Design and Economic Evaluation of a Coal-to-Synthetic Natural Gas

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Design and Economic Evaluation of Coal to Synthetic Natural Gas (SNG) Process Bor-Yih Yu, and I-Lung Chien Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/ie503595r • Publication Date (Web): 12 Feb 2015 Downloaded from http://pubs.acs.org on February 17, 2015

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Industrial & Engineering Chemistry Research

Paper submitted for publication in Ind. Eng. Chem. Res.

Design and Economic Evaluation of Coal to Synthetic Natural Gas (SNG) Process

Bor-Yih Yu and I-Lung Chien*

Department of Chemical Engineering National Taiwan University Taipei 10617, Taiwan

Revised: January 29, 2015

*

Corresponding author. I-Lung Chien, Tel: +886-3-3366-3063; Fax: +886-2-2362-3040; E-mail: [email protected]

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ABSTRACT In this work, the steady state design and economic evaluation for coal to Synthetic Natural Gas (SNG) process is rigorously studied, and this study could give a baseline for design and analysis for SNG production in Taiwan or other countries relied on importation of energy source. SNG is a product that holds very similar composition and heat value to typical natural gas, and can be used as a replacement in industrial and home usages. Natural gas is an important energy source in Taiwan, with increasing demand year by year. Due to the fact that over 99% of energy source depend on importation in Taiwan, and those advantages for coal over natural gas (lower importation price, great abundance, easier transportation and storage…etc), the route that converts coal into SNG is expected to benefit Taiwan if the related technology is successfully established. The whole process is divided into several parts, included Air Separation Unit (ASU), gasification, syngas treating section (water-gas-shift-reaction, syngas cooling and acid gas removal), methanation reaction section, and electricity production block from upstream to downstream. The overall energy conversion efficiency for the plant is 60.38%, with the SNG production cost to be 10.837 (USD/GJ), thus this process will be economically and practically attractive.

Keywords: Coal; Synthetic Natural Gas (SNG); Process Simulation; Economical Evaluation

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1.

Introduction Natural gas has long been considered as a cleaner energy source all over the world,

due to its lower C/H ratio. It generated less carbon dioxide, and no sulfur or nitrogen contained gas compared with burning coal. In Taiwan, there are several natural gas power plants that provide electricity to local industries, with over 1300 MWe generated each year. The largest natural gas power plant is in Datan, Taoyuan, which has six NGCC cycles and produce about 438 MWe, to support the electricity needed in northern Taiwan. Most of the other natural gas power plants locate at central and southern Taiwan, this is due to the fact that the heart of industry is located there. Thus, natural gas utilization is important in our country. Here, over 99% of energy source depend on importation, those included crude oil, coal, natural gas, and so on. For final energy products, electricity and petroleum product are the two largest categories, which combined for 90% of energy consumption. Natural gas used in Taiwan relies on importation (mostly from Qatar) in form of Liquefied Natural Gas (LNG), and the requirement of liquefying and pressurizing make the importation price to become high. The importation price for LNG is often within the range of 14~17 (USD/GJ), which makes the sold price of natural gas high in our country. Thus, to lower the cost on natural gas could have advantages in developing natural gas related process as well as promoting the economics. The major component of natural gas is methane, with slight amount ethane and propane to provide extra heat value. There is still some inert component, such as nitrogen, argon, and carbon dioxide inside. The composition for natural gas will vary with different site, which depend on the identities of geosphere. Also, the regulation for natural gas composition and heat value varies from country to country, but generally

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speaking, the inert component should be less than 4%, and the heat value should be in the range of 36~40 (MJ/Sm3). Synthetic Natural Gas (SNG) is a product with similar composition as natural gas, and which can be used in replacement of natural gas directly1-6. Basically speaking, SNG could be converted from coal, as well as biomass or other kind of fuels. The conversion process started from gasification, which converts the fuel to syngas (mainly CO and H2). The syngas is then passes the treating section, including Water-Gas-Shift Reaction (WGSR), Acid Gas Removal (AGR), and finally methanation section to form SNG. Comparing to natural gas, the heat value of SNG is often a little lower, this is because SNG is synthesized through methanation reaction and thus all the heat value comes from methane component. The typical regulations for SNG could be found in open literature. 1-4 Among the energy sources that imported into Taiwan, coal has good advantages over others, including lower price, great abundance, and easier transportation. These reasons lead to the fact that coal is the major energy source used annually in Taiwan. Compared with the high natural gas import price, the price for coal is about 3.179 (USD/GJ). Thus, the route that converts coal into SNG is expected to have both practical and economical advantages in Taiwan. There are already existing SNG production plants, or some relevant project worldwide, such as in U.S. and China. The U.S. used to be dependent on natural gas, thus it has begun the research in SNG producing early. But nowadays, the recently discovering of shale gas in U.S. makes the priority of producing SNG lower. For China, the natural gas demands are also highly reliant on importation. However, China is abundant in coal reserves, which has benefits in producing SNG. The SNG production is also listed in the key development in China

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in the near future, with a number of large demonstration projects ongoing, or already in operation.7 In Taiwan, due to the fact that the annual natural gas demand is quite large (and probably increasing year by year in the near future) but lacks domestic supply of nearly all energy sources, the technique to produce SNG is in urgent need. To reduce the cost for importing coal, seeking co-operation with countries which has large coal reserves could be the primary aim for developing our SNG plants in the near future. In this work, the design and economical evaluation of a plant-wide coal to SNG process in Taiwan are illustrated. The process is divided into several parts as mentioned above, and each part is rigorously studied with the detailed information stated in the following sections. The possible heat integration strategy and waste heat recovery in/between different parts are also proposed to produce electricity for using inside the coal to SNG plant. The concept of gasification-based process has been proposed earlier. In open literature, there are large amount of work performing the analysis energetically, economically as well as thermodynamically by simplified simulation studies. There are also works that studying each sub-part inside the plant-wide process by experiment or by modelling. However, the overall detailed analysis of the plant-wide process with rigorous modeling for each part inside the process is rarely found. Also, because of the flexibility in real operation, the operating condition for syngas treating section is determined based on optimization work which minimized the total annual cost, and there are several trade-offs between pipeline regulations and economical optimization. In this case, meeting the pipeline regulations becomes one that has higher priority. This contradiction is critical, but it was rarely studied before. After the design of each part is complete, the plant-wide economical evaluation is studied. The cost for producing SNG is then calculated and compared with the importation price of LNG in Taiwan. The

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primary aim of this work by the authors is to provide a baseline analysis for coal to SNG process for Taiwan, and any other country holds similar situation as Taiwan.

