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Sep 29, 2015 - synthetic natural gas (SNG) and ammonia are studied in this work. POLYGEN has been a widely studied topic recently, in which...
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Design and Economic Evaluation of a Coal-Based Polygeneration Process To Coproduce Synthetic Natural Gas and Ammonia Bor-Yih Yu and I-Lung Chien* Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan S Supporting Information *

ABSTRACT: The steady-state design and economic evaluation of a polygeneration (POLYGEN) process to coproduce synthetic natural gas (SNG) and ammonia are studied in this work. POLYGEN has been a widely studied topic recently, in which several products could be produced parallel at the same time. One of the two products in this study, SNG, has a composition and heat value very similar to those of typical natural gas, and can be used as a replacement in industrial and home usages. Another product, ammonia, is one of the most important inorganic chemicals in the world, and could be used as the precursor of various kinds of chemicals, as fertilizers, or as a cleaning agent. In the POLYGEN process, the relative production rates for different chemicals could be adjusted on the basis of different market demands, daily usages, and also changing political strategies. In our previous study (Yu, B. Y.; Chien, I. L. Design and Economical Evaluation of a Coal-to-Synthetic Natural Gas Process. Ind. Eng. Chem. Res. 2015, 54, 2339−2352), we illustrated that the SNG production price is lower than the liquefied natural gas importation price in Taiwan. The SNG production price is 10.837 USD/GJ (USD = U.S. dollars) in an SNG-only plant. With the POLYGEN process to coproduce SNG and ammonia, the SNG production cost could become even lower. If 20% of the syngas is used to produce ammonia, the SNG production price will drop to 9.365 USD/GJ, and if 40% is used for ammonia production, the SNG production price will drop further to 7.063 USD/GJ. Thus, although the POLYGEN process leads to an increasing total capital investment, it has positive influences from economic aspects. Besides, the flexibility of shifting the production rate of SNG or ammonia makes it possible to adapt to changes in the market demand.

1. INTRODUCTION The concept of the polygeneration (POLYGEN) process is to coproduce several different kinds of products in parallel, and along with CO2 capture. For each product, the relative production rate could be determined by market demands or other specific purposes if there are any. The origin of POLYGEN might come from the thought of energy storage.2 In an IGCC (integrated gasification combined cycle) power plant, syngas is used to produce electricity. The electricity demand fluctuates in a single day and also in different seasons; thus, the syngas needs a way to be stored once the electricity demand is low. However, the major components of syngas are CO and H2, which cannot be stored in a large amount easily. Thus, it can be used to produce other chemicals to help reach the goal. There are a lot of important chemicals that can be produced from syngas, such as alcohols, synthetic natural gas (SNG), ammonia, olefins, highcarbon liquid fuels, and so on.3,4 In recent years, the prices of the raw materials to produce these industrially important chemicals have risen and fluctuated, which causes the supply of these chemicals to become unstable. Thus, production of these chemicals through POLYGEN processes has been considered as a potential way to solve the problem in the past few years. Although the POLYGEN process will lead to higher capital investments, the fact that more kinds of chemicals can be produced will still enhance the flexibility for the plantwide operation and its economic performance. For this demonstration work, SNG and ammonia are chosen as the downstream products, and both are very important in Taiwan. Natural gas has long been considered as a cleaner energy source all over the world, due to its lower C/H ratio, and its © XXXX American Chemical Society

utilization has been very important in our country. In Taiwan, there are several natural gas power plants that provide electricity to local industries, with over 1300 MWe of electricity generated each year. However, over 99% of the energy source depends on importation. The recently developed shale gas has led to a dramatic drop in the natural gas price in the United States (4−6 USD/MMBTU (USD = U.S. dollars, and MMBTU = 1 million British thermal units) in 2010, 2.5−3 USD/MMBTU in 2015),5 but this does not benefit the Eastern Asian countries much, due to the requirement of liquefying and pressurizing and also the transportation distance. This leads to the fact that natural gas used in Taiwan still relies on importation (mostly from Qatar and Indonesia) in the form of liquefied natural gas (LNG). Although the prices of energy sources have dropped globally in recent years, the importation price of LNG has still been maintained at 11.20 USD/GJ in 2015, which is only an about 25% drop from that in 2013.6 Thus, lowering the cost of natural gas could have advantages in developing natural gas related processes as well as promoting the economics. In this work, a POLYGEN process to coproduce SNG and ammonia from coal is studied. Among the energy sources that are imported into Taiwan, coal has good advantages over the others, including a lower price, great abundance, and easier transportation. These reasons lead to the fact that coal is the major energy source used annually in Taiwan. The price for coal has Received: June 27, 2015 Revised: September 26, 2015 Accepted: September 29, 2015

A

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Figure 1. Block flow diagram for the plantwide POLYGEN process: (a) case 1, (b) case 2.

adopting carbon capture and storage can also be overcome by possible heat integration. In most of the works, a conceptual design with simplified modeling is included, but an overall plantwide process with rigorous simulation of each part is rarely found. Rigorous simulation is required to analyze the operating conditions, optimizations, and the possibility of scale-up. Thus, in this work, the steady-state design and the economical evaluation of a plantwide POLYGEN process to coproduce SNG and ammonia is illustrated, with each subprocess simulated rigorously. The process is divided into several subsections, including the air separation unit (ASU), gasification section, syngas treating section, methanation reaction section (SNG plant), ammonia synthesis reaction (ammonia plant), and power block. Among the subprocesses, the syngas conditioning step is the most important throughout the POLYGEN system. The configuration and operating conditions are flexible with different products produced. In this work, by producing SNG and ammonia, we proposed a modified configuration on the basis of our previous work.1 We clearly show the arrangement in the syngas conditioning step. Besides, in an ammonia plant, apart from the traditional process which uses an Fe-based catalyst (Habor−Bosch process), a novel Ru/C catalyst is assumed to be used to catalyze the reaction in this study. The Ru/C catalyst is more active than the traditional Fe-based catalyst at the moderate pressure, and operating at lower pressure will lead to better safety and economic performance.25−30 The catalyst is the core of ammonia synthesis, because it has a close relationship with the operating conditions and the related economic performance.25

also dropped along with that of LNG from 3.179 USD/GJ in 2013 to 2.674 USD/GJ in 2015.7 In our previous study, the SNG production cost was 10.837 USD/GJ in an SNG-only plant. After application of the new coal price, and a few corrections to our previous study,1 the SNG production cost will be 9.790 USD/GJ. Thus, the route that converts coal into SNG is still expected to have both practical and economical advantages in Taiwan. Ammonia is one of the most produced inorganic chemicals over the world, and it has a lot of uses. Most importantly, over 80% of ammonia is used in agriculture as fertilizers.8−10 For other uses, it is an important nutrient for organisms, or can be used as the raw material for producing urea or other related chemicals. aqueous ammonia can also be used to capture acid gases, as a refrigerant, or as a cleaning agent. Typically, ammonia is sold in the form of anhydrous liquid ammonia, with a purity of higher than 99.5 wt %.10 The normal boiling point of ammonia is −33 °C, which means that storage will require refrigeration or pressurization. There are research works that have studied the performance of different kinds of POLYGEN processes in the open literature, including transformation of coal to chemicals or liquid fuels or cofeeding of coal, biomass, natural gas, or coke oven gas to produce several products.11−24 Most of the works concentrate on energy or exergy analysis, thermodynamic performances, the optimization procedure, economic evaluation, and so on. The results from previous research indicate that POLYGEN exhibits better energy utilization and economic performances than a single-product plant. Also, the energy and cost penalty by B

