 Process simplification... 10 4.2.2 Coal analysis... 10 5 RESULTS AND DISCUSSION... 12 5.1 Parametric studies... 12 5.1.1 Gasification process... 12 5.1.2 Water Gas Shift process... 17 5.1.3 Methanation process... 18 5.1.4 Hydromethanation (Catalytic steam gasification) process... 18 5.2 Efficiency analysis... 21 5.2.1 Base case study: GPSP process with Lurgi gasifier... 24 5.2.2 Efficiency analysis implementing alternative gasifier... 24 5.2.3 Efficiency analysis implementing hydromethanation technology... 25 6. CONCLUSIONS... 27 REFERENCES:... 28 i

List of Tables and Figures Table 1. Typical operating conditions and gas compositions in the 3-stage methanation process.... 7 Table 2: Unit operation models and operating conditions... 10 Table 3: Proximate and ultimate analysis of coal samples... 11 Table 4: Operating condition of typical Lurgi dry ash and GE-Texaco gasifier[2, 11]... 16 Table 5: Syngas composition at process exit (with Lurgi Dry Ash gasification)... 22 Table 6: Syngas composition at process exit (with GE-Texaco gasification)... 22 Table 7: Comparison of efficiencies... 26 Figure 1. Simplifed process of the traditional methanation and hydromethanation technologies... 3 Figure 2. Detailed block flow diagram of the Great Plains Synfuel Plant.... 4 Figure 3. ADAM II 3-stage methanation process at 45 bar, 300-650 ºC... 7 Figure 4. The gas temperature and methane compositions in 3-stage and 1-stage methanation process at 30 bar... 8 Figure 5. Adiabatic gasification temperature map at varying O 2 /C and H 2 O/C ratios.... 13 Figure 6. Cold gas efficiency under adiabatic condition varying O 2 /C and H 2 O/C ratios.... 13 Figure 7. H2/CO ratio in syngas under adiabatic condition varying O 2 /C and H 2 O/C ratios... 14 Figure 8. Methane mole fraction in syngas under adiabatic condition at varying O 2 /C and H 2 O/C ratios... 15 Figure 9. Lurgi and GE-Texaco equilibrium compositions at elevated temperatures.... 15 Figure 10. Sulfur occurrence in Lurgi syngas (equilibrium at elevated temperatures).... 16 Figure 11. CO Conversion ratio in the WGS reactor at varying H 2 O/CO ratios (GE-Texaco syngas, 280 ºC, 40 bar).... 17 Figure 12. Conversion ratio of CO in the methanation process at varying T and P. (Initial H 2 /CO=3:1, at varying pressures and temperatures).... 18 Figure 13. Equilibrium compositions of hydromethanation at elevated temperatures... 19 Figure 14. Energy input required for hydromethanation process at elevated temperatures.... 20 Figure 15. CH 4 mole fraction in hydromethanation products at varying pressures and temperatures.. 20 Figure 16. Carbon conversion ratio in hydromethanation process at varying pressures and temperatures... 21 Figure 17: Heat balance in each component of the Coal-to-SNG process... 23 Figure 18: Energy conversion efficiency in Lurgi and GE-Texaco Coal-to-SNG process... 24 Figure 19: Comparison of energy conversion efficiency among different Coal-to-SNG processes.... 25 ii

1 INTRODUCTION 1.1 Energy Landscape Natural gas is one of the cleanest and most efficient of all energy sources and provides approximately 24% of the world s energy. It provides 23% of U.S. energy consumed and heats 60% of U.S. homes. It also provides a feedstock and fuel for the country's chemical and manufacturing industries and, as of 2007, powered 30% of all U.S. power plants. However, the reserves of natural gas are limited compared with its consumption in the U.S. The U.S. has been the largest importer of natural gas in the world[1]. According to the United States Department of Energy, 90% of all new baseload power plants will be fueled by natural gas. The sudden increase in demand for natural gas will make its price point skyrocket. This presents an opportunity for novel ways to introduce supply into the market. One of these novel ways includes the conversion of coal, an abundant fossil fuel resource in the U.S., to SNG (Synthetic Natural Gas). SNG can be produced from coal, petroleum coke, biomass, or solid waste. The carbon containing mass is gasified and then converted to methane, a large component of natural gas. From a national security standpoint, SNG presents a means to alleviate the reliance on imported energy resources by making the most of an abundant American resource. SNG could be liquefied and transported throughout the U.S. via existing pipeline infrastructure already in-place. Coal is much more evenly distributed throughout the world compared to oil and natural gas, and remains the world s most abundant fossil fuel, with an R/P ratio of more than 130 years, twice that of natural gas[1]. In countries with significant proven reserves of coal but a relative scarcity of natural gas, the Coal-to-SNG process is a promising technology that may provide clean synthetic natural gas for the growing demands of power generation and home utilization. 1.2 Motivation of Analysis The motivation of this analysis is to present an assessment of Coal-to-SNG technologies based on thermodynamic analysis and heat and mass balances, as well as other necessary simplified assumptions. Optimized operating conditions will be investigated via parametric study. The overall performance of the Great Plains Synfuel Plant traditional methanation processes and Great Point Hydromethanation process will be investigated as a benchmark and optimization solutions will be proposed. The implications of a widespread implementation of this technology, i.e. delivery of SNG at competitive prices, will be profound. 1