2. Design of Coal to SNG Process The overall block flow diagram for Coal to SNG process is illustrated in Figure 1. Firstly, coal is fed into gasifier with the high purity oxygen product from ASU, and is converted to raw syngas. The gasifier is operated at very high temperature and pressure, thus before sending to downstream, the raw syngas needs to be cooled by Radiant Syngas Cooler (RSC) and water quench. After cooling, the raw syngas enters the treating section, in which Sour Water-Gas-Shift-Reaction (SWGSR) is used to adjust the syngas composition, and Acid Gas Removal (AGR) process is used to capture H2S and carbon dioxide after syngas cooling. In SWGSR, the H2/CO mole ratio is adjusted to be higher than 3, due to the stoichiometric ratio of methanation reaction downstream. In AGR process, sulfur capture is desired to be nearly complete, because the catalyst used for downstream methanation reaction will be deactivated by sulfur containing components. The CO2 capture target is assumed to be 92%, this is because higher capturing target will increase the cost in AGR, which is not favorable, and 92% will be a reasonable assumption with existing capturing techniques.

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The last section for

this process is the methanation reaction. In this section, the treated syngas with the adjusted H2/CO ratio is reacted to form methane, which is the major component in SNG. After methanation reaction, there is a conditioning section to partly remove the small remaining CO2 and H2O by molecular sieve, and then pressurized the SNG product to meet the pipeline regulation.3-4 Besides, there is also a power block to generate steam by recovering waste heat from RSC and other cooling steps, and the generated steam is

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used to produce electricity by a steam cycle. In this work, Aspen Plus V7.3 is used for simulation. The major information and results for plant-wide process are described separately in the following subsections. Other detailed information and the simulation files provided by authors could be found in open access as the supporting information for the readers who are interested.

2.1.1 Coal Handling In coal handling section, coal is prepared to deliver into the plant. The coal is uploaded from the silo, dried, crushed, screened, and then transported to plant. In this SNG plant, the gasification system is operated by a slurry-feed, entrained flow gasifier, thus slurry preparation is necessary, while the pre-drying step is not. In slurry preparation step, the slurry medium, in this case, water, is added, and the coal is ground to the size that is allowed in entrained flow gasifier. Some additives are often added to reduce the viscosity of coal-water slurry, making it easier to transport to the gasification system. After the slurry is well-prepared, it is pressurized by a slurry pump, and is fed into gasifier. In this work, the processed coal sample is Kaltim-Prima coal (KPC) from Indonesia, which actually imported in Taiwan annually. The ultimate and proximate analysis data for KPC are listed in Table 1.7

2.1.2 Air Separation Unit In this study, cryogenic separation is used to produce large quantities of pure oxygen and nitrogen without argon separation. The block flow diagram is shown in Figure 2. The whole process can be divided into four blocks, including Main Air Compressor (MAC), further compression, Main Heat Exchanger (MHX), and rectifying

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section. In rectifying section, two columns, high pressure column (HPC) and low pressure column (LPC), are used to separate oxygen and nitrogen. The two columns are operating at different pressure, and the operating pressures are selected so that the reboiler of LPC and condenser of HPC are able to be integrated. For cryogenic air separation, it can be categorized into two types according to its operating pressure. They are elevated pressure ASU (EP-ASU) and low pressure ASU (LP-ASU). The operating pressure of HPC in EP-ASU is in range of 10-14 (atm), compared with that 4-7 (atm) in ASU. The ASU configuration and its operation condition for IGCC power plant have been studied in literature.11-13 From the energy-saving aspect, the EP-ASU is more appropriate to generate oxygen product that would be sent to gasifier which is operated at very high pressure. However, the separation efficiency of O2 and N2 becomes worse if the operating pressure is high. Compared with the oxygen product in IGCC plant (about 95 mol%) , the required purity for O2 in SNG plant is even higher. This is due to the fact that there are regulations for heat value in final SNG product, and the amount of inert compounds sent into gasifier and the downstream processes should be decreased3. Thus the oxygen purity from ASU is set to be 99 (mol%). For N2 product, the purity is set as 98 mol%. It can be used in recovering the physical solvent in AGR, as well as used in ammonia plant downstream the process which is not studied in this paper, Due to the high purity for both O2 and N2 product, LP-ASU is more suitable in this case. There are many configuration of ASU, and the performances were studied in literature.11 For conventional ASU, gaseous oxygen cycle (GOX) is often used if the products are desired at ambient temperature. In this kind of cycle, gaseous oxygen is taken out from the columns. For gasification-related system, the products from ASU have to be compressed to very high pressure. If GOX cycle operation is used, it will

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require large amount of electricity in compression. Another widely accepted concept is pumped liquid oxygen cycle (PLOX), in which liquid oxygen is taken out from the LPC. The liquid oxygen is then pumped, heated to become vapor in MHX, and compressed to the required pressure. PLOX cycle could have advantage in saving operating cost and capital cost, due to less electricity required in compressor. However, the refrigeration lost by taking out liquid oxygen has to be carefully recovered. This could be achieved by condensing a high pressure air stream by heat exchange with the vaporizing oxygen stream in MHX. The high pressure air stream could be obtained by further compression of a portion of the air feed that is already compressed by MAC. In that case, no extra air feed will enter the ASU and can avoid the increase of system loading.11-13 The overall process of LP-ASU with PLOX cycle is described below. The flowsheet and the simulation results of ASU are illustrated in Figure 3, with the main flows (air, nitrogen and oxygen related) labeled. The name of streams and units mentioned in the following text could also be found in fig.3. Air feed is first pressurized by MAC, and then cooled to room temperature. The compressed air is then split. The main stream (about 60~70%) is refrigerated by heat exchange with oxygen and nitrogen product in MHX, and is fed into the bottom of HPC (stream R-Air). The other air stream is further pressurized by the boosted air compressor (BAC), cooled, and then passes MHX to be further cooled to liquid (Stream L-Air). To enhance the nitrogen separation efficiency, a portion of liquid air is subcooled by heat exchanging with gaseous nitrogen at the top of LPC, and sent to the upper-middle section of LPC (stream L-Air-LP). The remainder of liquid air is sent to HPC (stream L-Air-HP). In order to reach the desired product purity, there is a recycle of liquid nitrogen taken from the top of the HPC. This recycle is split and sent back to HPC (stream HPC-R) and LPC

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(stream LPC-R), respectively, and the spilt ratio is dependent on the desired product purity and flowrate. In HPC, nitrogen and oxygen are separated to enrich the oxygen content at the bottom product. The bottom product from HPC is then cooled, and sent to the LPC for further separation (stream O2-I-LPC). Due to the fact that the boiling point of argon is in the middle of oxygen and nitrogen, and no argon separation in this case, the bottom product from HPC should be sent to the middle portion of LPC in order to reach the target purities for both O2 and N2 product. For products, liquid oxygen (stream L-O2) comes out at the bottom of LPC, and gaseous nitrogen (stream G-N2) comes out the top of LPC. These products are heated by exchanging heat with the air feed streams that are to be refrigerated and liquefied in MHX before going to downstream processes. The HPC is operated at 6.25 (atm), with 50 stages; while the LPC is operated at 1.3 (atm) and with 75 stages. The pressure drops for these two columns are 0.2 and 0.06 (atm), respectively.11 There are some design specifications for operation of ASU. The product purity of oxygen product is maintained at 99 (mol%) by adjusting the split ratio of the compressed air after MAC. The nitrogen product is maintained at 98 (mol%) by adjusting the split of liquid air (stream L-Air) into LPC and HPC. For better recovering of oxygen in HPC, the oxygen content is set to be less than 20 (ppm, mol%) in the overhead product. Finally, the total oxygen product is maintained at 5670 (kmol/hr) by adjusting the air feed flowrate, so that adequate amount of oxygen could be steadily provided into gasifer.