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Industrial & Engineering Chemistry Research Detailed information on the different catalysts can be found in section 2.1.5. After the design of each part is complete, the plantwide economical evaluation is again studied, and compared with that of a plant that produces SNG only. The primary aim of this work is to provide a baseline analysis for the POLYGEN process for Taiwan, and any other country in a situation similar to that of Taiwan.

Table 2. Summary of the Main Units in the Plantwide Process section ASU gasification syngas conditioning

2. DESIGN OF THE POLYGEN PROCESS The overall block flow diagram for the process is illustrated in Figure 1. First, coal is fed into the gasifier with the high-purity oxygen product from ASU, and is converted to raw syngas. Because the gasifier is operated at very high temperature and pressure, before being sent downstream, the raw syngas needs to be cooled by the radiant syngas cooler (RSC) and water quench. After cooling, the raw syngas enters the treatment section, in which the syngas is split, and the ratio of H2 and CO of two split streams is adjusted to the required target. Also, almost all the sulfur contents and 92% of CO2 are captured here. The split syngas streams are then sent to the SNG plant and ammonia plant, respectively. Besides, there is also a power block to generate steam by recovering waste heat from the RSC and other cooling steps, and the generated steam is used to produce electricity by a steam cycle. The aim of this work is to design a plantwide process that can be operated at different chemical production rates. Thus, there are two design cases that could be viewed as two different operation scenarios. Under the scope of this work, we considered SNG to be the primary product. For case 1, 80% of the syngas is sent to the SNG plant, and for case 2, 60% of the syngas is used to produce SNG. In both cases, the remaining syngas is all sent to the ammonia plant. For the equipment sizing in all subprocesses, the case with larger throughput is considered, and detailed information on the sizing can be found in Table 1. After the

case 1 case 2

syngas to ammonia plant

80% 60%

20% 40%

ammonia synthesis

methanation

P = 22.40 bar P = 0.20 bar P = 50.71 bar P = 51.71 bar, Tin = 220.0 °C, Tout= 459.8 °C WGSR2 P = 51.02 bar, Tin = 120.0 °C, Tout= 459.8 °C reactor bed 1 P = 100.00 bar, Tin = 370.0 °C, Tout= 459.8 °C reactor bed 2 P = 99.71 bar, Tin = 370.0 °C, Tout = 437.9 °C reactor bed 3 P = 98.89 bar, Tin = 370.0 °C, Tout = 425.1 °C reactor 1 P = 47.92 bar, Tin = 350.0 °C, Tout = 600.0 °C reactor 2 P = 46.57 bar, Tin = 350.0 °C, Tout = 461.1 °C reactor 3 P = 45.43 bar, Tin = 350.0 °C, Tout = 374.9 °C reactor 4 P = 40.69 bar, Tin = 280.0 °C, Tout = 301.8 °C

coal (KPC) from Indonesia, which is actually imported into Taiwan annually. The ultimate and proximate analysis data for KPC are listed in Table A1 in the Supporting Information. 2.2. Air Separation Unit. In this study, the cryogenic separation method is again used to provide a large amount of high-purity oxygen and nitrogen without argon separation. The overall process is an LP-ASU with PLOX cycle, and the units in the process are described in detail in the following text. The separation strategy (i.e., LP-ASU and HP-ASU) and its configurations (i.e., GOX, PLOX) have been discussed in the literature.1,41 The configuration of the ASU is similar to that in our previous study, but with some modifications. The whole process can be divided into four blocks, including the main air compressor (MAC), boosted air compressor (BAC), main heat exchanger (MHX), and rectifying section. In the rectifying section, two columns, a high-pressure column (HPC) and a low-pressure column (LPC), are used to separate oxygen and nitrogen. The two columns operate at different pressures, and the operating pressures are selected so that the reboiler of the LPC and condenser of the HPC are able to be integrated. In this POLYGEN plant, a high-purity nitrogen product (ultrapure N2) is produced along with an oxygen product and the original nitrogen product.31 In the ammonia synthesis plant, the ultrapure nitrogen is needed because the presence of oxygen in the synthesis process will cause the catalyst to deactivate. Also, this ultrapure nitrogen product can be used as a stripping gas in the acid gas removal process and pressure swing adsorption (PSA) unit. This nitrogen product is withdrawn from the top of the HPC, at which the nitrogen purity is nearly pure. The other part of the distillate from the HPC still serves as reflux to itself and to the LPC, to increase the separation efficiency of oxygen and nitrogen. The overall process flowsheet is illustrated in Figure 2. For the HPC and LPC, the operation conditions and

Sizing Basis coal handling ASU gasification section syngas treatment ammonia synthesis methanation power block

temperature/pressure P = 1.32 bar P = 6.34 bar P = 56.74 bar, Tout = 1309.1 °C P = 56.40 bar, Tout = 593.3 °C P = 51.71 bar

H2SCON TH-REG CO2ABS WGSR1

Table 1. Detailed Information for Sizing of Each Subprocess syngas to SNG plant

name LPC HPC gasifier RSC H2SABS

case 1 = case 2 case 1 = case 2 case 1 = case 2 case 2 case 2 case 1 case 1

rigorous design, the plantwide economical evaluation is performed, and the results for coproduction are compared with those of our previous work, in which SNG is the only product. In this work, Aspen Plus V8.4 is used for simulation. The operating conditions of the major units are summarized as in Table 2, and the reaction information is included in Table A2 in the Supporting Information.29,31−40 The major information and results for the plantwide process are described separately in the following subsections. Other detailed information and the simulation files can be found in the Supporting Information for the readers who are interested. 2.1. Coal Handling. The coal sample and the other steps required in coal handling are assumed to be the same as those in our previous work.1 The processed coal sample is Kaltim-Prima C

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Figure 2. Simulation flowsheet for the cryogenic air separation unit.