2 OVERVIEW OF THE COAL-TO-SNG PLANTS IN OPERATION A Coal-to-SNG system converts solid hydrocarbons such as coal, biomass or petroleum coke into SNG. A conventional approach for Coal-to-SNG is by a process of gasification, gas shift, and methanation. This indirect approach has been demonstrated in the Great Plains Synfuel Plant for 20 years and proven to be successful in application. More recently, advancements have been made on the direct gasification approach by the Great Point Energy. This mechanism involves hydromethanation and circumvents the processes of gasification and water gas shift. Figure 1 shows the simplified flow diagram for these two processes. Both indirect and direct technologies for SNG production will be detailed in this report. 2.1 GREAT PLAINS SYNFUELS PLANT The traditional Coal-to-SNG processes have been demonstrated in the Great Plains Synfuel Plant for 20 years and proven to be successful in application. The Coal-to-SNG concept has grown in interest recently due to its capability for CO 2 capture and utilization in enhanced oil recovery. In the Great Plains Synfuel Plant, more than 5 million tons of CO 2 have been sequestered to date, while doubling the oil recovery rate of an oil field in Saskatchewan[2]. A detailed process diagram of the GPSP is shown in Figure 2[2]. The plant consists of coal and ash handling unit, Air Separation Unit (ASU), steam generation, gasifier, water gas shift, and methanation unit, as well as the AGR (Rectisol) and Flue Gas Desulphurization (FGD) unit to remove acid gas in syngas and flue gas. Besides methane, the plant also produces ammonia, ammonium sulfate, naphtha, and phenol as by-products. Although demonstrating successful and economical clean synthetic fuels production, the GPSP can be optimized in many aspects. For instance, the gasification technology, a Lurgi system, was adopted by GPSP 20 years ago and may not be the most favorable option today because of its small coal processing throughput and large production of waste water. Choosing a technology that produced less waste could eliminate or diminish ancillary processes such as gas liquor separation, wastewater treatment, ash handling, and so on. However, replacing the gasification system would require the adjustment of other processes, principally the water gas shift and methanation. We will conduct a parametric study comparing different gasification technologies in the system, to investigate the favorable configuration for optimized efficiency. 2.2 GREAT POINTS PILOT PLANT The hydromethanation, or catalytic steam gasificaiton technology, is considered to be more energy-efficient than the traditional methanation processes. This process was initially developed by Exxon in the 1970s using potassium carbonate (K 2 CO 3 ) as a catalyst. However the process is still under development and not commercialized. 2

O 2 Coal H 2O Steam Steam H 2O H 2S, CO 2 H 2O ASU Gasifier Quench Or Cooling Water Gas Shift Cooling Gas cleaning & CO 2 capture Methanation CH 4 Air Q Q Q Q Q Q Traditional methanation process H 2O H 2S, CO 2 H 2 H 2O Coal H 2O Hydromethanation Cooling Gas cleaning & CO 2 capture H 2 separation CH 4 Q Q Q Q Hydromethanation process Figure 1. Simplifed process of the traditional methanation and hydromethanation technologies. It is better known as direct gasification and involves a lesser number of components than the indirect process. Most notably, it replaces the gasifier and water-gas-shift reactor with a hydromethanation unit which essentially uses steam to react with the coal to form methane. The advantage of hydromethanation is that it doesn t require air separation unit; hence there is less energy penalty for the process. The challenges of catalytic steam gasification are the separation of catalyst from ash/slag and the loss of reactivity of the catalyst[3]. Great Point Energy is currently building a testing plant facility in Somerset, Massachusetts. The design and operating data are limited to obtain because of confidential issues. 3

Figure 2. Detailed block flow diagram of the Great Plains Synfuel Plant. 4

3 FUNDAMENTALS OF COAL-TO-SNG PROCESSES 3.1 Gasification The gasification of coal is a process that ultimately breaks down the fuel into its basic chemical constituents. Instead of burning the coal directly to perform energy conversion, gasification allows for energy in coal to be stored in the form of a gas. Better known as syngas, it consists primarily of carbon monoxide and hydrogen[4]. More specifically, carbonaceous material undergoes the following processes during gasification: pyrolysis (or devolatilization), combustion, gasification and water Gas Shift reaction. In pyrolysis, the carbonaceous material is heated after which volatile products are released and char is produced[5]. It is the char that undergoes the gasification reaction. When the volatiles and char react with oxygen in the gasifier, carbon dioxide and carbon monoxide are formed. Thirdly, the gasification reaction occurs when carbon dioxide and steam produce carbon monoxide and hydrogen. Lastly, the water gas shift reaction occurs when carbon monoxide reacts with water to form carbon dioxide and hydrogen. The gasification reactions can be described by 1 Cs () + O2 CO (3.1) 2 Cs () + CO 2CO (3.2) 2 Cs () + HO CO+ H (3.3) 2 2 Gasification process can be accomplished via different established technologies. These include: entrained flow, fluidized bed, and transport reactor. Commercial vendors that sell gasifiers types include ConocoPhillips, GE Energy/Texaco, Shell, Siemens, KBR Transport, and Lurgi. Lurgi dry ash gasifier is a typical moving-bed gasification technology. A fixed bed of coal or biomass is subjected to the gasification agent, i.e. steam and oxygen or air. Ash is removed as a slag; this type of gasifier requires a high ratio of steam and oxygen to carbon in order to supersede the ash fusion temperature. The thermal efficiency of the process is high, while the syngas exit temperature is low. GE-Texaco gasifier is a typical entrained-flow gasification technology. Slurry of atomized coal becomes gasified in-flow with oxygen. High temperatures of this gasifier ensure that methane and tar are not part of the product gas, mostly H 2 and CO. Due to high temperatures, this gasification method requires great deal more oxygen than the other gasifiers. As a result it places more strain on the air separation unit than the other gasification technologies. Fluidized bed gasifier is another technology option for low rank coal. The carbonaceous material is fluidized in this type of reactor. The lower operating temperatures in dry ash gasifiers require that the fuel be highly reactive. 3.2 Water Gas Shift (WGS) The gas shift reaction can be described by: CO + H2O CO2 + H2. (3.4) 5