2.1.3 Gasification In pursuit of clean, and various energy uses, gasification technology has received much more attention in recent years. Essentially, gasification process converts coal or

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other fuels to syngas (mainly CO, H2, CO2 and H2O) through partial oxidation. The most promising characteristic for gasification process is that various kinds of fuel can be used. Among the kind of gasifiers, Entrained Flow Gasifier (EFG) is the most selected for gasification-based process nowadays, due to its high operating temperature (> 1300 ºC) and pressure (>30 atm), short residence time (95%).1, 14-16 EFG can be categorized as dry-feed or slurry-feed. For these two types, slurry-feed operation will result in lower Cold Gas Efficiency (CGE), this is because of the presence of slurry medium, typically water, provides a thermal load inside gasifier. Before gasification reaction occurs, all the slurry water should be vaporized, thus it leads to higher oxygen demand to combust and to provide the required heat, and meanwhile reduce CGE. However, dry-feed operation has several problems. First, the dry feeding system is quite expensive, and the cost will be increased largely as the operating pressure increases. Second, pneumatically transportation of the solid fuel often causes blocking.1 Operation difficulties are the major reason that slurry-feed operation is more favored. In this work, an industrial scale, slurry-feed, entrained flow, oxygen blown gasifier is used to convert coal into syngas. Yu and Chien17 have studied the performance of gasifier, and established a modified 1-dimensional model for simulation with appropriate assumptions. The model is validated with the technical data from Tampa Electric Company and DOE/NETL, and the simulation results are in good agreement with those data in open access. The model is capable of predicting the gasification performance of coal/water slurry gasification, coal/CO2 slurry gasification, and coal-biomass co-gasification. Thus the 1-dimensional model is used here to simulate the gasification performance, and the flowsheet is illustrated in fig. 4.

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In this work, it is assumed that the processed coal flow rate is 5760 (ton/day), with two gasifier trains operate in parallel. The coal/water slurry is at 63 wt%. The residence time inside gasifier is assume to be 3 seconds, the operating pressure is 56.8 (bar), while the coal particle is assume to be homogeneously spherical, with 80 (µm) diameter. All the reactions involved in gasification system are listed in Table 2.18-24 The high temperature syngas at the gasifier outlet needs to be cooled before downstream usage. Here in this paper, radiant quench method is utilized, in which a Radiant Syngas Cooler (RSC) is used to cool the syngas.3,9,

25-26

High temperature

syngas is slowly cooled by heat exchange with cooling medium, thus an amount of heat could be recovered as medium (or high) pressure steam. Due to the low cooling rate for radiant quench, some reversible reactions, mainly WGSR, is expected to continue even if the raw syngas has already left gasifier. The raw syngas is cooled to 593.3 (⁰C) in RSC as suggested in literature.25 For downstream uses, the syngas must be further cooled. Thus after RSC, there is a water quench unit to cool the syngas to its dew point. In this step, the low temperature quench water is sprayed, and having direct contact to syngas, which makes it cool quickly without any remaining reactions.3,9 After the water quench, the solvable components are washed out by a scrubber, with outlet temperature of syngas to be near 200 (⁰C).

2.1.4 Syngas Treating Section: SWGSR and Syngas Cooling The H2/CO mole ratio of syngas from coal-water slurry gasification is usually between 1 to 1.5, and it should be adjusted based on different downstream uses. Thus, to reach complete CO reaction, the H2/CO ratio is adjusted to higher than, or equal to the

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stoichiometric ratio of methanation reaction. In this work, the H2/CO ratio is adjusted to 3.5. There are two operating cases depending on the position of WGSR. The typical operating case is that the WGSR unit is in the middle of AGR. The sulfur content in syngas is first captured by the first stage of AGR, and the sulfur-free syngas is then passed the WGSR unit to adjust the H2/CO ratio. The treated syngas is then entered the second stage of AGR to capture CO2. The drawback for this operation is that AGR process is operated at temperature much lower than WGSR, thus there will be large amount of energy loss in changing temperature between those operations. A better improvement from the aspect of energy saving is that the WGSR comes before the two-staged AGR process, this operation is called Sour-Water-Gas-Shift Reaction (SWGSR), in which the sulfur-tolerant catalyst should be used. Besides, the sulfur content in coal is appeared in the form of carbonyl sulfide (COS) in syngas, it will also be hydrolyzed to hydrogen sulfide (H2S) in WGSR stage. In this work, SWGSR operation in accompany with COS hydrolysis is used to adjust the syngas composition. Typically, the SWGSR catalyst will also enhance the reaction rate of COS hydrolysis. Due to the fact that SWGSR and COS hydrolysis are both equilibrium limited with similar reactants, these two reactions are simulated in the same reactor. Mo-based catalyst, which will not deactivate under the presence of sulfur, is used to catalyze these reactions. The reaction kinetics for SWGSR is taken from open literature.24 Due to the lack of information about COS hydrolysis kinetics under SWGSR catalyst, another set of kinetic parameters under Al-based catalyst is used here.20 The sulfur content is very slight but not able to be neglected, thus the above

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compromise is believed to be adequate in describing COS hydrolysis performance in the reactor. SWGSR is equilibrium limited, with stoichiometric ratio between CO and H2O to be 1. Thus, feed with excess H2O will be able to enhance the reaction rate. In this work, the H2O/CO feed ratio is assumed to be 1.5, with enough steam added. The steam required here could be generate from the power block in the plant. Both SWGSR and COS hydrolysis reaction are exothermic, and the reactor is assumed to be tubular reactor operated in adiabatic mode. Due to the fact that our targeting H2/CO ratio is in the mid-way to the reaction equilibrium, thus the required reactor volume should be found. This could be viewed as an optimization problem with several constraints. The objective function is to minimize the reactor size, under the following constraint: (1) Length to diameter equal to 10, (2) Reactor outlet temperature less than the upper operating limit of catalyst (520 ⁰C), (3) H2/CO mole ratio at reactor outlet is equal to 3.5. The optimization work for SWGSR reactor is studied in the Aspen built-in SQL tool. Because WGSR is exothermic and the reactor is operated in adiabatic mode, the reacted syngas should be further cooled to the operating temperature in AGR. The cooling step is separated into three stages, part of the waste heat could be recovered by generating different levels of steam.9 The process condensate contains water and other soluble contaminants that removed from the syngas, which need to be further treated before disposed or recycled. The flowsheet and simulation results for SWGSR and syngas cooling are illustrated in Figure 5.