Figure 3. Simulation flowsheet for the 1-dimensional, slurry-feed, entrained-flow gasifier.

comes out at the bottom of the LPC, and gaseous nitrogen (stream G-N2) comes out at the top of the LPC. All three products are heated by exchanging heat with the air feed streams that are to be refrigerated and liquefied in the MHX before going to the downstream processes. In the simulation, there are also some design specifications for operation of the ASU. The product purity of the oxygen product is maintained at 99 mol % by adjusting the split ratio of the compressed air after the MAC. The ultrapure nitrogen product is maintained at 99.9 mol % by adjusting the amount of distillate at the HPC that is sent back to itself as reflux. The ultrapure nitrogen flow rate is maintained at 2500 kmol/h by adjusting the amount of distillate at the HPC to be withdrawn. This flow rate will be adequate for ammonia

the design variables are assumed to be the same as those of the ASU in our previous study.1 Overall, there are three product streams provided in the ASU. One is ultrapure nitrogen (99.9 mol %), which is described above. Another one is pure oxygen (99 mol %); this product is sent to the gasifier as a gasifying agent. The high purity of oxygen is required because of the regulation of SNG downstream. Argon and nitrogen in the oxygen product will travel through the process and be present in the SNG product. If too much argon and nitrogen are present, the SNG will not reach the required pipeline heat value for direct use. The remainder is pure nitrogen (about 95.3 mol %), which could be used to promote pressure in the gas turbine, or in other ways. Liquid oxygen (stream L-O2) D

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Figure 4. Simulation flowsheet for the syngas treatment step.

previous study. The detailed information on the gasification reaction and also the pros and cons of different kinds of gasifiers have also been discussed in the open literature.1,3,4 2.4. Syngas Treatment Section. A lot remains to be discussed regarding the arrangement of subprocesses in the plantwide POLYGEN process, and the focus is mostly on the configurations of the syngas treatment section. In this section, the syngas produced from the gasifier is purified and split, and the composition of the split syngas streams is adjusted to the required value for producing chemicals. The whole section can be divided into two parts, teh water−gas shift reaction (WGSR) and acid gas removal (AGR). In AGR, H2S and CO2 are captured by the SELEXOL method. The advantages and potentials of the SELEXOL method have been discussed in previous literature.1,42−44 The characteristics of different kinds of configurations for the syngas treatment section were also discussed in our previous work.1 From the research of the U.S. Department of Energy National Energy Technology Laboratory (DOE/NETL),11 the syngases for producing different chemicals were produced independently, in which two WGSR sections and two dualstaged SELEXOL acid gas removal process are required. In this work, a modified process configuration with one dual-staged SELEXOL process based on our previous work is assumed. The simplified configuration could reach the treatment target, as well as decrease the process complexity and investment cost. The overall flowsheet of the syngas treatment process is illustrated in Figure 4, with the major design and operating variables of the units summarized in Table 2. Basically, it can be

synthesis in case 2, in which the ammonia production rate is larger and requires more nitrogen. Finally, the total oxygen product is maintained at 5670 kmol/h by adjusting the air feed flow rate, so that an adequate amount of oxygen can be steadily provided into the gasifer. Throughout the ASU, the Peng− Robinson equation of state is used to describe the thermodynamic behavior of the system. 2.3. Gasification. In this work, an industrial-scale, slurryfeed, entrained-flow, oxygen-blown gasifier is used to convert coal into syngas. The 1-dimensional model established in our previous work is again used to simulate the gasification performance, and the flowsheet is illustrated in Figure 3.1 Throughout the gasification section, the Redlich−Kwong−Soave equation of state (RKS) method is selected to represent the thermodynamic behavior of the system. It is assumed that the processed coal flow rate is 5760 tons/day, with two gasifier trains operating in parallel. The coal/water slurry is at 63 wt %. The residence time inside the gasifier is assumed to be 3 s, the operating pressure is 56.8 bar, and the coal particle is assumed to be homogeneously spherical, with an 80 μm diameter. All the reactions involved in the gasification system are listed in Table A2 in the Supporting Information. At the gasifier outlet, an RSC is used to cool the syngas to 593.3 °C.1 A water quench unit is used to cool the raw syngas to its dew point for further downstream uses. After the water quench, the solvable components are washed out by a scrubber, with the outlet temperature of the syngas near 200 °C. In the gasification section, all the operation variables, the related assumptions, and the model used are the same as those in our E

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Industrial & Engineering Chemistry Research ⎛ −76800 ⎞ 0.13 ⎟( P ) rCO = (4.3 × 105) exp⎜ (PH2O)0.49 ⎝ RT ⎠ CO (PH2)−0.45 (PCO2)−0.12 (1 − b)

divided into four sections, COS hydrolysis, H2S capture, WGSR (including syngas cooling and water condensation), and CO2 capture, with WGSR lying between H2S capture and CO2 capture. The specifications and thermodynamic model selection of each section are described in the following. For COS hydrolysis, 99.9% COS conversion to H2S is assumed. For AGR, 99.7% sulfur recovery and 92% CO2 recovery are assumed. For two-staged WGSR, 98% CO conversion is targeted. For simulation, the Peng−Robinson equation of state is chosen for WGSR and ELECNTRL is chosen for syngas cooling. PC-SAFT (perturbed-chain statistical associating fluid theory) is chosen as the thermodynamic model for AGR, with the binary interaction parameter cited from the literature, and the Aspen built-in DEPG component is used to represent the solvent.1,42 After the gasification process, the sulfur content is mostly in the form of H2S, with some portion in the form of COS. However, COS has less affinity with the SELEXOL solvent compared with H2S. If COS enters the downstream turbine, it will possibly lead to SO2 formation through combustion. Thus, to avoid this, COS is hydrolyzed to form H2S before sulfur capture. After COS hydrolysis, the raw syngas stream is cooled to 35 °C, and most of the water content inside is condensed out. Because two syngas streams with different compositions are needed in the POLYGEN process, part of the syngas will bypass the WGSR, and this stream is finally mixed with the other stream that passes the WGSR and AGR with very high CO conversion and decent CO2 capture. Thus, sulfur capture should be the first step of syngas treatment to avoid sulfur contents being sent to the downstream chemical plant, and causing catalyst deactivation. The AGR process (including H2S capture and CO2 capture) is modified from that of our previous work. In H2S capture, the cooled raw syngas and the stripped gas recycled from the H2S concentrator (H2SCON) are both sent to the H2S absorber (H2SABS). A small portion of the loaded, semilean solvent from the CO2 absorber (CO2ABS) is used to contact the acid gases for capturing. Here, 99.7% of H2S from the inlet stream is captured. The H2S-rich stream from the absorber is first stripped by contact with pure nitrogen in the H2SCON, and then is thermally stripped in the TH-REG. The overhead product from the THREG is ready to feed into the Claus plant for sulfur recovery. The sulfur-free syngas from the H2SABS is first heated by heat exchange with the water−gas shift reactor outlet to 204 °C, and then passes into a zinc oxide bed for deeper sulfur removal. The zinc oxide bed is common equipment for sulfur removal, and the separation ability is up to 99.9% overall capture. However, due to the fact that ZnO is hard to regenerate, it is restricted to a slight amount of sulfur removal.11 Here we assume the H2S removal in the ZnO bed reaches 99.9% from the stream after H2S absorption. The reaction equation for WGSR is CO + H2O ↔ CO2 + H2. It is an exothermic and an equilibrium-limited reaction, with a stoichiometric ratio between CO and H2O of 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 generated from the power block in the plant. The reactor is assumed to be a tubular reactor operated in adiabatic mode. Due to the fact that our target CO conversion is higher than the equilibrium conversion at a given temperature, a two-staged reaction is needed. The first stage is high-temperature shift (HT-Shift), with the kinetic expression illustrated in eq 1. The kinetic information is cited from Germani and Shurman.37