The hydrogen content of the syngas coming from the gasifier must be enhanced before it can be converted into methane. An additional injection of steam into the syngas flow under certain conditions promotes the water-gas shift (WGS) reaction of Equation (3.4). The ideal output of the WGS process is a stoichiometric mixture that can be reformed into methane as completely as possible. Carbon dioxide-with hydrogen-may be reformed into methane, or it may be separated from the syngas and sequestered or used for a different purpose to further improve the stoichiometric ratio. The principle reactants for methanation are carbon monoxide and hydrogen. The WGS reaction is exothermic, generating 42 kj per mole[6]. Le Chatelier s principle or a more explicit equilibrium calculation shows that the WGS reaction proceeds further at relatively low temperatures and is pressure insensitive, i.e. thermodynamic equilibrium favors high conversion of CO and steam to hydrogen and carbon dioxide at low temperatures. However, such low temperatures impede the kinetics of the reaction. The productive use of the WGS process at low temperatures therefore requires catalysts to accelerate the reaction rate to an acceptable throughput. Current implementations of the WGS use two stages with different catalysts[7]. The first shift occurs at relatively high temperature to accelerate the reaction. This High Temperature Shift (HTS) catalyst has conventionally been composed of magnetite (Fe 2 O 3 ) and chromium oxides and operates at 400 500 C[8]. The syngas flow is then cooled and sent through a Low Temperature Shift (LTS), at 200-300 C, the catalyst of which may be composed of copper or, for example, the Raney nickel-aluminum alloy. Ultimately, the kinetics of the WGS is the limiting factor, and a compromise must be made between throughput and lowering the reactor temperature to further the shift. Additionally, these conventional catalysts have many downsides: the HTS catalysts may be inactive below a certain temperature, LTS catalysts may degrade at a temperature slightly above their operating range, these catalysts may require complex activation procedures, and they may even be pyrophoric[9]. The Department of Energy has researched other catalysts without such downsides; indeed, progress has been made in developing catalysts that would operate in one stage (at a relatively low temperature)[9]. Such catalysts, possibly composed of platinum/mixed oxides, non-precious metals/mixed oxides, or vanadium-cobalt oxides, have the potential to function with greater activity and can survive exposure to air and a wide temperature range[9]. The improved activity of better catalysts would allow WGS systems to be built smaller and with greater throughput. 3.3 Methanation The methanation reaction can be described by CO + 3H CH + H O (3.5) 2 4 2 CO + 4H CH + 2H O (3.6) 2 2 4 2 As the methanation reactions are highly exothermic and pressure-favorable, the methanation reactors are designed running at low temperature and high pressure with catalysts. Three stages of methanation are a conventional design for methanation technology. Figure 3 shows the schematic diagram of the ADAM II methanation process, illustrating the three stages[10]. Three adiabatic methanation reactors, D 201, D 202 and 6

D 203 are equipped with fixed catalytic beds. The syngas coming from the WGS unit is preheated to a temperature above the starting temperature of the catalyst. At each methanation reactor outlet, the gas compositions are approximately at chemical equilibrium. Heat is generated during the methanation reactions, so syngas cooling is needed between stages. The typical operating temperatures and compositions at the inlets and outlets of the three stages are shown in Table 1. Figure 3. ADAM II 3-stage methanation process at 45 bar, 300-650 ºC Table 1. Typical operating conditions and gas compositions in the 3-stage methanation process. 1 st stage 2 nd stage 3 rd stage Product Unit Inlet Outlet Outlet Outlet (dry basis) Gas flow rate m3/h 962 832 282 271 149 P bar 26.85 26.7 26.6 26.5 26.5 T ºC 306 651 485 343 16 H2O mol% 15.58 28.7 39.57 44.96 0 CH4 mol% 23.29 34.68 43.15 46.79 85.01 CO mol% 5.74 1.81 0.04 0 0 CO2 mol% 6.69 4.79 2.47 0.71 1.29 H2 mol% 44.90 25.70 9.97 2.63 4.78 N2 mol% 3.8 4.32 4.8 4.9 8.92 An alternative technology for the methanation process is 1-stage methanation with cooling. The temperature and conversion rate for the 3-stage and 1-stage methanation process are compared in Figure 4. 7

Figure 4. The gas temperature and methane compositions in 3-stage and 1-stage methanation process at 30 bar 3.4 Hydromethanation (Catalytic steam gasification) The direct method of Coal-to-SNG is a process that performs the same function as the indirect method, except eliminates three of the six steps. With less steps of energy conversion, there is an increase in end-to-end efficiency of producing SNG from the coal. As mentioned previously, the Great Point pilot plant is experimenting with the direct process via hydromethanation. In this process, gasification and methanation occur in the same reactor in the presence of a catalyst[3]. Steam is the only gasification agent used so that gas shift and methanation steps are no longer necessary. The ideal reaction route is: 2C+ 2H O CH + CO (3.7) 2 4 2 However with steam, low temperatures greatly limit the rate of reaction. At high temperatures, the thermodynamic environment is not favorable for methane production. The solution; introduce a catalyst to facilitate the reaction. Alkali metals catalyze carbon with steam to form CO and H 2, by so doing, increasing the reaction rate several fold. Selection of catalyst is based on affinity to reacting with coal. KCl and K 2 SO 4, for example, are ineffective despite their belonging to the alkali family[11]. In the processs of gasification, the actual catalyst is not retained in the gasifier but is carried out with the ash. In order for a commercial plant to maximize profits, it is essential that a recycling loop be implemented to recover the catalyst for re-use in the coal gasification process. From end-to-end in hydromethanation, coal is pulverized and mixed with the selected catalyst. Before feeding the impregnated catalyst-coal into the gasifier, it is dried to remove as much moisture from the fuel as possible. Gasifiers are then fed with the feedstock and begin introducing steam into the environment to perform the gasification. Post the hydromethanation process, carbon monoxide and hydrogen must be separated from the methane product. The cryogenic distillation process effectively separates methane from the synthesis gas (a process with an energy penalty lower than the oxygen separation from air in an ASU). 8

In all practicality, the use of catalysts in the methanation process enable lower temperature and operating pressures, thereby reducing the mechanical and thermal standards for materials used in the gasifier. 9