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2.1.5 Syngas Treating Section: AGR In this work, the commercially available physical absorption method, SELEXOL process, is used for Acid Gas Removal (AGR). In this method, a mixture of diethylene glycol is used as the solvent, which has much greater affinity to acid gas than other components.8-9, 27-28 In this work, H2S and CO2 are selectively captured in two stages. The captured H2S could be sent into Claus process for sulfur recovery, which is not considered in this work, while the captured CO2 is assumed to be compressed and sequestrated. For simulation, PC-SAFT is chosen as the thermodynamic model, with the binary interaction parameter cited from literature, and the aspen built-in DEPG component is used to represent the solvent.9,28 The capturing targets are assumed to be 99.7% for H2S and 92% for CO2. The overall flowsheet of AGR (dual-staged SELEXOL) process is illustrated in Figure 6, with the detailed information described in the following sections. The overall process for AGR is quite complicated, with many design variables to be determined. It can be roughly divided into three parts, H2S and CO2 absorbers (abbreviated as H2SABS and CO2ABS), regeneration of CO2-rich solvent, and regeneration of H2S-rich solvent. The syngas from previous sub-process, and the stripped gas recycled from H2S concentrator (H2SCON) are both sent to H2SABS. About 10% of the loaded, semi-lean solvent from CO2ABS is used to contact the acid gases for capturing. Due to the fact that the affinity of acid gas to DEPG depends on partial pressure, and the CO2 content in syngas is much higher than H2S, thus using the rich solvent from CO2ABS to contact with syngas will prevent the over-capture of CO2 inside H2SABS. The sulfur-free syngas from H2SABS, and the recycled gas from high pressure flash unit (HP-Flash) are sent to CO2ABS. Lean solvent recycled from

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regeneration stage is used to contact the gases to reach the desired CO2 capturing target. The cleaned syngas from CO2ABS is then ready to enter the downstream methanation process. For designing the absorbers, the determination of total solvent flowrate, and the split ratio of semi-lean solvent from CO2ABS are important. The total solvent flowrate is adjusted to let the CO2 capture reach 92%, while the split ratio of semi-lean solvent is adjusted to let the H2S capture reach 99.7%. DEPG has a boiling point of 275(⁰C), and the maximum operation temperature is around 175(⁰C).27 Also, DEPG solvent has better capturing ability at lower temperature and higher pressure. To avoid the deterioration of solvent and to reduce the energy required for solvent regeneration, water is often added along with DEPG into the system. According to the literatures, the maximum allowable water content in solvent is about 6 wt%.8-9 However, too much water content in the solvent will result in less capturing ability. It will need much more solvent to reach the capturing target for higher water content solvent, which will largely increase the operating cost and capital cost. Thus in this work, the solvent concentration is assumed to be 99 wt% DEPG and 1 wt% water, and the inlet temperature is refrigerated to 10 (⁰C). This will be reasonable assumptions for effectively capturing acid gases. The operating pressure of H2SABS is set at 51.7 (bar), while the CO2ABS is operated at 50.3 (bar). In both absorbers, pressure drops are calculated by the Aspen built-in tray rating function. The capturing efficiency for acid gas will be greater at higher pressure, thus the operating pressure for the absorber is set at the highest pressure allowable, with reasonable pressure drop assumption between units. The total stages of H2SABS and CO2ABS are listed as optimization variables, which will be discussed in the later part.

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The other part of semi-lean solvent from CO2ABS will enter the regeneration stage, in which the solvent and CO2 is separated by pressure drop in several flash units. In flash regeneration stage, three levels of pressure drop are used to strip the rich solvent. The high pressure flash unit (HP-Flash) is used to recover the H2 that absorbed by DEPG. The gas stream from HP-Flash is then sent back to CO2ABS, while the liquid is sent to other two flash units for CO2 separation. The medium pressure (MP-Flash) and low pressure flash unit (LP-flash) are used to separate the sequestration-ready CO2 products. The pressure level of the regenerative flash units should be determined as a trade-off between CO2 compression duty, regenerated solvent compression duty and solvent flowrate. For simplicity in design, those pressures are reasonably assumed to be 19/9/1.5 (bar). Under these assumptions, about 90% of H2 captured in CO2ABS could be recovered by HP-Flash. In this overall plant, SNG is the final product, thus less loss of H2 will be desired. The regenerated solvent from LP-Flash is pumped back to the upper section of CO2ABS, and the CO2 products from MP-Flash and LP-Flash units will be compressed to 152 (bar) for sequestration. In this work, multi-stage compressors with inter-stage cooling are used for simulating the compression steps. For regenerating solvent from H2SABS, the rich solvent is first contacted with stripping gas in the concentrator (H2SCON). In this case, high purity nitrogen from ASU is used as stripping gas. A portion of the absorbed H2S, and most of the absorbed CO2 are stripped. The stripped gas is then compressed, and recycled back to H2SABS. The solvent comes out from H2SCON has much higher composition in H2S, and H2SCON is necessary for producing Claus feed. The Claus process converts H2S into elemental sulfur, and the H2S content in Claus feed is better to be within 20~50 (mol%).10, 27 A feed with too less H2S into Claus process will cause its configuration to

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become more complicated, and the capital investment higher. Although the Claus process is not studied in this work, it is still reasonable to assume the H2S content in Claus feed to be higher than 30 (mol%) due to the fact that KPC is low in sulfur content (30.5 mol% in this work). The total number of H2SCON is assumed to be 12 in order to have adequate space for rich solvent and stripping gas to contact. The pressure of H2SCON is an optimization variable. The stripping performance is better under lower operating pressure, in which it requires less stripping gas to reach the desire target in Claus feed. However, the lower operating pressure makes the compression cost for recycle gas increase. Thus it is believed to have a trade-off in optimization of AGR. After H2SCON, the remaining H2S-rich solvent stream is then passed to a thermal regenerating column (TH-REG), in which the absorbed gas and solvent are totally separated. The lean solvent comes out from the bottom of TH-REG, which is recycled back to CO2ABS. The distillate from the column is first cooled to its dew point, and then passes a flash unit to condense the water. The vapor product from the flash is the feed to Claus process, while the condensate is sent back to TH-REG as a reflux. The condensate is almost all water, with slight amount of H2S and CO2 in ppm level. To both avoid water accumulating inside the column and also to maintain the H2S recovery at the target, part of the condensate is purged from the system. The purged condensate could be recycled for other uses after treatment. There are two specifications for maintaining the separation target. The composition at bottom is maintained at 99 wt% DEPG by adjusting the reboiler duty of the column for directly recycled back to CO2ABS. Another specification is that 99.95% of H2S is recovered at the claus feed, and it is adjusted by changing the purge ratio of condensate.