b=

PCO2PH2 KeqPCOPH2O

⎛ 4577.8 ⎞ Keq = exp⎜ − 4.33⎟ ⎝ T ⎠

(1)

The second stage is low-temperature shift (LT-Shift), with the kinetic expression listed in eq 2. The kinetic information is cited from Choi and Stenger,38 with the same definition of parameter b, and the same value of the equilibrium constant (Keq) as in HTShift. ⎛ −47400 ⎞ ⎟P P (1 − b) rCO = (2.96 × 105) exp⎜ ⎝ RT ⎠ CO H2O

(2)

Because the reaction is exothermic and is operated in an adiabatic reactor, an intermediate cooler is needed to lower the temperature for LT-Shift. The cooling could be used to provide the energy for generating steam in the power block. The stream after the WGSR is then cooled in two stages. The water content in the syngas is again condensed. The cooled syngas from the WGSR and syngas cooling stage and the recycle stream from the high-pressure flash (HP-Flash) unit are then combined, entering the CO2ABS at the bottom. The gas makes close contact with the lean solvent DEPG in the CO2ABS, which enters at the column top. After close contact between the gas stream and the solvent, 92% of CO2 is captured, and the solvent comes out of the CO2ABS at the bottom as a rich solvent. The rich solvent from the CO2ABS is split. A small portion enters the H2S capture stage, which is described above. The remainder will enter the regeneration stage, in which the solvent and CO2 are separated by pressure drop in several flash units. In the flash regeneration stage, three levels of pressure drop are used to strip the rich solvent, which are HP-Flash, medium-pressure flash (MP-Flash), and low-pressure flash (LP-Flash). HP-Flash is used to recover the H2 in the rich solvent, which is sent back to the absorber. MP-Flash and LP-Flash are used to separate out the sequestration-ready CO2, which is then compressed to 152 bar to be sequestrated. There are also some design specifications in AGR. First, H2S removal in the H2SABS is targeted at 99.7% by adjusting the split ratio of rich solvent from the CO2ABS. Second, the CO2 capture in the CO2ABS is targeted at 92% by adjusting the required lean DEPG solvent flow rate that enters the CO2SABS. Third, the stripping nitrogen flow rate is adjusted to let the SNG final product meet the pipeline requirement ([CH4] > 95.5 mol %). This nitrogen stream will travel with the syngas to the methanation reaction and be present in the SNG final product; thus, the determination of the required flow rate of stripping nitrogen should be based on the SNG pipeline regulation. Because H2S capture in this process is now in front of the WGSR, the extent of CO2 capture in the H2SABS will be less. This leads to higher purity of H2S (about 50 mol %) compared to that of our previous case (about 30 mol %) in the Claus feed. Thus, the required stripping nitrogen flow rate that enters the H2SCON could also be decreased. Other major design variables of this section (summarized in Table 2) are based on our previous work. In this work, the Claus plant is not studied, but the cost and energy consumption of this part are still included. F

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Figure 5. Simulation flowsheet for ammonia synthesis.

The final step of the syngas treatment section is to adjust the H2/CO mole ratio, and the flow rate into the SNG plant and ammonia plant. Thus, the syngas after AGR is split again. One part is combined with the gas that bypasses the WGSR and AGR. The mixing target is a H2/CO mole ratio of 3.5 to serve as the methanation feed. There is still some amount of CO2 inside the methanation feed, but no further CO2 capture is needed. This is because CO2 could react with H2 to form methane. The direct reaction between CO2 and H2 to form methane is very slow at high temperature; however, CO2 and H2 could react to form CO via the reverse WGSR. By properly designing the system to remove water from the process, the methanation conversion could be very decent. Another split stream from the AGR outlet gas should be further purified before being sent to the ammonia plant. This is because the remaining oxygen-contained species will be a potential gasifying agent for causing the oxidation of the graphite support in the Ru/C catalyst. Thus, a pressure swing adsorption (PSA) unit using nitrogen as the purging gas is used for further purification. Here, the PSA unit is modeled on the basis of literature data, by assuming that 94% of H2, 35% of Ar, and 4% of N2 are recovered.45 2.5. Ammonia Synthesis Process. Ammonia synthesis was developed in the early 1900s (known as the Habor−Bosch process), and has become a mature industrial process. For almost all ammonia plants, the reaction is catalyzed with an Fe-based catalyst.8,9,46,47 However, for the Fe-based catalyst, the catalysis performance at low- to medium-pressure operation is not as good as that at high pressure. Also, the produced ammonia will provide a kinetic inhibition of the catalyst, which causes the conversion to be far lower from the equilibrium conversion. These reasons lead to the fact that typical ammonia synthesis processes were operated at very high temperature (>400 °C) and very high pressure (>200 bar). The operation at such extreme conditions will lead to higher maintenance cost and also higher risks of industrial accidents. For the ammonia synthesis, the catalyst is the heart of it all. It affects the capital expense, operating expense, performance, and the safety of the process.25 Thus, many research groups have spent a lot of effort in improving the catalyst performances recently, especially in the lower pressure region. Among the research in the open literature, a Ru-based catalyst could meet the target, by providing over 20 times the activity of the Fe-based catalyst at lower pressure.26−29 However, ruthenium is a noble

metal, and its scarcity makes it much more expensive than the Febased catalyst.30 The price issue causes the Fe-based catalyst to continue to play a lead role in ammonia synthesis. Despite the much higher catalyst cost, it is still considered that the use of the Ru-based catalyst could lead to a lower cost for the system. The KAAPplus (Kellogg advanced ammonia process) is the first highpressure ammonia synthesis process that makes ammonia with the aid of a Ru-based catalyst. The successful operation of the KAAPplus paved for the way for the Ru-based catalyst to become a possible candidate for a more efficient, much safer, and less energy intensive ammonia synthesis in the future.25 In the ammonia synthesis process, the reaction is N2 + 3H2 ↔ 2NH3. In this work, a novel Ru/C catalyst is used to catalyze the reaction, and the information on the reaction kinetics is cited from Rossetti et al.29 rate = k f λ(q) (a N2)0.5 (a H3)0.375 (a NH3)−0.25 −