4 APPROACHES AND BASIC ASSUMPITIONS OF THE STUDY 4.1 Thermodynamic analysis[12] As most of the gasification processes are thermally auto-balanced, the equilibrium states (Temperature, compositions) under adiabatic conditions are calculated. At given constant pressure and initial enthalpy, the equilibrium state is reached when ds 0 at constant ( p, H ) (4.1) Therefore, at equilibrium, when conditions of constant pressure and enthalpy are applied, the total entropy is at maximum. Some of the processes are at specific pressure and temperature, exothermic or endothermic. Constraining the unit to constant T and p, we find that dg = dsg, and at equilibrium under these conditions, the following equation must be satisfied dg 0 at constant ( p, T ) (4.2) Therefore, at equilibrium, the Gibbs free energy must reach a minimum when the state is defined by the pressure and temperature. 4.2 Assumptions 4.2.1 Process simplification The traditional coal-to-sng processes are simplified in this study, key components such as gasifier, quenching chamber, WGS reactor, and methanation reactor are modeled using thermodynamic equilibrium methodology. Other components such as ASU, hydrogen separation and AMINE process are simplified based on the second law efficiency or other models in the 2.62 class. A schematic diagram of the traditional methanation and hydromethanation process is shown in Figure 1. The basic assumptions for these processes are listed in Table 2, as well as their operating conditions. Table 2: Unit operation models and operating conditions Process Simplified unit model Temperature (ºC) Pressure (bar) ASU Second law efficiency[13] 20 25/40 Gasifier Equilibrium 450/1300 25/40 Quench/cooling Steam-water phase equilibrium 200/230[11] 1/25/40 Water Gas Shift reactor WGS reaction equilibrium 250/280[2] 25/40 Gas cleaning and CO 2 capture Simplified AMINE process[13] 35 1/25/40 H 2 separation Second law efficiency[13] 35 1/40 Methanation Equilibrium 300 25/40 Hydromethanation Equilibrium 450 1/40 4.2.2 Coal analysis Different coal types are adopted in the calculation for Lurgi and GE-Texaco gasification process, due to different coal type adaptability in these technologies. The reported proximate and ultimate analyses are given in Table 3, showing high moisture, high volatiles and low heating value for lignite coal than the bituminous coal. 10

Table 3: Proximate and ultimate analysis of coal samples Lignite coal #1[11] For Lurgi dry ash case study Approximate analysis (as received) wt% Moisture 36 8.5 Ash 7.4 10.7 Fixed Carbon 29.4 44.8 Volatiles 27.2 36 Ultimate analysis (dry ash free) wt% Carbon 71.5 78.4 Hydrogen 4.8 5.7 Oxygen 21 10.6 Nitrogen 1.4 1.2 Sulfur 1.3 4.0 HHV (MJ/kg) 16.8 26.3 Bituminous coal #2[14] For GE-Texaco case study 11

5 RESULTS AND DISCUSSION Parametric studies are carried out to understand the fundamentals of three key processes in indirect methanation: Gasification, WGS and Methanation, as well as the hydromethanation concept. Base on the results for each component through thermodynamic simulation, the efficiency of the traditional and alternative gasification coal-to-sng processes are calculated. Although the hydromethanation process is a new concept (hence very limited design data is available), its efficiency analysis is also achieved with some assumptions and simplifications. 5.1 Parametric studies 5.1.1 Gasification process Gasification is such a process that the chemical energy of carbon is used to produce synthetic gas products like CO, H 2 and CH 4. The carbon-steam gasification reactions o C+ H2O CO+ H2 Δ H298 =+ 131.3kJ/mol (5.1) C+ 2H O CO + 2H Δ H =+ 90.1kJ/mol (5.2) o 2 2 2 298 are highly endothermic, so the heat should be supplied through the combustion reactions o C+ O2 CO2 Δ H298 = 393.5 kj/mol (5.3) 2C+ O 2CO Δ H = 221.0 kj/mol (5.4) o 2 298 If all the heat is supplied by direct carbon combustion in the oxygen, rather than by indirect means via electricity or a heat transfer medium, then the carbon-oxygen-steam ratio should be optimized such that the global reaction is thermal self-balanced. The thermodynamic equilibrium temperature and compositions are obtained under adiabatic condition, in which the total enthalpy of products are the same as the feeding stocks (carbon, steam and oxygen at 20 ºC and 25 bar). The map of cold gas efficiency, CH 4 yields, and H 2 /CO ratios at varying O 2 /C and H 2 O/C ratios are calculated, as shown in Figure 5, Figure 6, and Figure 7 respectively. The adiabatic gasification temperature increases with increasing O 2 /C ratio and decreasing H 2 O/C ratio, as shown in Figure 5. Oxygen-carbon reactions are highly exothermic and can supply heat to other gasification reactions or increase the product gas temperature, while the steam-carbon gasification will absorb the heat into chemical energy in syngas. High temperature is favorable to chemical kinetics, so the entrained flow gasification technologies such as GE-Texaco runs at relatively higher O 2 /C ratios and relatively lower H 2 O/C ratios. In contrast, Lurgi dry ash gasification runs at lower O 2 /C ratios and relatively higher H 2 O/C ratios, so the syngas resident time in Lurgi gasifier is much longer. However, the greater oxygen consumption in entrained flow gasifier will require more energy consumption for air separation, and thus lead to higher efficiency penalty. 12

O 2 /C ratio 1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 3.2E+03 2.4E+03 2.8E+03 1.6E+03 8.0E+02 Texaco 1.2E+03 4.0E+02 2.0E+03 Lurgi T ad ( o C) of product syngas Adiabatic condition @ 25 bar 0.0 0 1 2 3 4 5 H 2 O/C ratio Figure 5. Adiabatic gasification temperature map at varying O 2 /C and H 2 O/C ratios. 1.0 0.9 0.8 0.30 0.075 0.23 0.15 O 2 /C ratio 0.7 0.6 0.5 0.4 0.3 0.2 0.1 Texaco 0.53 0.68 0.45 0.45 0.38 0.60 Lurgi Cold gas efficiency Adiabatic condition @ 25 bar 0.38 0.30 0.23 0.15 0.075 0.0 0 1 2 3 4 5 H 2 O/C ratio Figure 6. Cold gas efficiency under adiabatic condition varying O 2 /C and H 2 O/C ratios. The cold gas efficiency ( η cg ) is a key factor assessing the energy conversion efficiency in gasification process, which is defined as nh LHV 2 H + n 2 COLHVCO + nch LHV 4 CH4 ηcg = 100% (5.5) LHV coal 13