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In order to effectively and economically regenerate DEPG solvent, TH-REG is expected to operate at a vacuumed pressure at 0.2 (bar). Under this operating pressure, the reboiler could be heated by an industrial medium pressure steam (11 bar), and it can be generated by the recovered waste heat inside the plant-wide system. The total stage number for TH-REG is assumed to be 5, this is because of the least reboiler duty and ease of separation under this assumption from the preliminary testing. In optimization works, three design variables are selected to study, including the total stage of H2S absorber, total stage of CO2 absorber, and the pressure of the concentrator. The original works for optimization are studied based on English units, thus the English units are remained here in order to clearly illustrate the results. Some necessary unit conversion factors are also listed. The optimization algorithm is describing below: (1)

Guess a number of total stages for H2SABS (NTH2S).

(2)

Guess a number of total stages for CO2ABS (NTCO2).

(3)

Guess a H2SCON pressure (PCON).

(4)

Change the total DEPG recirculation rate, the ratio of semi-lean solvent split, and the stripping gas flowrate until all design specification can be met. The specifications are 92% overall CO2 capture, 99.7% overall H2S capture, and 30.5 mol% H2S in claus feed.

(5)

Go back to (3) and change a PCON until TAC is minimized.

(6)

Go back to (2) and change a number of NTCO2 until TAC is minimized.

(7)

Go back to (1) and change a number of NTH2S until TAC is minimized. The hierarchy of the iterative optimization procedure was arranged so that the

design variable for the outermost iterative loop is the most sensitive one in terms of

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TAC changes, while the innermost one affect the TAC the least. The optimization results are illustrated in Figure 7. As NTH2S increases, TAC will continue to decrease, this is probably because less semi-lean solvent is used to capture, which reduces the loading in the thermal-regeneration step. Thus, a reasonable upper limit for NTH2S is assumed to be 60 stages in this optimization study. The optimized condition from minimizing TAC occurs when NTH2S is 60, NTCO2 is 14, and PCON at 30.3 bar (440 psia). However, there is another limitation in stripping gas flowrate. The stripping gas will leave the AGR system with cleaned syngas, and it is acted as an inert in the methanation reaction section which will make the SNG heat value lower. The variation of required stripping gas flowrate and minimized TAC with changing PCON under fixed NTH2S equals to 60 is studied in Figure 8. The stripping gas flowrate for maintaining H2S composition in Claus feed is decreased with lower PCON. This is because the larger pressure difference between the H2SABS and H2SCON make the acid gas and solvent be separated more easily, thus less stripping gas is required. However, as PCON decreases, TAC increases, this is how the trade-off behaves. The methane concentration and SNG heat value related to amount of stripping gas are both illustrated in Figure 9. The inert of SNG includes nitrogen and argon, and nearly all nitrogen comes from AGR due to the fact that the ASU separates the oxygen to 99 mol% with the remained to be argon. Thus, for considering both TAC and the regulation of SNG, the chosen operating condition from this constrained optimization work is at NTH2S is 60, NTCO2 is 14, and PCON at 22.1 bar (320 psia). The SNG regulations are: CH4 > 95 mol%, HHV > 36 (MJ/sm3), which is described in detail in section 2.1.6.

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Some simulation results for the constrained, optimized case are described below. In H2SABS, 10.73% of semi-lean solvent from CO2ABS is used to capture H2S. In CO2-rich solvent regeneration stage, about 90.93% of CO2 will be recovered to sequestrate (8% in the cleaned syngas to downstream, and the remainder in Claus feed). Totally, the solvent recirculated inside the system is about 220 (ton/day), and about 132.43 (kmol/hr) stripping gas is required in H2SCON. Other detailed results are also illustrated in fig. 6 along with the AGR flowsheet.

2.1.6 Methanation Reaction Section In methanation reaction section, the treated syngas from AGR is reacted to form methane. The possible reactions are also listed in Table 2.29-30 For the SNG regulation in this work, the high heat value should be higher than 36 (MJ/sm3), the methane content should be larger than 95 mol%, and the CO concentration should be less than 100 ppm for the final SNG product. The pipeline pressure for SNG is 62 (bar), while the temperature is below 40 (⁰C). The overall process flowsheet is illustrated in Figure 10, in which a series of adiabatic plug flow reactors with inter-stage cooling is arranged. The overall process could be roughly divided into two stages, the first one is High Temperature Reaction (HTR) stage, and the other is the Low Temperature Reaction (LTR) stage, also called clan-up stage.4, 29-30 In this work, the Ni/MgAl2O4 catalyst is used in HTR stage. There are some characteristic for this catalyst. If the operating temperature is too low, carbonyl nickel will form at the surface of catalyst, which leads to deactivation. However, if the operating temperature is too high, sintering will occur. Thus, the feasible operating temperature for this catalyst is in the range of 350 to 600 (⁰C).29

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However, methanation reactions are exothermic, thus the reaction equilibrium is favored to the product side at lower temperature. Thus, another catalyst, Ni/Al2O3 catalyst, which can operate in the range of 280 to 400 (⁰C), is considered to be used in LTR stage.30 All the kinetic expressions are cited from literature. The

treated

syngas

from

AGR

is

first

heated

by

a

series

of

feed-effluent-heat-exchangers, and mixed with a cooled, compressed recycle stream from the first reactor outlet to reach 350 (⁰C). The methanation reactions are exothermic, thus the outlet for reactors in HTR stage should be cooled back to 350 (⁰C) before entering next reactor in order to make the overall conversion high. The reactions are mostly occurred in the first reactor, thus a large cooled outlet stream recycled from its outlet to the entrance after cooling could be helpful to avoid runaway inside reactor. After the third reactor, CO is nearly completely reacted, but there is still some amounts of H2 remained, which will be dangerous in storage and operation with SNG. Thus, at the outlet of the third reactor, a flash unit is used to separate water with gas before entering the LTR stage. This could drive the reaction between CO2 and H2 forward and is helpful for methane production. In the LTR stage, CO and H2 are reacted near completely to form methane. After the cleanup reactor, the produced syngas still contains some CO2 and H2O, thus another separation step is required. In this work, molecular sieve is used to separate 99% of remaining CO2 and all water inside the SNG product.3 The treated SNG is then compressed to the pipeline pressure, and cooled to required temperature. The composition of final SNG product is 95.5% CH4, 4.2% for combination of N2 and Ar, while CO is only 7 ppm. The high heat value of SNG product is 36.06 (MJ/m3), which matches the regulation.

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2.1.7 Power Block There is waste heat that could be recovered by generating different level of steam through a Heat Recovery Steam Generator (HRSG).9, 32 Most of the generated is used to drive steam turbine for producing electricity, while the remainder of them could be used in the SNG plant (i.e. steam used in WGSR and thermal-regeneration in AGR). In this work, a system of heat recovery and steam cycle is studied. The system is shown in Figure 11, and the waste heat to recover is listed in Table 3. In a typical steam cycle, the sub-cooled boiler feed water is heated by contacting the hot streams that is to be cooled, to make it become superheated steam. The temperature for the superheated steam is a design variable; the efficiency will become higher as it increases. Typically, this temperature will be less than 600 (⁰C), this is the operating limit for the equipment. In this work, this temperature is assumed to be 566 (⁰C) based on the assumption by NETL.3 For expanding step, the high pressure steam passes through a series of turbines, and four pressure levels are assumed and also illustrated in Table 3. The High-Pressure Steam (HPS) is expanded in the first turbine to become intermediate-pressure steam (IPS), and the IPS is reheated to 566 (⁰C) before further expanding, this step could enhance the power production efficiency. The reheated IPS is expanded in the second turbine to become low-pressure steam (LPS), and is then further expanded to become condensate to recycle back to system. In simulation, the isentropic efficiency of each turbine is assumed to be 87.5%. An important thing to notice is that the isentropic efficiency will influence the simulation directly and largely, thus the assumption should be made with care. In Table 3, the electricity demands inside the SNG plant are listed. As the table illustrated, ASU and AGR combines for over 80% of total electricity consumption. In

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ASU, the electricity is mostly used in the MAC, and in AGR, over 70% of electricity is used to compress CO2 up to the sequestration pressure. Thus, to improve the energy required in ASU and AGR will lead to benefits for the SNG plant. For other electricity utilization, such as electricity required in coal handling, slag handling, waste water treatment, and other miscellaneous uses, it is assumed to be 10% of total electricity requirement of the whole SNG plant. After those considerations, the net electricity generation for the system is at 29.151 (MW), this could be viewed as the bonus from this plant. In methanation section, there are also some waste heat streams that could be recovered at relatively low temperature. Those are also recovered by generating the industrial low pressure steam (LP-steam) for ancillary uses inside the plant. In economical evaluation, the LP-steam is also viewed as a bonus from the plant-side process.

2.1.8 Summary The simulation result for main streams connecting each part inside the plant-wide coal to SNG process is illustrated in Table 4, and calculation results for the energy efficiency is in Table 5. The overall energy conversion efficiency is at 60.38%; for energy that transferred from coal to SNG, it is at 58.66%.

3.

Economical Evaluation for Coal to SNG Process After the design and optimization for each part is studied, the economical

evaluation for SNG production is also necessary before the plant is actually established. Economical evaluation is critical, but the results will largely depend on the assumption

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of the financial structure, which will vary with different region. The economics for SNG production with different financial structure are studied in open literature.3, 33 In Taiwan, Chang et al. did a rough economical evaluation based on just correcting the fuel price with these economical models, and most of the result are roughly between 11.37~13.26 (USD/GJ).34 Here in this work, a complete economical evaluation with a reasonable financial structure is studied. The most commonly used unit to calculate the natural gas price over the world is USD/MMBTU, thus the SNG price is also converted to that unit in calculation for easier comparison. According to open literature, the items to be studied include capital investment (CAPIN), fixed operating and maintenance cost (Fixed O&M cost), and variable operating and maintenance cost (Variable O&M cost). Each part is explained in detail separately.3, 33-35

3.1 Capital Investment (CAPIN) In this study, the cost parameters and terminology used are based on the methodology developed from Electric Power Research Institute (EPRI).35 The input parameters, the calculation stage, and the items for calculating capital expense for coal to SNG plant are listed below, with other information required for calculation listed in Table 6. (1)

Direct construction cost. These are the costs for equipment material. The direct construction costs are provided in equation form. Chen regressed the empirical equation that calculated the direct construction cost for each part of IGCC plant without CO2 capture.35 For the parts that are not included in the discussed IGCC plant (i.e. dual-staged AGR process, Methanation, SNG purification), the direct

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construction costs are predicted based on empirical equations provided by Luyben.8 (2)

Indirect construction cost. The indirect construction cost included the cost for labor in installing the equipment, and other portion of capital expense that is not included.

(3)

General facility cost. This term includes the construction cost of roads, office buildings, labs, pipelines, electrical systems and so on.

(4)

Engineering & Home office fees. This is an overhead fee paid to the architect and engineering company.

(5)

Project Contingency Cost. The project contingency is a factor that covering the cost of additional equipment or other cost resulting from a more detailed design.

(6)

Process Contingency cost. This is referred to the cost for design uncertainty, and is always applied on an area-by-area basis. For the processes that are more mature (i.e. commercialized processes), the factor for process contingency cost will be smaller.

(7)

Royalty Fee. Royalty fee is a set amount of money that a business franchise owner must pay to be part of a franchise system.

(8)

Startup Cost. This part calculated the cost for pre-production, which is actually referred to the cost for operator training, equipment testing, extra maintenance, inefficient uses of fuels and consumables.

3.2 Fixed operating and maintenance cost (Fixed O&M)

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The items studied in this part are operating labor cost, maintenance cost, taxes and insurances, and administrative and support cost. The detailed information required for calculation is listed in Table 7.

3.3 Variable operating and maintenance cost (Variable O&M) For Variable O&M, the items studied are fuel price, maintenance material cost, consumable cost, and also the profit of by-products. Here in this work, the coal price sold by Tai-power company in Taiwan is used. The primary by-products are sulfur, electricity, and LP-steam from this SNG plant. However, KPC is coal sample with very low-sulfur content, thus the credits for selling sulfur is omitted in this work. The consumables included the cost for WGSR catalyst, methanation reaction catalyst, and DEPG solvent that used in AGR. All the information required for calculation is listed in Table 7.

3.4 Financial Structure The financial structure is very important to economical evaluation. Actually, the calculated SNG production cost is the required sold price to reach the target under the assumed financial structure. Those assumed targets are often in terms of payback period or internal rate of return (IRR). A stringent financial structure should take consideration on many things, such as the depreciation years and its type, rate of interests, inflation rate, discount rate, and cash flow. The product cost will be too high to be realizable under an overly strict financial structure. In this work, a simpler financial structure is assumed that only inflation rate and capital depreciation is accounted. It is assumed that the operation target is that the

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payback period for capital investment is 20 years, with 3.5% inflation rate during the operation period. All the capital investment will be linearly depreciated in 20 years. According to Gibsin Engineers, Ltd (GIBSIN), a Taiwan-U.S. joint venture consulting company, these assumptions will be corresponded to the investment regulation for public enterprises in Taiwan.36

3.5 Result for Economical Evaluation for Coal to SNG plant. The results for CAPIN, fixed O&M and variable O&M are listed in Table 8. For CAPIN, ASU, gasification, and syngas cleaning and treating process are combined for about 60% overall capital expenses. This is in agreement with the results from Chandel and Williams as well as DOE/NETL.3, 33 For variable O&M, the fuel price has the largest impact. Due to the fact that the coal in Taiwan depends on importation, thus the unit price is about 2.5 times higher than the coal price in U.S. The production cost is calculated in US dollar per unit heat value of SNG, thus the overall energy conversion efficiency will also influence the results of economical evaluation. Here in this work, the calculated SNG production cost is at 10.837 (USD/GJ). It is much lower than the LNG import price in Taiwan, which is 14~17 (USD/GJ). If this SNG plant could be co-operated with other country that has large amount of coal reserves, the SNG production cost will become even lower. Briefly, it seems to be economically attractive for Taiwan in the foreseeable future.

4.

Conclusion In this work, the plant-wide Coal to SNG process is studied. In this process, coal is

first converted to syngas through gasification, adjusted composition by WGSR, cleaned

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though AGR process, and reacted to form SNG through methanation reaction. A waste heat recovery system is also studied to produce different levels of steam, and then it is used in the SNG plant or to generate electricity for auxiliary use. From economical evaluation, the cost to produce SNG is at 10.837 (USD/GJ), which is much lower than the price to import LNG in Taiwan. Thus, to establish a SNG plant will have practically and economically advantages in the near future.

Acknowledgements The research funding from the National Science Council of R. O. C. under grant no. NSC 102-3113-P-002-014 is greatly appreciated.

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Table 1. Proximate and Ultimate analysis for Kaltim-Prima Coal Kaltim Prima Coal (KPC)

Proximate analysis

Ultimate analysis (DB*, wt%)

HHV (DB)

Moist

10.50

FC

51.58

VM

43.16

Ash

5.26

Ash

5.26

C

75.79

H

5.21

N

1.52

CL

0

S

0.66

O

11.56

(MJ/kg) (*DB; dry basis)

31.231

33



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Table 2. Kinetics for Reactions in each Section Reactions References Homogeneous Reactions CO + H2O↔ H2 + CO2 Kumar and Ghoniem (2012)18 (in gasifier) CO + H2O ↔H2 + CO2 Karan et. al (2002)19 (in RSC) COS+H2O ↔ H2S+CO2 Svoronous et. al (2002)20 C18H20+9 O2→18CO+10H2

Westbrook and Dryer (1981)21

Gasification

CO + 0.5 O2 →CO2

Kumar and Ghoniem (2012) 18

Section

H2 + 0.5 O2 → H2O

Kumar and Ghoniem (2012) 18

CO + 3H2 ↔ CH4 + H2O

Watanabe and Otaka (2006) 22

Heterogeneous Reactions

SWGSR Section Methanation Section

C+0.5O2 → CO

Wu et. al (2010) 23

C+H2O → CO+H2

Wu et. al (2010) 23

C+CO2 → 2CO

Wu et. al (2010) 23

C+2H2 → CH4

Wu et. al (2010) 23

COS+H2O →H2S+CO2

Svoronous et. al (2002)20

CO+H2O↔CO2+H2

Lund (1996) 24

CO+3H2↔CH4+H2O CO2+4H2↔CH4+2H2O CO+H2O↔CO2+H2

HTR: Xu and Froment (1989) 29 LTR: Kopyscinski et. al (2010) 30

34


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Table 3. Information for calculating energy balance Waste Heat to be Recovered for Generating Electricity Sources

Duty (MW)

Gasifier RSC

Temperature In Temperature Out (⁰C)

(⁰C)

95.970

1582.3

866.5

Methanation HPS

220.520

600

350.0

Syngas Cooling MPS

125.010

350

148.9

Syngas Cooling LPS

57.760

148.9

60

Steam Levels in Power Block HPS

125.1 (bar)

LPS

44.8 (bar)

IPS

34.5 (bar)

BFW

0.1 (bar)

Electricity utiliization Uses

(unit: MW)

ASU Total

69.728

AGR Total

36.868

Methanation Total

6.929

Power Block Utility

1.871

Other Uses

11.528

Total Electricity Uses

126.936

Total Electricity Generation

156.087

Net Electricity Generation

29.151

35



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Page 36 of 52

Table 4. Information for main streams inside the Process

65.6

O2 ASU 96.7

Gasifier Outlet 1309.1

RSC outlet 593.3

To WGSR 198.9

To Syngas Cooling 356.0

To AGR 35.0

To Methanation 13.7

To Claus 0.8

MP-CO2

LP-CO2

114.5

163.4

SNG Product 35.0

72.39

72.39

67.57

56.74

56.40

55.85

54.31

53.59

50.33

0.2

151.99

151.99

62.0

5760 (ton/day)

5083.6

5662.8

23393.0

23393.0

28875.9

32686.2

25472.1

16954.0

126.3

3201.4

4409.3

4216.0

-

0.0000

0.9900

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0006

0.0000

0.0000

0.0002

CO

-

0.0000

0.0000

0.4172

0.3563

0.2886

0.1099

0.1461

0.2104

0.0000

0.0071

0.0003

0.0000

H2

-

0.0000

0.0000

0.2738

0.3347

0.2711

0.3846

0.5115

0.7401

0.0000

0.0064

0.0001

0.0000

H2 O

-

1.0000

0.0000

0.1992

0.1383

0.3031

0.2383

0.0014

0.0001

0.0353

0.0007

0.0035

0.0019

CH4

-

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.9555

H 2S

-

0.0000

0.0000

0.0004

0.0004

0.0003

0.0012

0.0016

0.0000

0.3050

0.0000

0.0000

0.0000

N2

-

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0072

0.0402

0.0001

0.0000

0.0291

AR

-

0.0000

0.0100

0.0024

0.0024

0.0020

0.0017

0.0023

0.0031

0.0078

0.0013

0.0002

0.0124

CO2

-

0.0000

0.0000

0.0966

0.1575

0.1274

0.2586

0.3370

0.0390

0.6061

0.9844

0.9959

0.0010

COS

-

0.0000

0.0000

0.0013

0.0013

0.0011

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

NH3

-

0.0000

0.0000

0.0090

0.009

0.0064

0.0056

0.0000

0.0000

0.0000

0.0000

0.0000

0.0000

Coal

Water

T (⁰C)

65.6

P (bar) Flow rate (Kmol/h) O2

36


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Table 5. Energy Calculation for the Plant-wide Process (MW) Energy Input

Coal

1690.45

Energy Output

SNG

991.63

Electricity

29.15

Overall Energy Conversion

60.38 (%)

Coal to SNG Energy Conversion 58.66 (%) (1 MW = 3.412 MMBTU/h)

37



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Page 38 of 52

Table 6. Information for Capital Investment in Economical Evaluation Calculation Calculation Stage (1) Coal Handling (2) ASU (3) Gasification System (4)Syngas Cooling (5) WGSR (6) WGSR Items Studied

(7) Dual-Staged AGR and CO2 Compression (8) Methanation and SNG purification (9) Claus

Total Direct Cost

Process (10) HRSG and Steam turbine (11) Ash/spent sorbent handling Indirect Cost General Facility Cost Engineering/Office Fees Process Contingency Cost Project Contingency Cost Royalty Cost

20% Total Direct Cost 15% Total Direct Cost

Total Facility Cost

10% Total Facility Cost

5~10% of Total Facility Cost

15% of Total Facility Cost

0.5% of Total Facility Cost

6 months of operating labor

1 month maintenance materials, non-fuel consumables, and waste disposal at full

Startup Cost

capacity, 25% of 1 month’s fuel cost at full capacity

2% of total plant cost

38


Total Plant Cost

Total Capital Investment

Page 39 of 52

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Table 7. Information for Fixed and Variable operating and maintenance cost in Economical Evaluation Calculation Fixed Operating and maintenance cost Operating Labor Labor Burden

Insurances

18 engineers total 30% of total operating labor cost.

Maintenance Labor Taxes

Averaged to be 55000 (NTD/month), or 1550 (USD/month)

and

40% of total maintenance cost. (maintenance cost = 2.5% total capital investment) 25% of total plant cost. Variable Operating and maintenance cost

Fuel (KPC)

3.179 (USD/GJ)

Catalysts

1000 (USD/ton)

Chilled Water

4.43 (USD/GJ)

Electricity

16.8 (USD/GJ)

LP Steam

7.79 (USD/GJ)

DEPG Solvent

13.4 (USD/gal)

Maintenance

60% of total maintenance cost. (maintenance cost = 2.5% total

Material

capital investment)

39



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Page 40 of 52

Table 8. Result for Economical Evaluation Cost

Cost

(USD/GJ)

(USD/MMBTU)

3.700

3.907

Operators

0.022

0.023

Maintenance Labor

0.390

0.412

0.103

0.109

Taxes and Insurances

0.925

0.977

Total Fixed O&M

1.440

1.520

Coal Price

5.420

5.723

SWGS Catalyst

0.002

0.003

0.454

0.480

Total Capital Investment

Fixed operating and

Administrative and

maintenance cost

Support Labor

Chilled Water (SELEXOL+ASU) Variable operating and

Methanation Catalyst

0.036

0.038

maintenance cost

SELEXOL Solvent

0.134

0.141

Maintenance Material

0.527

0.556

LP Steam

-0.382

-0.405

Net Electricity

-0.493

-0.521

Total Variable O & M

5.697

6.015

Cost for SNG Production

10.837

11.443

Typical LNG Importation Price

14.21~17.05

15~18

(*1MMBTU = 1.05587 GJ)

40


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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure Captions Figure 1

Block flow diagram for plant-wide Coal to SNG process.

Figure 2

Block flow diagram for cryogenic air separation unit.

Figure 3

Simulation flowsheet for cryogenic air separation unit.

Figure 4

Simulation flowsheet for 1-dim, slurry-feed, entrained flow gasifier.

Figure 5

Simulation flowsheet for SWGSR and syngas cooling.

Figure 6

Simulation flowsheet for dual-staged AGR process.

Figure 7

TAC comparison for varying NTH2S, NTCO2, and H2SCON pressure.

Figure 8

Comparison of Stripping gas required and the minimized TAC with varying H2SCON pressure in AGR system at NTH2S = 60.

Figure 9

The methane concentration and SNG heat value related to amount of stripping gas.

Figure 10 Simulation flowsheet for methanation section. Figure 11 Simulation flowsheet for waste heat recovery and power block.

41



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Figure 1

42


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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 2

Air

Compression and Cooling

dehydration and Purification

Main Heat Exchanger

Further Compression

43


Rectification of O2 and N2


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Figure 3

44


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Figure 4

45



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 46 of 52

Figure 5

CW

LPS 148.9

60 ( C) 60

Waste Heat Streams (To Power Block)

35 ( C)

Steam 3810.2 (Kmol/h)

To AGR MPS To Syngas Cooling

To-WGSR (Raw Syngas)

355.6

148.9 ( C)

SWGSR Reactor

Flash1

Flash2

L=11.36 (m) D=1.136 (m)

Flash3 Process Condensate (To sour water treatment)

46


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Figure 6

47



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Page 48 of 52

Figure 7 NT (H2SABS) =60

35580

(b)

35560

35680

35540

35660

35520

35640

35500

NT (H2SABS) =55

35700

(a)

13

14

15

35620

16

NT (H2SABS) = 50

35880

13

14

15

16

NT (H2SABS) = 45

36100 (d)

(c) 35860

36080

35840 36060

35820

36040

35800

13

14

15

16

TAC Comparison for NT (H2SABS) =45, 50, 55, 60 (e)

36000 35800

13

P_Conc = 400 (psia) P _Conc = 420 (psia) P _Conc = 440 (psia) P_Conc = 460 (psia) **1 (bar) = 14.503 (psia)

35600 45

50

55

60

48


14

15

16

*For sub-figures (a) to (d): x-axis: NT(CO2ABS) y-axis: TAC, 1000(USD) *For sub-figure (e): x-axis: NT(H2SABS) y-axis: TAC, 1000(USD)

Page 49 of 52

Figure 8 200

35750

Minimized TAC (1000 USD)

190 180

STRGAS (Kmol/h)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


170 160 150 140 130 120 110 300

320

340

360

380

400

35700 35650 35600 35550 35500 300

P_H2SCON (Psia)

320

340

360

P_H2SCON (Psia)

49


380

400


Figure 9 0.965

CH4 Composition (mol %)

36.4 36.2

HHV (MJ/Sm3)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 50 of 52

36.0 35.8 35.6 35.4 100

120

140

160

180

200

0.960 0.955 0.950 0.945 0.940 100

120

STRGAS (Kmol/h)

140

160

STRGAS (Kmol/h)

50


180

200

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Figure 10 SNG Product MS CW1 0.97 (MW)

SNGCOOL 2.47 (MW)

ME-COMP2 2.00 (MW)

CO2/H2O

LP1 4.93 (MW)

ME-HX3

To Methanation ME-HX1

R4 L=3.50(m) D=1.75 (m)

ME-HX4

ME-HX2

ME-HX5

ME-COMP1 4.93 (MW)

LP2 8.00 (MW)

HEAT 3.77 (MW)

R2out

R1 L=10.0 (m) D=5.0 (m)

R3 L=1.92 (m) D=0.96 (m)

CW2 5.61 (MW) FLASH1

R1out-Rec H2O R1out HPS 220.50 (MW)

R2 L=2.87 (m) D=1.43 (m)

51


R3out

R4out


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Figure 11

52


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Design and Economic Evaluation of a Coal-to-Synthetic Natural Gas

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