1 (a )0.75 (a H2)−1.125 K eq NH3

1 + K H2(a H2)0.3 + K NH2(a NH3)0.2 a N2 =

1 − ηβ(q) Pγ 1 + q − 2ηβ(q) N2

a H2 =

q − 3ηβ(q) Pγ 1 + q − 2ηβ(q) H2

a NH3 =

2ηβ(q) Pγ 1 + q − 2ηβ(q) NH3

⎛ −23000 ⎞ ⎟ k f = (9.02 × 108) exp⎜ ⎝ RT ⎠ ⎛ 4529.22 ⎟⎞ K H2 = exp⎜ −6.844 + ⎝ T ⎠ ⎛ 3522.73 ⎟⎞ K NH3 = exp⎜ −4.176 + ⎝ T ⎠ (3)

In eq 3, α is the defined activity of the specific component, kf is the pre-expenential factor, and KH2 and KNH3 are absorption constants for H2 and NH3. λ(q) and β(q) are parameters that depend on the H2/N2 mole ratio inside the system, in which λ(q) G

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Industrial & Engineering Chemistry Research Table 3. Status of Variables in Ammonia Synthesis variable

status

R1 inlet T R2 inlet T R3 inlet T R1 inlet H2/N2 ratio AMMSEP recycle T R1/R2/R3 length

DS/vary DS/vary DS/vary optimization optimization optimization DS/vary

R1/R2/R3 diameter HPS generation

fixed DS/vary

ammonia product purity

DS/vary

specification value 370 °C 370 °C 370 °C X X X enough for reaction equilibrium 4m max heat exchange duty for HX3 99.5 wt %

varies R-SPL1 ratio R-SPL2 ratio R-SPL3 ratio X refrigeration temperature exchange duty for HX1 reactor bed length

variation of optimization variables and related constraint

2.0−2.4 R1 outlet T < 460 °C 120−140 °C

X exchange duty of HPS generation pressure of liquid ammonia after decompression

stream from the flash contains a lot of unreacted H2 and N2, and is recycled back to the reaction section. The liquid stream from the flash is rich in ammonia, and is decompressed and heated and then passes through another flash unit for further separation. The liquid stream from the second flash is the final ammonia product, and the vapor stream could be used as a fuel gas to provide some energy. If the ammonia purity does not satisfy the target (99.5 wt %), a series of absorbers and strippers with water as the absorbent can be used for further purification again. Also, if the system contains too much inert gas, such as CH4 and Ar, a small portion of the vapor recycle stream could be purged out from the system to avoid its accumulation. However, because there is a PSA unit for purifying the H2-rich feed, the inert amount inside the system is already very low. Therefore, neither the absorber−stripper series for further purification nor the purge stream is necessary. There are many design variables to be determined in the ammonia synthesis process. Some of them could be used to specify other variables, some of them are assumed to be reasonable values, and the remaining variables are viewed as optimization variables. The variables and their status are listed in Table 3. The objective function for optimization is to minimize the total annual cost (TAC), with the assumption that the payback period is 3 years. From the preliminary studies, the TAC decreases as the system pressure increases. Although the increase in system pressure causes the compression cost to rise, the enhanced reaction rate decreases the required reactor volume and the catalyst loading amount. Because the Ru/C catalyst is much more expensive than the traditional Fe-based catalyst, a decrease in catalyst loading can lead to a better economical performance (Ru/C catalyst, 116.86 USD/kg; Fe-based catalyst, 2.20 USD/kg).30 Also from our preliminary studies, the TAC decreases as the reactor bed inlet temperature increases. This results from the fact that chemical equilibrium favors the products at lower temperature. Due to the limitation of equilibrium conversion, operating at lower temperature leads to higher per pass conversion, which effectively decreases the loading of the system, and this leads to less utilities required. Although the lower reaction temperature results in a lower reaction rate and also longer reactor, the increment in capital cost can be overcome by the decrease in utility cost. Thus, from our preliminary work before optimization, we assume the system pressure is at 100 bar, and the inlet temperatures for each reactor bed are set at 370 °C. For each reactor bed, the length of the reactor is assumed to be enough for reaching chemical equilibrium. In the calculation, we divide the plug flow reactor into 10 equal segments, and the criterion for equilibrium is that

affects the reaction rate directly, and β(q) affects the reaction equilibrium. When H2/N2 is equal to 3.0, λ(q) and β(q) are both equal to 1.0. When H2/N2 is equal to 1.5, λ(q) is 1.2 and β(q) is 0.5. The λ(q) value at H2/N2 = 1.5 is refitted by the authors, while the remaining parameters in the equation set are suggested by Rossetti et al.29 In their work, the reaction experiments were conducted at temperatures between 370 and 460 °C and pressures of 70−100 bar; thus, the synthesis process should be within these operation ranges. For the Ru-based catalyst, the most distinct feature is that H2 acts as an inhibitor of the catalyst in the chemical absorption step; thus, the operation at a nonstoichiometric ratio with N2 in excess is preferred. However, from a thermodynamic point of view, operation with a nonstoichiometric ratio has a negative effect on the reaction equilibrium. Therefore, the ratio between H2 and N2 will be an important variable and should be carefully determined. However, as the reaction goes on, the H2/N2 ratio will continue to drop once it is less than 3. Thus, in this work, linear interpolation of the parameters λ(q) and β(q) is used to express the kinetic behavior at the condition in which the H2/N2 ratio is between 3 and 1.5. The process flowsheet is illustrated in Figure 5, and can be roughly divided into two sections, the reaction section and the purification section. In the reaction section, a feed-quench-type process configuration is applied.8,9 H2 and N2 feeds are first compressed to the system pressure (100 bar), and mixed with the recycled, unreacted reactants. The combined feed is then split into four streams. One of them passes the feed effluent heat exchanger (FEHE) to increase the temperature by exchanging heat with the outlet of the last reactor bed (R3). This heated split stream is then mixed with another, and enters the first reactor bed (R1). The reaction is exothermic; thus, the reactor outlet needs to be cooled before the stream is sent into the next reactor bed (R2). Cooling is reached by mixing the cold, split stream with the reactor outlet directly. At the second reactor outlet, the remaining split stream is mixed to cool it, and then sent to the third reactor bed (R3). The outlet from R3 is cooled by passing through the FEHE, and sent to the purification section. In the purification section, the outlet from the reactor section is first cooled by passing through two heat exchangers. One of them cools the reactor outlet by heating the recycle stream, and the other further cools it by heating the ammonia liquid that separated from the flash unit downstream. After passing through the heat exchangers for heat recovery, the reactor outlet is then further cooled by using chilled water, to the temperature at which ammonia is capable of being separated by a flash unit. The vapor H