1.0 0.9 0.8 H 2 /CO ratio in syngas Adiabatic condition @25 bar O 2 /C ratio 0.7 0.6 0.5 0.4 0.3 0.2 0.59 Texaco 1.2 2.3 9.4 4.7 Lurgi 38 2.3E+02 1.5E+02 3.0E+02 5.3E+02 6.8E+02 6.0E+02 0.1 3.8E+02 19 4.5E+02 0.0 5.3E+02 0 1 2 3 4 5 H 2 O/C ratio Figure 7. H2/CO ratio in syngas under adiabatic condition varying O 2 /C and H 2 O/C ratios. Figure 6 and Figure 7 show the cold gas efficiency and H 2 /CO ratio of syngas at varying O 2 /C and H 2 O/C ratios. η cg is mainly a function of the O 2 /C ratio, while the H 2 /CO ratio is a more determined by the H 2 O/C ratio. 0.5 is the stoichiometric ratio of oxygen-carbon reaction to produce carbon monoxide, smaller O 2 /C ratio will not convert carbon sufficiently, and CO will burn at greater O 2 /C ratios. So η cg reaches the maximum value at this ratio. On the other hand, the water gas shift reaction CO + H O( g) CO + H Δ H = 41.2 kj/mol (5.6) o 2 2 2 298 is nearly thermal-neutral compared with other reactions, so it will convert CO into H2 without changing the η cg much. If the optimized O 2 /C ratio (about 0.5) is chosen, although lower H 2 O/C ratio will slightly increase the η cg, like that of GE-Texaco gasification technology, meanwhile lower O 2 /C ratio and higher H 2 O/C ratio is favorable to produce H 2. Were CH 4 the final product, then low temperature moving-bed gasification technology like Lurgi is favorable in terms of its higher H 2 /CO ratio (about 3-40) in syngas according to the ideal gasifier model. Figure 8 plots the CH 4 mole fractions in syngas at varying O 2 /C and H 2 O/C ratios. Unlike H 2 /CO ratio, the CH 4 mole fraction shows a maximum (about 14%) region around O 2 /C=0.25 and H 2 O/C=1.5. The operating condition of Lurgi dry ash gasification technology used in GPSP is near this region, but with slightly higher O 2 consumption and higher steam injection, as a result, the actual CH 4 mole fraction in Lurgi dry ash gasifier off syngas is about 6% in wet basis, and 11% in dry basis. The higher oxygen and steam ratio design might consider the heat loss and temperature control in the gasifier. Again, as 14

CH 4 is the final product, the high steam-carbon ratio low temperature Lurgi gasification will reduce the cost of water gas shift and methanation in the following processes, and might be favorable in terms of economic optimization. 1.0 0.9 0.8 CH 4 mole fraction in syngas Adiabatic condition @25 bar 0.7 O 2 /C ratio 0.6 0.5 0.4 0.3 0.2 0.1 Texaco 0.064 0.032 0.11 0.14 0.080 0.016 Lurgi 0.096 0.048 0.0 0 1 2 3 4 5 H 2 O/C ratio Figure 8. Methane mole fraction in syngas under adiabatic condition at varying O 2 /C and H 2 O/C ratios. Mole fraction Mole fraction 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0 0.6 0.5 0.4 0.3 0.2 0.1 0.0 H2 H2O CO CO2 CH4 N2 H2O(L) C(S) Lurgi dry ash gasification@25 bar 0 200 400 600 800 1000 1200 1400 GE-Texaco gasification @40 bar T ( o C) Figure 9. Lurgi and GE-Texaco equilibrium compositions at elevated temperatures. Lurgi dry ash gasification and GE-Texaco gasification processes are chosen as typical gasification technologies for coal-to-chemical in this case study. The typical operating 15

conditions of these two processes are listed in Table 4, and coal proximate and ultimate analysis data are shown in Table 3. Table 4: Operating condition of typical Lurgi dry ash and GE-Texaco gasifier[2, 11] Characteristic Lurgi Dry Ash GE-Texaco Entrained Flow Pressure, bar 25 40 Temperature, ºC Combustion zone 980-1370 1300-1500 Gasification zone 650-815 1400 Gas exit 370-540 1300 Gas quenching 200 230 Material input Coal input, kg/s Lignite, 281.8 Bituminous, 105.1 Steam/water, kg/s 297.8 51.8 O 2 injection, kg/s 57 88 Syngas compositions (wt%)* CO 7.6 37.6 CO2 15.9 11.2 H2 20.7 27.0 H2O 48.4 23.0 CH4 6.4 0 N2 and other balance balance * Lurgi case is cited from reference [11], GE-Texaco case is from the equilibrium calculation, compositions frozen at 1300 ºC. 0.004 Mole fraction in equilirium products 0.003 0.002 0.001 0.000 Sulfur occurrence in syngas Lurgi case study H2S S(L) S(S) -0.001 0 200 400 600 800 1000 1200 1400 T ( o C) Figure 10. Sulfur occurrence in Lurgi syngas (equilibrium at elevated temperatures). Given the feeding stocks and operating pressure, the equilibrium compositions of the two gasification technologies are calculated at elevated temperatures shown in Figure 9. At the high gas exit temperature (about 1300 ºC), the composition of GE-Texaco syngas is almost in equilibrium state. But the case of Lurgi syngas is complicated, as the kinetics of different reactions from equation (5.1)-(5.4) and (5.6) are different, the compositions are 16

not in equilibrium state at lower gas exit temperatures (about 370-540 ºC). As have been discussed in the oxygen-steam-carbon feeding ratio parametric study, more H 2 and CH 4 are produced in the Lurgi gasification process, indicating less WGS and Methanation duty in the down flow processes. Sulfur occurrence in the syngas is also plotted in Figure 10, showing the major species is H 2 S at temperatures higher than 400 ºC. 5.1.2 Water Gas Shift process Water gas shift process is employed in GPSP to partially convert carbon monoxide into hydrogen. It is also widely used in other coal-to-chemical processes required hydrogen production. In the coal-to-sng process, the favorable H 2 /CO ratio is 3:1 to avoid carbon deposition and achieve high conversion ratio in the methanation process. As shown in equation (5.6), the WGS reaction is slightly exothermic, nearly thermalneutral, and pressure-neutral, so given the operating temperature, only the H 2 O/CO ratio is variable to achieve the optimized H 2 /CO ratio objective. Figure 11 shows the CO conversion ratio and H 2 /CO ratio in product gas as function of feeding H 2 O/CO ratios given the upstream GE-Texaco syngas. The optimized steam/carbon monoxide ratio is 0.62, at which 57% of carbon monoxide will be converted to hydrogen. 1.1 1.0 Water Gas Shift Parametric Study Syngas from GE-Texaco gasifier 100 80 Conversion ratio 0.9 0.8 0.7 0.6 0.5 0.4 Conversion ratio H2/CO 0.5 1.0 1.5 2.0 2.5 3.0 H2O/CO 60 40 20 0 H2/CO ratio Figure 11. CO Conversion ratio in the WGS reactor at varying H 2 O/CO ratios (GE-Texaco syngas, 280 ºC, 40 bar). 17

1.00 Carbon Conversion Ratio 0.98 0.96 0.94 0.92 0.90 1 20 40 Methanation process H 2 /CO=3, varying T and P 200 300 400 Temperature ( o C) Figure 12. Conversion ratio of CO in the methanation process at varying T and P. (Initial H 2 /CO=3:1, at varying pressures and temperatures). 5.1.3 Methanation process The compositions of feeding stream to methanation reactor are mainly CO, H 2 and CH 4, after removing all the CO 2, H 2 S and other impurities. From the methanation reactions CO + 3 H CH + H O( g) Δ H = 210 kj/mol (5.7) o 2 4 2 298 CO + 4H CH + 2 H O( g) Δ H = 113.6 kj/mol (5.8) o 2 2 4 2 298 We know these reactions are highly exothermic, pressure-favorable. So low temperature and high pressure will facilitate the methanation reactions. Figure 12 shows the carbon monoxide conversion ratio in methanation at varying pressures and temperatures. The carbon monoxide conversion ratio increases with increasing pressure and decreasing temperature as predicted. At atmospheric pressure, the conversion ratio increases from ~90% to 99% with temperature decreasing from 400 to 200 ºC. However, at elevated pressure such as 20 bar and 40 bar, the temperature is not so sensitive, only 2 percentages change within the same temperature range. 5.1.4 Hydromethanation (Catalytic steam gasification) process 18

0.6 H2 CO CH4 C(S) H2O CO2 CH3OH Hydromethanation C/H 2 O=1:1, 1 bar 0.4 Mole fraction 0.2 0.0 200 400 600 800 T ( o C) Figure 13. Equilibrium compositions of hydromethanation at elevated temperatures. Traditional coal-to-sng process has been demonstrated to be feasible in synthetic fuel production, however it involves a serial of separated reactors and processes. An integrated hydromethanation concept is raised based on the gasification, WGS and methanation reactions in equation (5.1), (5.6) and (5.7). Add these reactions one obtains C+ H O 0.5CH + 0.5CO Δ H =+ 7.7 kj/mol (5.9) o 2 4 2 298 which is slightly endothermic, and more endothermic if written in terms of coal[11] CH + 0.82H O 0.41CO + 0.59CH Δ H =+ 17.6 kj/mol (5.10) o 0.72 2 2 4 298 The equilibrium compositions of hydromethanation at atmospheric pressure are plotted in Figure 13 as function of temperatures. The steam to carbon ratio is in stoichiometric 1:1 in this case study. CH 4 and CO 2 are main compositions at temperature lower than 300 ºC, although this temperature is low in terms of reaction kinetics, the H 2 yield will increase with increasing temperatures, which is not favorable to produce CH 4 as the final product. Moreover, the input heat required thermodynamically increases with increasing temperatures as well as shown in Figure 14. So the operating condition should be maintained at lower temperature. Two measures can be implemented to increase the CH 4 yield: by using catalytic and elevating the system pressure. Figure 15 shows the CH 4 yield in the equilibrium state at elevated pressures. Higher pressure converts H 2 to produce more CH 4. A problem for the hydromethanation is low carbon conversion ratio. Figure 16 shows the total carbon conversion ratio in hydromethanation at elevated pressures. It remains about 40% at any pressure in the feasible temperature range for CH 4 production. The features of hydromethanation such as low temperature, high pressure, and low 19

carbon conversion ratio determines that it is not feasible for entrained flow reactor, but other reactors in which the solid carbon and catalytic can be recirculated, or excessive steam can be used. Fluidized bed is one of the options for these characteristics. Heat needed (kj/kg Carbon) 10000 8000 6000 4000 2000 Hydromethanation C/H 2 O=1:1, 1 bar 0 200 400 600 800 T ( o C) Figure 14. Energy input required for hydromethanation process at elevated temperatures. 0.15 1 10 20 40 Hydromethanation C/H 2 O=1:1, varying pressures 0.5 0.4 CH 4 mole fraction 0.10 0.05 CH 4 0.3 0.2 0.1 H 2 mole fraction H 2 0.00 0.0 200 400 600 800 T ( o C) Figure 15. CH 4 mole fraction in hydromethanation products at varying pressures and temperatures. 20