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Figure 6. Optimization results for ammonia synthesis: (a) TAC, (b) first reactor outlet temperature.

the reactor temperature variation ratio is less than 10−3 (i.e., (T10 − T9)/T9 < 10−3). Generally, only three variables are left in the optimization work; they are the H2/N2 ratio at the R1 inlet, the ammonia separation (AMMSEP) ratio, and the recycle temperature (TREC). The importance of the H2/N2 ratio has already been discussed above. The ammonia separation ratio is defined as the portion of ammonia in the reactor outlet to be recovered by flash as liquid ammonia. Because the reactor outlet is a vapor stream, more ammonia could be recovered at lower temperature. The unrecovered ammonia is sent back to the reaction section with the unreacted species in the vapor recycle stream. However, the recycled vapor ammonia could act as a thermal carrier in R1, in which the reaction proceeds much more widely than in the other

rector bed. Thus, one constraint that holds in the design is that the ammonia separation ratio should keep the R1 outlet temperature below the upper limit of the catalyst operation temperature range. Last but not least, the recycle temperature will affect the split ratio of four streams because there are inlet temperature specifications for each reactor bed. Thus, it is also listed as an optimization variable. The optimization procedure is described below: (1) Guess a value of the H2/N2 ratio. (2) Guess a value of the AMMSEP ratio. (3) Change the T-REC to minimize the TAC under the satisfaction of the ammonia purity specification. The ammonia purity specification is 99.5 wt %. I

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Figure 7. Simulation flowsheet for the methanation process.

Table 4. Information for Calculating the Energy Balance Waste Heat To Be Recovered for Generating Electricity POLYGEN case 1 source

duty (MW)

gasifier RSC methanation HPS ammonia synthesis WGSR intermediate WGSR bypass AGR input condensing

191.932 174.684 28.098 25.154 7.910 151.086

HPS IPS

POLYGEN case 2

temp in (°C)

temp out (°C)

1315.6 593.3 600.0 350.0 284.8 135.9 409.1 301.2 204.1 60.0 199.1 60.0 Steam Levels in the Power Block 125.1 bar 34.5 bar Electricity Utilization (MW)

duty (MW)

temp in (°C)

temp out (°C)

191.932 129.060 55.404 34.042 5.845 151.086

1315.6 600.0 283.0 409.1 204.1 199.1

593.3 350.0 135.0 280.6 60.0 60.0

LPS BFW

4.48 bar 0.1 bar

use

POLYGEN case 1

POLYGEN case 2

ASU total AGR total methanation total ammonia total power block utility other uses total electricity uses total electricity generation net electricity generation

81.704 31.490 4.160 5.186 1.625 18.625 142.790 150.153 7.363

81.704 35.880 2.763 10.735 1.834 19.997 152.854 139.802 −13.052

recorded. From the figure, the operation at a H2/N2 ratio of 2.2, ammonia separation ratio of 0.53, and recycle temperature of 120 °C has the lowest TAC. This operation point also matches the requirements that the outlet temperature at the first reactor bed is less than 460 °C. Therefore, this result is then adopted as the case for final economic evaluation.

(4) Go back to step 2 and change the AMMSEP ratio until the TAC is minimized. (5) Go back to step 1 and change the H2/N2 ratio until the TAC is minimized. The optimization results are illustrated in Figure 6, in which the TAC and the outlet temperature of the first reactor bed are J

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Figure 8. Simulation flowsheet for the power block.

2.6. Methanation Reaction Section. In the methanation reaction section, the treated syngas from AGR is reacted to form methane. The possible reactions are also listed in Table A2 in the Supporting Information.39,40 The process configuration is similar to that in our previous work, in which four adiabatic plug flow reactors in series with intermediate coolers are adopted.1 The flowsheet of the methanation process is illustrated in Figure 7. The SNG regulations are listed in the following: the high heat value should be higher than 36 MJ/sm3, the methane content should be larger than 95.5 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 °C. There are some adjustments between this new work and our previous work in the methanation part. First, there is more carbon dioxide in the methanation feed in this new case, because it contains a syngas stream that does not undergo CO2 capture. This could have some benefits for the entire system. The reaction between CO2 and H2 is very slow; thus, CO2 inside the system could act as a thermal carrier due to the fact that methanation reactions are all exothermic. This leads to a result in which a smaller portion of the R1 outlet needs to be recycled to maintain the outlet temperature within the feasible catalyst operation temperature range. It can also decrease the loading of the recycle compressor. Second, less stripping gas enters the AGR in POLYGEN cases comparing with the previous case that only

produces SNG. This could enhance the concentration of methane in the SNG product, and also lead to a higher heat value. For comparison with the previous results, assume 98.5% of CO2 is removed by molecular sieve in the final SNG treatment stage to maintain the heat value around 36 MJ/sm3. This is a reasonable assumption and will not cause the CO2 amount to exceed the required regulation for SNG (less than 2%).11 2.7. Power Block. The process configuration for the power block is also similar to that in our previous study, and the key concept is again briefly mentioned here.1 In the steam cycle, the subcooled boiler feedwater (BFW) is first pumped and heated to the furnace temperature to become a superheated high-pressure steam (HPS), and the large amount of waste heat inside the system can be recovered by heating the BFW. The superheated HPS is then expanded by turbines to generate electricity through three pressure levels (HPS, intermediate-pressure steam (IPS), low-pressure steam (LPS)), and the pressure levels are listed in Table 4. After the low-pressure expansion turbine, the steam becomes a partial vapor−liquid condensate at the same pressure as the BFW. This condensate is then further condensed, to close the steam cycle in the power block. A reheating step of the IPS is also helpful for enhancing the overall electricity generation efficiency. The steam required in the WGSR and that used to heat the thermal stripper (TH-REG) in AGR can also be produced here. Some low-temperature waste heat could also be used to generate the industrial medium- or low-pressure steam. This K

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Industrial & Engineering Chemistry Research Table 5. Energy Calculation for the Plantwide Process a

energy input (MW ) energy output (MWa)