1.0 Carbon conversion ratio 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0 Operating pressure (bar) 1 10 20 40 Hydromethanation C/H 2 O=1:1, varying pressures 200 400 600 800 T ( o C) Figure 16. Carbon conversion ratio in hydromethanation process at varying pressures and temperatures. 5.2 Efficiency analysis To determine the overall efficiency of each methanation system design, we must first define an appropriate measure of efficiency. The yield of such a process is of prime concern, but do we define yield as the mass of product gas per mass of coal processed or do we examine the heating value of product gas compared with the heating value of coal to determine an efficiency of chemical energy conversion? Both measures of efficiency have merit: the mass yield can be easily considered in economic terms and the conversion efficiency may be compared with the efficiencies of other refining and energy production processes. The cold gas efficiency of equation (5.5), illustrating the conversion efficiency of chemical energy of coal into methane, must be altered to account for any enthalpy inputs into the system, for example heat for the hydromethanation reactor for this inherently endothermic direct conversion to occur. The altered efficiency used is, η alt n HHV + n HHV = HHV + h CO CO CH4 CH4 coal processing (5.11) where we include the heating value of hydrogen gas produced in hydromethanation separately, for it is removed from the product stream and is a by-product (the small fraction of CO is allowed to remain in the synthetic natural gas). The enthalpy of processing, h processing, is the additional energy required in the system for processing which is not provided by the exothermic reactions of conversion. 21

Table 5: Syngas composition at process exit (with Lurgi Dry Ash gasification) Gasifier exit Quenching WGS Cooling AGR Methanation Condensation T, ºC 450 200 250 35 35 300 20 Pressure, bar 25 25 25 25 25 25 20 Mole fraction H2O 0.48 0.50 0.50 0.0013 0.0019 0.33 H2 0.21 0.20 0.20 0.41 0.60 CO 0.076 0.074 0.068 0.14 0.20 CO2 0.16 0.15 0.16 0.32 CH4 0.064 0.062 0.062 0.12 0.18 0.63 0.95 N2 0.0072 0.0070 0.0070 0.014 0.020 0.034 0.051 H2S 0.0012 0.0012 0.0012 0.0023 Table 6: Syngas composition at process exit (with GE-Texaco gasification) Gasifier exit Quenching WGS Cooling AGR Methanation Condensation T, ºC 1300 230 280 35 35 300 20 Pressure, bar 40 40 40 40 40 40 20 Mole fraction H2O 0.23 0.23 0.018 0.0014 0.0022 0.49 H2 0.27 0.27 0.48 0.49 0.75 CO 0.38 0.38 0.16 0.16 0.25 CO2 0.11 0.11 0.33 0.33 CH4 9.3E-05 9.3E-05 9.3E-05 9.5E-05 0.00014 0.50 0.98 N2 0.0033 0.0033 0.0033 0.0034 0.0051 0.01 0.020 H2S 0.0094 0.0094 0.0094 0.0095 This enthalpy is potentially a summation of many elements inside a methanation plant. Here, we will examine the effects of the Air Separation Unit (ASU), the hydrogen gas separation required in the hydromethanation process, hydrogen sulfide separation (required when using coal as a feedstock), and CO 2 capture. The first two will be estimated using fundamental thermodynamic principles, while the thermodynamics of CO 2 and H 2 S capture will examined by implementing a unique absorption/desorption process. The energies for separation of oxygen, hydrogen, and hydrogen sulfide are estimated using the isentropic enthalpy of gas separation and estimations of second law efficiencies for real separator systems. The work necessary for separation of one component from the others in such a model is given by the equation, w RT p p = χ ln χ + ( 1 χ ) ln ( 1 χ ) 0 1 1 sep a a a a ηsecond law pa2 p2 (5.12) 22

where X a is the molar fraction of the component to be separated, P a2 is pressure of separated stream, and P 1 and P 2 are the pressure of the main stream before and after separation, respectively. The efficiencies of separation will be estimated as 0.15 for oxygen[13] and 0.3 for hydrogen the low molecular weight of hydrogen gas relative to the other components aids considerably in separation. The enthalpy required to operate the CO 2 and H 2 S capture system used in this analysis is heat energy (unlike for the other separators, for which we assume they are powered by another form of energy, e.g. electricity or shaft work). The input of heat to capture CO 2 and H 2 S works in the following way[13]: the gas stream enters an absorber chamber at low temperature into which a spray of water and amine (NH 2 -CH 2 -CH 2 -OH) is injected. Two moles of amine reacts and bonds with each mole of CO 2 captured. The water and amine solution does not evaporate and instead leaves the absorber rich in CO 2. It is then heated in a desorber chamber to 130 C to break the CO 2 -amine bonds. Additionally 1 mole of water evaporates for each mole of CO 2 captured. As the CO 2 leaves the desorber in gas form, the water and amine solution can be cooled and returned to the absorber. The rather considerable energy necessary to break the CO 2 -amine bonds and to evaporate a large mass of water, however, can be regenerated from cooling processes needed between processing operations according to our analysis. See Table 2 for temperatures of individual processing steps. Heat balance (kj/mole carbon in coal) 100 80 60 40 20 0-20 -40-60 -80 ASU Lurgi GE-Texaco Gasification WGS Quenching AGR Syngas cooling Methanation Condensation Figure 17: Heat balance in each component of the Coal-to-SNG process 23

Energy conversion efficiency 1.2 1.1 1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0 Original coal and power for ASU Gasification WGS Methanation Lurgi GE-Texaco Figure 18: Energy conversion efficiency in Lurgi and GE-Texaco Coal-to-SNG process. 5.2.1 Base case study: GPSP process with Lurgi gasifier The GPSP process, with the Lurgi gasifier, is capable of converting 45% of the carbon in lignite feedstock into methane as shown in Table 7. The only external energy needed to operate the process in our simplified model is for the ASU, which demands 0.26 MJ/kg lignite coal. The chemical energy of lignite was converted to the energy in methane at an efficiency of 0.67, considering the energy demands of the ASU. 5.2.2 Efficiency analysis implementing alternative gasifier The GE-Texaco entrained gasifier operates with a higher quality bituminous coal (Illinois # 6 in this study). GE-Texaco gasifier has a higher O 2 /Carbon ratio for high operating temperature and high coal throughout capability, however the higher temperature achieved in the gasifier occurs through a greater extent of combustion and conversion of carbon into CO 2. Only 33% of the carbon in the coal is therefore converted into methane. The greater oxygen requirements in this gasifier also demand more energy input for the ASU, 1.06 MJ/kg bituminous coal as shown by the External Energy Input in Table 7. The chemical conversion efficiency is, therefore, lower, at 0.57. The higher mass yield of methane per kilogram coal, also shown in the table, only illustrates the higher carbon content of the feedstock. The heat balance of each component in the traditional methanation processes are shown in Figure 17 (positive means energy input, negative means heat output). Energy are required in ASU and AGR (Acid Gas Removal) units, while release in gasification, quenching, syngas cooling, methanation and condensation units. The WGS unit are close to thermal-neutral. GE-Texaco gasification consumes more oxygen, which leads to a 24

larger ASU, higher exothermic degree in gasification and quenching. Also because the syngas contains less H 2 and less CH 4, larger WGS and Methanation units are required, which all lead to energy penalty to the Coal-to-SNG process. The overall energy conversion efficiency through the whole Coal-to-SNG process descends because of heat loss or energy consumption. Figure 18 illustrates the chemical energy conversion efficiency (the chemical energy in syngas over total chemical energy in gasifier inlet, i.e, in coal, steam and power for ASU). Gasification and Methanation are the major components consuming chemical energy. However, exhausted heat shown in Figure 17 can be optimized and recovered, to supply heat for AGR process, or for power generation. 5.2.3 Efficiency analysis implementing hydromethanation technology Opposed to the indirect methanation techniques of gasification, water gas shift, and methanation, hydromethanation is fundamentally an endothermic process and therefore requires an external source of energy in the form of heat for the reaction to proceed. This energy input is the sole component of h processing, and is considered at 100% efficiency. Energy Conversion Efficiency 1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 H2 CH4 0.1 0.0 Lurgi GE-Texaco HM 1 bar HM 40 bar Process Figure 19: Comparison of energy conversion efficiency among different Coal-to-SNG processes. As discussed above, the carbon conversion ratio is only about 50% in hydromethanation process within one throughput. In this efficiency analysis, we assume the unconverted carbon will be recirculated and reused until the carbon conversion is 100%. The CH 4 yield is slightly positively correlated with pressure, which is illustrated in Table 7. However, the product gas composition is enriched with CO at higher pressures (0.6% and 2.2% CO at 1 bar and 40 bar, respectively). The appropriateness for operating at high pressures will depend on the desired specifications of the product gas. 25

Table 7: Comparison of efficiencies Process CH 4 yield Chemical energy External conversion efficiency energy input [kg/kg Coal] [mol/mol C] [MJ/kg Coal] GPSP/Lurgi Plant 0.262 0.451 0.67 0.26 GE-Texaco gasifier 0.312 0.330 0.57 1.06 Hydromethanation (1 bar) 0.62 0.46 0.70 12.5 Hydromethanation (40 bar) 0.63 0.47 0.69 12.3 Hydrogen is produced in the hydromethanation process, it can be recirculated into the hydromethanation reactor to produce CH 4, or separated as a by-product. The latter is chosen as an option in this study to avoid the analysis of recirculation. The overall energy conversion efficiency of indirect and direct methanation processes are shown in Figure 19, hydromethanation processes are higher than the indirect methanation processes, at about 70%, close to the results given by Chandel et. al[3]. Although the pressure doesn t affect the efficiency much, it may beneficial to chemical kinetics and the coal handling capacity. Energy conversion efficiency for indirect methanation implementing Lurgi gasification technology and GE-Texaco gasification technology are 67% and 57% respectively, which shows that the low oxygen consumption, low temperature, moving-bed gasification technology is more favorable in terms of energy conversion efficiency. The efficiency is also close to that given by Gray et. al[15]. The gap between these two gasification technologies might be close after optimizing the heat management. Many other considerations should be taken into account besides the energy efficiency, for instance, the waste water disposal, capacity optimization, and so on. More detail studies are needed in order to optimize the overall efficiency of the Coal-to-SNG process. 26

6. CONCLUSIONS The fundamentals, research and development status of Coal-to-SNG processes are reviewed followed by thermodynamic analysis in this study. Oxygen-steam-carbon constraint under adiabatic condition is investigated, and results show the practical constraint for syngas production. Optimized oxygen-steam-carbon ratios are different for different final products, efficiency, operating conditions and product compositions should be considered comprehensively. For Coal-to-SNG production, the O 2 /Carbon ratio ~0.25-0.3 and H 2 O/Carbon ratio ~1.5-2 are favorable to produce a CH 4 -rich syngas with high H 2 /CO ratio at higher energy conversion efficiency. The optimized steam/co ratio is investigated for WGS reaction given the syngas compositions. The optimized steam/carbon monoxide ratio is 0.62 for methane production using GE-Texaco gasifier. High pressure is favorable for methanation reactions, the carbon monoxide conversion ratio increases from 97% to 99% as pressure increase from 1 to 10 bar at 300 ºC. Pressure doesn t change the compositions much for hydromethanation process, while lower temperature (350-500 ºC) is favorable for CH4 yield within acceptable energy conversion ratio. Catalysts are the most critical technique in this reaction route. The analysis of the simplified models illustrates that coal methanation is practical thermodynamically. Two alternative gasification technologies are compared for the Coalto-SNG processes. The Great Plains method with Lurgi dry ash gasifier recovers 67% efficient at converting coal into methane, comparing with 57% with GE-Texaco gasificaiton. The Lurgi gasifier operates with lower quality coal than the GE-Texaco gasifier and combusts less of that coal for syngas production. Additionally, the ASU needed for the Lurgi gasifier also consumes much less energy than for the oxygen-rich GE-Texaco gasifier. The relatively higher efficiency of hydromethanation (about 70%) indicates that this technique has potential to supplement or supplant the Great Plains model, especially when considering the simpler plant design (shown in Figure 1). Ultimately, the kinetics of the three plant designs will also be of importance when analyzing the efficiency and practicality of the conversion processes. The long residence time for coal in the Lurgi reactor is an economic penalty, for it requires a larger reactor for a certain throughput. The infancy of the hydromethanation industry suggests that catalysts have yet to be developed or commercialized. Future coal methanation analyses will wish to examine in further detail the design of heater exchangers for the regeneration of heat to operate the amine scrubbing operation. 27