SNG-only plant1

POLYGEN case 1

POLYGEN case 2

1690.41 0 1020.78 991.63 62.11 0 58.66 3.67

1690.41 0 1015.19 785.51 7.36 223.85 59.71 0.44

1690.41 10.790 1023.49 582.47 0 441.02 60.55 −0.64

coal electricity total SNGb electricity ammoniab

HHVchemicals/HHVcoal (%) electricity/HHVcoal (%) a

1 MW = 3.412 MMBTU/h. bHHV for methane, 55.50 MJ/kg; HHV for ammonia, 22.5 MJ/kg.

gasifier outlet is first chilled to a very low temperature for H2S recovery, and after that the stream is heated again for the WGSR. After the WGSR, it is chilled again to achieve CO2 capture. The repeated chilling and heating steps probably cause energy to be lost.

medium- or low-pressure steam can be used in other plants that only require a moderate heat source nearby. Thus, it is viewed as another bonus from the system. The flowsheet of the power block can be found in Figure 8. In Table 4, the waste heat information, the pressure levels in the steam cycle, and the electricity demands inside the POLYGEN process for case 1 and case 2 are all listed. In the POLYGEN process, the loading of the ASU is higher than that in the SNG-only plant, because an ultrapure nitrogen stream is needed for ammonia synthesis. This leads to a 17.2% increase of the electricity demand in the ASU for POLYGEN cases. However, because there is a portion of syngas that bypasses the WGSR and AGR, the CO2 compression cost is less in the POLYGEN cases. In comparison with the SNG-only plant, a 14.5% decrease and a 2.7% decrease of the electricity demand in AGR could be found in case 1 and case 2, respectively. Generally, the ASU and AGR still combine 75∼80% of total electricity consumption in both cases. For ammonia synthesis, there is 3.7% and 7.2% overall electricity consumption for case 1 and case 2, respectively. The electricity demand in ammonia synthesis is in compressing the H2 and N2 feed, and also the recycled unreacted vapor stream. Also, for other electricity utilization, such as electricity required in coal handling, slag handling, wastewater treatment, and other miscellaneous uses, it is assumed to be 15% of the total electricity requirement of the whole plant. After those considerations, the net electricity generation for the system is 7.363 MW for case 1 and −13.052 MW for case 2. The results mean that, in case 1, there is some electricity that could be viewed as a bonus from the system, but in case 2, for reaching energy balance, some electricity input is needed. 2.8. Summary. The simulation results for the main streams connecting each part inside the plantwide POLYGEN process for both cases are illustrated in Tables A3 and A4 in the Supporting Information. The calculation results for the energy efficiency are given in Table 5. The energy conversion efficiency (HHVchemicals/HHVcoal) is defined as the ratio between the total heat value from the chemical products and the total input heat value. From Table 5, the energy conversion efficiency for the SNG-only plant (58.66%) is a bit lower than that for the POLYGEN plant (59.71% for case 1 and 60.55% for case 2). The result is in correspondence with those in the open literature that the POLYGEN process can have better energy conversion efficiency. For two POLYGEN cases, case 2 has a higher energy conversion efficiency, and this is because producing ammonia can lead to better energy utilization. For the SNG-only plant, 3.44% (electricity/HHVcoal) of the heat value in coal is transferred to electricity as a bonus. For POLYGEN case 1, only 0.44% of the heat input is transferred to electricity, and for case 2, some electricity is input into the system. The difference between these cases is that, in the POLYGEN process, the

3. ECONOMICAL EVALUATION OF THE POLYGEN PROCESS 3.1. Information Required in Economical Evaluation. After the design and optimization of each part is studied, the economical evaluation of SNG production is also necessary before the plant is actually established. The results for economical evaluation of the POLYGEN process are also compared with those for the SNG-only plant. The financial structure suggested by Gibsin Engineers, Ltd. (GIBSIN), a Taiwan−United States joint venture consulting company, is considered in the evaluation work. The operation target is that the payback period for capital investment is 20 years, with a 3.5% inflation rate during the operation period. The items to be studied are also divided into three parts, capital investment (CAPIN), fixed operating and maintenance cost (fixed O&M), and variable operating and maintenance cost (variable O&M). The calculation strategies are the same as in our previous studies. In this work, the POLYGEN plant is designed as a plant that operates with different scenarios, with different product throughputs in different time periods due to changing market demands or political strategies. Thus, for capital investment, the plant must be designed to be capable of shifting the production rate of SNG and ammonia flexibly. Thus, in economical evaluation of each subpart, the one with larger throughputs in the two cases are considered. The details for consideration in economical evaluation are illustrated in Table 1. Other required information for calculating CAPIN, fixed O&M costs, and variable O&M costs is collected in Tables A5 and A6 in the Supporting Information.1,48 According to the data provided by Taipower Co., the leading company for importing coal and the internal supply, the coal importation price has dropped from 3.179 USD/GJ in 2013 to 2.674 USD/GJ in 2015. The import price of coal is the most sensitive factor in the economical evaluation. Thus, to compare the results of the evaluation, the results of our previous study are corrected by the new coal unit price. On the other hand, the global selling price of ammonia has a lot to do with the natural gas price. As the natural gas price goes down, so does the ammonia price. From the report of ICIS published in 2014,49 the spot price of ammonia is about 0.50−0.52 USD/kg in Taiwan. Also, ammonia storage often requires a huge, thick, two-layer tank, refrigeration-related equipment, and other ancillary equipment. For a conservative purpose, when calculating the income, we assume that the ammonia selling price will drop to 90% of the L

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Industrial & Engineering Chemistry Research Table 6. Results of Economical Evaluation total capital investment fixed operating and maintenance cost

variable operating and maintenance cost

revenue from other products

operators maintenance labor administrative and support labor taxes and insurances total fixed O&M coal price WGSR catalyst methanation catalyst ammonia catalyst chilled water refrigerant ZnO SELEXOL solvent waste disposal maintenance material electricity total variable O&M electricity steam ammonia

cost for SNG production (USD/GJ) cost for SNG production (USD/MMBTUa) IRR (%) typical LNG importation price (USD/MMBTUa) a

SNG-only plant1

POLYGEN case 1

POLYGEN case 2

3.940 0.022 0.372 0.098 0.985 1.477 4.560 0.002 0.007 X 0.454 0.044 0.000 0.134 0.071 0.502 X 5.774 −1.051 −0.350 X 9.790 10.336 12.04 11.30−11.70

5.258 0.030 0.497 0.132 1.314 1.973 5.734 0.007 0.007 0.166 0.542 0.085 0.000 0.135 0.089 0.671 X 7.438 −0.157 −0.358 −4.789 9.365 9.888 12.84

6.990 0.041 0.658 0.175 1.748 2.622 7.762 0.009 0.010 0.440 0.844 0.177 0.000 0.151 0.121 0.889 0.311 10.714 X −0.485 −12.778 7.063 7.457 12.95

1 MMBTU = 1.05587 GJ.

can just be overcome by the revenue from selling ammonia. If the production rate of ammonia becomes larger (case 2), the economic performance of the plantwide process will become even better. Due to the fact that, in Taiwan, natural gas is a very important energy source but we lack it, in this work, we consider SNG as a primary product, and more than half of the syngas produced from gasification is used to produce SNG. The consideration of case 1 and case 2 illustrates the condition of different production rates in a single plant as time varies. The evaluation result shows no matter how and when we shift the throughput of SNG and ammonia between these two scenarios, the economic performance will be better than a plant that only produces SNG. Besides, one more advantage of the POLYGEN process is that two or more products can be produced from a single process. This could provide a variety of energy sources, and also effective energy storage. For our country and other countries relying on energy source importation, the POLYGEN process will continue to be attractive economically and practically in the foreseeable future.

spot price in January 2014, along with a 2−3% selling cost. The ammonia storage cost is calculated to be about 0.06 USD/kg according to the literature.50,51 Thus, a selling price of 0.38 USD/ kg for ammonia in this evaluation work is a reasonable assumption. 3.2. Evaluation Results and Comparisons. The results for CAPIN, fixed O&M, variable O&M, and IRR (internal rate of return) are listed in Table 6. The most commonly used unit to calculate the natural gas price over the world is U.S. dollars per MMBTU; thus, the SNG price is also converted to that unit in the calculation for easier comparison. For CAPIN, the ASU and gasification section combines for about 50−55% of the total capital expense. Another 30% of the total capital expense is used in the syngas treatment section and chemical plants. For variable O&M, the fuel price still has the largest impact. After correction for the coal price, the SNG production cost is 9.790 USD/GJ for the SNG-only plant in our previous research. For POLYGEN plant case 1, the SNG production cost is 9.365 USD/GJ, and for POLYGEN case 2, the SNG production cost is 7.063 USD/GJ. From the perspective of the IRR, it is revealed that the IRR for the POLYGEN plant is higher than that of the SNG-only plant, which could also imply the better investment enticement from shareholders. In comparison with the dropping LNG importation price (11.09 USD/GJ), the SNG-only process is still economically attractive, but if SNG is coproduced with ammonia, the economical performance becomes even better. Although the inclusion of the ammonia plant leads to higher capital investment, the value that the ammonia product brings in could still have a positive effect on plantwide economical performances. However, if only 20% of the total syngas is used to produce ammonia (case 1), the benefits that ammonia brings are not apparent. This is because the increase in capital investment

4. CONCLUSION In this work, the plantwide, coal-based POLYGEN process to coproduce SNG and ammonia is rigorously studied. As the results from this work show, the multiproduct characteristic of the POLYGEN configuration can be more economically attractive than a single-product plant. If 20% of the syngas is used to produce ammonia, the SNG production price will drop to 9.365 USD/GJ, and if 40% of the syngas is used for ammonia production, the SNG production price will further drop to 7.063 USD/GJ. Both of the prices are lower than that calculated from the SNG-only plant, 9.790 USD/GJ. Thus, although the POLYGEN process leads to increasing total capital investment, M

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PLOX = pumped liquid oxygen cycle POLYGEN = polygeneration PSA = pressure swing adsorption RSC = radiant syngas cooler SNG = synthetic natural gas TAC = total annual cost TH-REG = (in ASU) thermal regeneration column T-REC = (in ammonia plant) recycle temperature variable O&M = variable operating and maintenance cost WGSR = water−gas shift reaction

it still has positive influences from the economic aspects. Due to the fact that over 99% of the energy sources in Taiwan depend on importation, and also the cheaper price and easier transportation of coal over LNG, the route to convert coal into SNG and other kinds of chemicals will continue to be economically attractive in the near future.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.5b02345. Tables providing more information on the POLYGEN process (PDF) Simulation files of each subprocess created with ASPEN V 8.4 (ZIP)

■

Capital Letters

Keq = equilibrium constant Ki = adsorption equilibrium constant (i = H2 or NH3) P = system pressure (bar) Pi = partial pressure (i = species) R = ideal gas constant (=8.314 J/(mol K)) Ti = temperature of ammonia reactor bed in different sections (i = 9, 10) (°C) T = temperature (K)

AUTHOR INFORMATION

Corresponding Author

*Tel.: +886-3-3366-3063. Fax: +886-2-2362-3040. E-mail: [email protected].

Lowercase Letters

b = equilibrium quotient calculated in WGSR kf = pre-exponential factor in ammonia synthesis reaction q = H2/N2 mole ratio in ammonia synthesis reaction r = reaction rate (kmol/(m3 s))

Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The research funding from the National Science Council of ROC under Grant No. MOST 103-2221-E-002-257 is greatly appreciated.

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Greek Symbols

NOMENCLATURE

■

Acronyms

ASU = air separation unit AGR = acid gas removal AMMSEP = (in ammonia plant) ammonia separation ratio BAC = (in ASU) boosted air compressor BFW = (in power block) boiler feed water CAPIN = capital investment CO2ABS = (in AGR) CO2 absorber FEHE = (in ammonia plant) feed effluent heat exchanger fixed O&M = fixed operating and maintenance cost GOX = gaseous oxygen cycle HHV = higher heat value H2SABS = (in AGR) H2S absorber H2SCON = (in AGR) H2S concentrator HP-ASU = high-pressure ASU HPC = (in ASU) high-pressure column HP-Flash = (in AGR) high-pressure flash unit HPS = (in power block) high-pressure steam HT-Shift = (in WGSR) high-temperature shift reaction IGCC = integrated gasification combined cycle IPS = (in power block) intermediate-pressure steam IRR = internal rate of return (%) KPC = Kaltim-Prima coal LNG = liquefied natural gas LP-ASU = low-pressure ASU LPC = (in ASU) low-pressure column LP-Flash = (in AGR) low-pressure flash unit LPS = (in power block) low-pressure steam LP-Shift = (in WGSR) low-pressure shift reaction MAC = (in ASU) main air compressor MHX = (in ASU) main heat exchanger MP-Flash = (in AGR) medium-pressure flash unit

α = defined activity in ammonia synthesis reaction kinetics (bar) β = defined parameter in ammonia synthesis reaction kinetics γ = fugacity coefficient of species λ = defined parameter in ammonia synthesis reaction kinetics

REFERENCES

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DOI: 10.1021/acs.iecr.5b02345 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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DOI: 10.1021/acs.iecr.5b02345 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX