Efficiency of Gas-to-Liquids Technology with Different Synthesis Gas

Feb 5, 2014 - The purpose of this work is to study the effect of different methods of natural gas-to-syngas conversion on the GTL efficiency. .... In ...
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Efficiency of Gas-to-Liquids Technology with Different Synthesis Gas Production Methods Ilya S. Ermolaev,*,†,‡ Vadim S. Ermolaev,†,‡ and Vladimir Z. Mordkovich†,‡ †

Department of New Chemical Technologies and Nanomaterials, Technological Institute for Superhard and Novel Carbon Materials, 7a Centralnaya Street, Troitsk, Moscow, 142190, Russia ‡ INFRA Technology Ltd., 11-3B Mokhovaya Street, Moscow, 125009, Russia ABSTRACT: The design and optimization of a gas-to-liquids technology (GTL) is considered, mostly from the view of an optimal choice of a synthesis gas (syngas) production method. Material balance and energy efficiency are calculated for a number of flowsheets, which comprise syngas production by various techniques and Fischer−Tropsch (FT) synthesis of liquid hydrocarbons (LHs). Three different methods are considered for syngas production: combined steam and dry reforming; autothermal reforming; and partial oxidation. The FT process is considered in the version, which produces LHs, not waxes. The results of modeling manifest that the choice of syngas production method influences dramatically the efficiency of overall GTL technology. Ways to improve the efficiency of GTL technology are discussed.

1. INTRODUCTION

A combined steam and dry reforming (SDR) method as a version of steam reforming is the most common method for syngas production. The injection of carbon dioxide into the steam reformer allows the control of the H2/CO ratio. It was shown elsewhere6 that the efficiency of SDR-including GTL can be controlled by recirculation technique of FT tail gases, which allows acquisition of a high energy efficiency of the GTL process. However, since recirculation itself represents an extra burden on both capital and operational costs, it is important to consider alternative methods of syngas production. The alternative methods are represented by more recently developed autothermal reforming (ATR)1,8,9 and partial oxidation (POX),1,10,11 which are used by the most recent industrial GTL plants ORYX GTL by Sasol-Chevron and Pearl GTL by Shell, respectively. The purpose of this work is to study the effect of different methods of natural gas-to-syngas conversion on the GTL efficiency.

The gas-to-liquids (GTL) technology for liquid hydrocarbons (LHs) production on the basis of Fischer−Tropsch (FT) process includes two main stages: (1) Conversion of methane to produce syngas (mixture of CO and H2); (2) Fischer− Tropsch synthesis to produce liquid hydrocarbons from the syngas. The great majority of GTL versions existing in industry or in literature imply also the stage of hydrocracking of heavy waxes produced by the FT process.1 The thermal efficiency of such GTL processes is discussed in the literature.2−4 In particular, it is shown that application of different types of syngas production technology matters, as well the use of recirculation of the FT tail gas.3 It is shown that the thermal efficiency of the conventional GTL technology (FT+hydrocracking) may reach as high as 67%.4 The present work, however, is devoted to the analysis of performance of a “shortcut” version of GTL, where the FT process produces lighter fractions of liquid hydrocarbons (LH) directly, so there is no need for hydrocracking.5,6 This technique can significantly reduce the capital cost of the entire process. However, the energy efficiency of the “shortcut” version of GTL is more difficult to control, so the contribution of syngas production into GTL efficiency becomes more important.7 Syngas production technology is well developed and available commercially in many versions, so it may seem the easiest part in constructing the GTL flowsheet. However, it is the most capital-intensive stage, the cost of which is up to 60% of the construction costs of a GTL plant. Selection of technological solutions for the production of syngas is usually determined by the quality of the natural gas to be converted, as well as by the required H2/CO molar ratio. For FT synthesis, the optimum molar ratio of H2/CO is 2.0− 2.2.1,7 This ratio can be obtained using various methods of conversion of hydrocarbon gases. © XXXX American Chemical Society

2. MODELING AND PROCESS DESCRIPTION Flowsheets for the GTL process with three different syngas production methods are shown in Schemes 1 to 3. Simulation of these flowsheets was done using mathematical models, which allow a consideration of all main steps of this technology. The composition of natural gas (NG) in mole basis is as follows: CH4, 79.64%; C2H6, 7.8%; C3H8, 4.0%; C4H10, 2.23%; C5H12, 2.73%; CO2, 0.07%; H2S, 0.02%; CH3SH, 0.01%; N2, 3.4%; H2, 0.1%. This composition is typical for petroleum associate gases of Western Siberia. The first step in all flowsheets is desulfurization, which is done through hydrogenation of organic sulfur compounds by Received: July 17, 2013 Revised: December 19, 2013 Accepted: January 27, 2014

A

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Scheme 1. GTL process flowsheet with syngas production via combined steam and dry reforming (SDR): (1) desulfurization unit; (2) pre-reformer; (3) main reformer (SDR); (4) tubular furnace; (6) water separator; (8) unit for carbon dioxide (CO2) extraction from flue gases (FG); (9) unit for carbon dioxide (CO2) extraction from syngas (SG); (10) unit for hydrogen (H2) extraction from syngas (SG); (11) Fischer−Tropsch reactor (FT); (12) separator of endproducts; (13) stabilization unit of end-products; (NG) natural gas; (LHs) liquid hydrocarbons; (SC) small circulation loop of tail gas; (BC) big circulation loop of tail gas; (SBC) stabilized blowdown circulation loop.

Scheme 3. GTL process flowsheet with syngas production via partial oxidation (POX): (1) de-sulfurization unit; (2) pre-reformer; (3) partial oxidation reactor (POX); (4) fired heater; (5) reactor of water gas shift reaction; (6) water separator; (7) air separation unit (ASU) for oxygen (O2) extracting; (9) unit for carbon dioxide (CO2) extraction from syngas (SG); (10) unit for hydrogen (H2) extraction from syngas (SG); (11) Fischer−Tropsch reactor (FT); (12) separator of end-products; (13) stabilization unit of endproducts; (NG) natural gas; (LHs) liquid hydrocarbons; (SC) small circulation loop of tail gas; (BC) big circulation loop of tail gas; (SBC) stabilized blowdown circulation loop.

Scheme 2. GTL process flowsheet with syngas production via authothermal reforming (ATR): (1) de-sulfurization unit; (2) pre-reformer; (3) autothermal reformer (ATR); (4) fired heater; (6) water separator; (7) air separation unit (ASU) for oxygen (O2) extracting; (9) unit for carbon dioxide (CO2) extraction from syngas (SG); (10) unit for hydrogen (H2) extraction from syngas (SG); (11) Fischer− Tropsch reactor (FT); (12) separator of end-products; (13) stabilization unit of end-products; (NG) natural gas; (LHs) liquid hydrocarbons; (SC) small circulation loop of tail gas; (BC) big circulation loop of tail gas; (SBC) stabilized blowdown circulation loop.

(1) Hydrogenation: CH3SH + H 2 ↔ CH4 + H 2S

(1)

(2) Absorption: ZnO + H 2S ↔ ZnS + H 2O

(2)

After desulfurization the residual sulfur content does not exceed 0.5 ppm. The purified gas is mixed with superheated steam in a molar ratio of H2O/C = 0.4,1 heated up to a temperature of 520 °C, and fed to an adiabatic prereformer (item 2, Schemes 1−3). The adiabatic prereformer ensures oxidation of heavier hydrocarbons in the reaction:1 CnHm + nH 2O = nCO + (n + m /2)H 2

(3)

Thus the heavier hydrocarbons are converted into a mixture of H2, CO, and CH4. The reaction runs with heat absorption. Temperature at the outlet of the prereformer depends on the temperature and composition of the mixture at the inlet and is in the range of 400−450 °C. The resulting mixture is fed to the main reformer (item 3, Schemes 1−3) for producing syngas of a required composition. Further on we consider three cases separately: (1) Combined steam and dry reforming (SDR) is carried out through methane oxidation by water and carbon dioxide at a constant temperature in the following reactions:10−12 CH4 + H 2O ↔ CO + 3H 2

hydrogen and consequent removal of hydrogen sulfide by solid sorbent. The hydrogen is added to the natural gas in the amount of about 5%.12 This stream is heated up to the temperature of 400 °C and goes to the desulfurization unit (item 1, Schemes 1−3). The necessary hydrogen is extracted from syngas in a membrane unit (item 10, Schemes 1−3). The process of extraction of sulfur compounds is thus carried out by chemisorption purification in two steps:12

CH4 + CO2 ↔ 2CO + 2H 2

CO + H 2O ↔ CO2 + H 2

(− 206.4 kJ/mol) (− 248.3 kJ/mol)

(+ 41.0 kJ/mol)

(4) (5) (6)

A superheated steam is added before the main reformer so the molar ratio of H2O/CH4 = 2.5/1.10,12 Carbon dioxide is also added to the stream before the main reformer in order to control the molar ratio of H2/CO in a produced syngas. This B

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mixture is heated to 650 °C and goes to reaction tubes of the main reformer (item 3, Scheme 1). The heat required for the endothermic conversion reaction is obtained by the combustion of tail gas of FT synthesis in the tubular furnace (item 4, Scheme 1). If the tail gas is not enough, an additional portion of the natural gas is burned. The temperature of converted gas at the outlet of the reaction tubes is 900 °C. The temperature of the flue gases passing from the radiation chamber of the tubular furnace to the convection chamber is taken as equal to 1150 °C, which allows proper calculation of required amount of gas supplied to the burner. (2) Autothermal reforming (ATR) is carried out through methane oxidation by oxygen and water on the following reactions:1,10,11 CH4 + 1.5O2 → CO + 2H 2O

CH4 + H 2O ↔ CO + 3H 2 CO + H 2O ↔ CO2 + H 2

(+ 519 kJ/mol)

(− 206 kJ/mol) (+ 41 kJ/mol)

by an aqueous solution of methyldiethanolamine (MDEA).15 The purified gas leaves the absorber of this unit with the concentration of carbon dioxide at about 300 ppm and goes to the hydrogen extraction unit, where the molar ratio of H2/CO in syngas is reduced to the desired value of about 2.2. In the case of SDR the extracted carbon dioxide goes back to the main reformer for syngas production. Since the carbon dioxide extracted from syngas is not enough to provide a H2/ CO ratio of 2.2, extra CO2 is extracted from flue gases in unit 8, Scheme 1. The resulting syngas is fed to the FT reactor (item 11, Schemes 1−3) at a pressure of 22 bar and temperature of 150 °C. This reactor is a shell and tube apparatus, in which a cobalt catalyst loaded into a tubular space provides the conversion of syngas into a mixture of hydrocarbons and water in the following reactions:1,16

(7) (8)

nCO + (2n + 1)H 2 → CnH 2n + 2 + nH 2O ( +160 kJ/mol)

(9)

The stream from the prereformer is heated to a temperature of 650 °C and goes to the autothermal reformer (ATR) (item 3, Scheme 2). Oxygen is fed into the reactor at a temperature of 200 °C, keeping the molar ratio of O2/CH4 = 0.59.1,10 To ensure the desired molar ratio of H2/CO in a produced syngas, superheated steam is also added into the stream. The temperature of converted gas at the exit of the reactor is equal to 1050 °C. Necessary oxygen is produced from air in the ASU (item 7, scheme 2). In the calculations it is assumed that the tail gases of FT synthesis enough to preheat the process streams and steam generation. (3) Partial oxidation (POX) is carried out through methane oxidation by oxygen only in the following reactions:10,13 CH4 + 0.5O2 → CO + 2H 2

(+ 36 kJ/mol)

CH4 + O2 → CO + H 2O + O2 CH4 + 1.5O2 → CO + 2H 2O CH4 + 2O2 → CO2 + 2H 2O

( +278 kJ/mol) (+ 519 kJ/mol)

(+ 802 kJ/mol)

nCO + 2nH 2 → CnH 2n + nH 2O CO + H 2O ↔ CO2 + H 2

(+ 41 kJ/mol)

(+ 180 kJ/mol)

(+ 41 kJ/mol)

(16) (17)

Temperature of synthesis is constant and equal to 260 °C. Heat of the exothermic reaction is removed by boiling water, which is circulated through the reactor’s shell. The volumetric vapor content at the outlet of the cooling circuit of the FT reactor is assumed to be 30%. The FT reactor is modeled as a stoichiometric reactor with isothermal temperature distribution. The distribution of synthesis products corresponds to Anders−Schulz−Flory molecular weight distribution:1,17 mn = (1 − [α , β ]) ·[α , β ]n − 1

(10)

(18)

where m is the mole fraction of molecules with a carbon number n and α and β are parameters of the molecular weight distribution of paraffinic and olefinic hydrocarbons, respectively. CO conversion in the FT reactor is taken 80% in all simulations.18 Parameters of the molecular weight distribution of the synthesis products are α = 0.82, β = 0.72. The amount of olefins in the final product is assumed equal to 28%. The resulting mixture is cooled and fed to the separator (item 12, Schemes 1−3), where at a temperature of 40 °C it is divided into three streams: liquid hydrocarbons (LHs), water and tail gas (a mixture of gaseous products of the Fischer− Tropsch synthesis, and unreacted syngas). Liquid hydrocarbons under pressure go to the stabilization unit (item 13, Schemes 1−3), where the pressure is released and some of the hydrocarbons go to a gas phase, while the remaining part of the liquid hydrocarbons is taken away in the form of the end product. In considered flowsheets the tail gas is used as a fuel gas for implementation of conversion reaction and heating process streams. Also tail gas can be used for efficiency increase of the overall GTL process by recirculation to various stages of technology. We considered flowsheets with three different circulation loops as shown in Schemes 1−3 by dashed line. A small circulation loop (SC) implies that a part of the tail gas mixes up with fresh syngas and goes back to the FT reactor. A big circulation loop (BC) implies that a part of tail gas mixes up with a stream of natural gas, goes at first in a prereformer, and

(11) (12) (13)

The stream from the prereformer is heated up to a temperature of 650 °C and goes to the partial oxidation reactor (POX) (item 3, Scheme 3). Oxygen is fed into the reactor at a temperature of 200 °C keeping the molar ratio of O2/CH4 = 0.65.1 Converted gas leaves the reactor with a molar ratio of H2/CO = 1.8−1.9 and temperature about 1300−1400 °C. Necessary oxygen is produced from air in the ASU (item 7, Scheme 3). In the calculations it is assumed that the tail gases of FT synthesis enough to preheat the process streams and steam generation. To obtain the desired molar ratio of H2/CO, the resulting converted gas is cooled down to the temperature of 370 °C and fed to the reactor of the high-temperature water gas shift reaction (item 5, Scheme 3), where CO reacts with the superheated steam in the following reaction:12,14 CO + H 2O ↔ CO2 + H 2

(15)

(14)

The temperature of the stream at the outlet is 420 °C. Syngas produced by all the methods is cooled down to 40 °C, so the excess water can be separated (item 6, Schemes 1−3). The dry syngas is directed to the carbon dioxide extraction unit (item 9, Schemes 1−3), where CO2 is extracted C

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then to the main reformer. There is also a stabilized blowdown circulation loop (SBC) in the flowsheets, where the tail gases leaving the stabilization unit are fed into the prereformer. However the amount of these gases is insignificant. Our simulation considers the following parameters common for all flowsheets: (i) input parameters of natural gas, water and air are 1 bar and 20 °C; (ii) pressure drop in devices 1, 2, 3, 8, 10 is 1 bar; (3) hydrogen is released from the block 9 under the pressure of 5 bar;19 (iii) air for oxygen extraction is compressed up to 10 bar, oxygen is produced at 1 bar;20 (iv) all the machines used for pumping and compression are characterized by the same value of efficiency;21 (v) heater air excess factor is equal to 10%. The efficiency of GTL was estimated as a carbon efficiency (%) expressed as the ratio of the number of carbon atoms in the final product to the amount of carbon atoms in the feed gas; and also as a thermal efficiency (%) expressed as the ratio of low heating value of final product to the low heating value of the feed gas. It is important to note that both carbon efficiency and thermal efficiency are closely related with a practically important process productivity, which is usually estimated as a yield of the end product in kilogram per one thousand 1000 m3 of natural gas consumed for all purposes. Also, we estimated the energy consumption on compression and pumping, expressed in MJ/m3 of natural gas, taking into account the following machines: − compressors of natural gas, syngas, carbon dioxide and hydrogen − circulating compressors of three circulation loops − compressor of air and oxygen of the ASU − flue gas blower − air blowers of burners − pumps for water supply on technology − circulating pumps of MDEA solution − circulating pumps of the first and second cooling loops of the FT reactor, as well as feed pumps

Figure 1. Influence of the SDR process pressure on carbon efficiency and energy consumption of GTL technology. Solid lines, results for ordinary reformer; dashed lines, “ideal” reformer (see text).

no natural gas goes to burning. This kind of a flowsheet means high-quality integration of the conversion block by the organization of thermal and material streams both in the main reformer and its furnace. In contrast to SDR, pressure dependencies for autothermal reforming ATR and partial oxidation POX are not analyzed in this work due to more technical difficulties in pressure variation for these two methods. So the present work considers ATR and POX at the pressure of 24 bar,1,10 which is similar to the values used in industrial technology GTL, and which is similar to the pressure of conventional SDR. Figure 2 shows the calculations results of the role of circulation loops in the GTL flowsheets with three different methods of syngas production.

3. RESULTS OF SIMULATION It is well-known that combined steam and dry reforming SDR is more efficient at lower pressures.10,12 However, the decrease of pressure at the SDR stage leads to higher expenses on syngas compression up to the level necessary for the FT process. So optimization of SDR-including GTL requires analysis of the influence of the pressure at the SDR stage. Figure 1 shows how the pressure variation in the main reformer influences GTL efficiency. The increase of pressure from 4 to 24 bar leads to insignificant (1.7%) carbon efficiency decrease. Heat of tail gases combustion is not large enough, so it is necessary to burn an additional amount of natural gas in order to maintain the SDR process. The higher the pressure at SDR is, the larger amount of extra natural gas is required, so it leads to the observed decrease in carbon efficiency. However the energy consumption drops 2-fold down from 2.5 to 1.3 MJ/m3 of natural gas with that same pressure decrease. Also Figure 1 shows the results of simulation of a flowsheet with a very efficient main reformer, where no extra natural gas is necessary to burn to maintain SDR process. The dashed lines in Figure 1 indicate the results for such ideal reformers. In that case, the carbon efficiency of the process increases significantly (by 7.7%) with the SDR pressure increase up to 24 bar. Energy consumption increases as a natural consequence of the fact that

Figure 2. Effect of the circulation loops of different tail gases on the efficiency of the GTL process using SDR, ATR, and POX.

It can be seen from Figure 2 that the best no-circulation process efficiency is achieved in the case of SDR. Carbon efficiency of this process is 52.2% (productivity, 427 kg of LHs per 1000 m3), while thermal efficiency is 43.5%. In the cases of ATR and POX the carbon efficiency is 46.9% (384 kg/1000 m3) and 46.1% (377 kg/1000 m3), respectively, while the thermal efficiency of the process is 39.1% in the case of ATP and 38.4% in the case of POX. D

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It can also be seen from Figure 3 that the introduction of circulation loops leads to further reduction of energy consumption for SDR-based flowsheet. The best result for the most carbon-efficient SC loop is 1.19 MJ/m3. In the cases of ATR and POX the use of SC loop does not affect the energy consumption. However, the use of SBC and BC loops leads to the increase of energy consumption in both ATR and POX cases. In the case of the most carbon-efficient combined loop SC−SBC, the energy consumption is 2.24 MJ/ m3 for the ATR case and 2.41 MJ/m3 for the POX case.

The introduction of circulation loops on the synthesis stage (SC) leads to the efficiency increase for all three flowsheet types. Moreover, the optimal circulation rate (optimal value here is the value where no further significant increase in efficiency is observed with circulation rate increase) is approximately the same for all three conversion methods and is equal to 1.3. The value of carbon efficiency at the optimal circulation rate is ca. 56% (460 kg/1000 m3) for all three flowsheet types, while the thermal efficiency is ca. 47%. The introduction of SBC and BC circulation loops leads to a different effect on the GTL process efficiency. Figure 2 shows that the use of these loops in the case of SDR does not affect the process efficiency. However, the introduction of SBC or BC loops into ATR- or POX-based flowsheets may lead to significant increase in efficiency. This dependence is linear for the BC loop. In the case of ATR the carbon and thermal efficiencies increase up to 59.4% (486 kg/1000 m3) and 49.5%, respectively. In the case of POX the carbon and thermal efficiencies increase up to 61.15% (500 kg/1000 m3) and 50.9%, respectively. The use of the SBC loop only has no significant effect on the efficiency; however the SBC−SC combination provides the greatest growth of efficiency. For the ATR case the carbon and thermal efficiencies of the process reach 63.05% (516 kg/1000 m3) and 52.6%, respectively, while for the POX case the values of 62.05% (507 kg/1000 m3) and 51.7% are reached. It should be noted that the use of the combined SC−BC circulation loop (not shown in the figures) leads to some insignificant further increase in the efficiency of GTL. However, the use of such a loop leads to a very significant increase in the energy consumption for compression and pumping, which affects negatively the economic performance of the process. Figure 3 shows the results for energy consumption of the GTL process at various methods of natural gas conversion and different tail gas circulation loops.

4. DISCUSSION The results of simulation show that the best carbon efficiency for no-recirculation flowsheets is achieved in the case of SDR and reaches 52.15%, which is 5.25% higher than the efficiency of ATR-based flowsheet and 6.05% higher than that of the POX-based flowsheet. It is important to notice that the SDRbased process consumes almost two times less energy than ATR- or POX-based flowsheets do. Moreover, the introduction of circulation loops leads to even lower energy consumption values for the SDR case. The ATR- and POX-based processes, in contrast, increase their energy consumption with introduction of circulation loops. The carbon efficiency of these processes, however, can be significantly increased by introduction of circulation loops. As a result, the highest carbon efficiency of ca. 63% was demonstrated by the ATR-based flowsheet with the combined SDC−SC loop. Application of the most obvious SC loop in the GTL process, wherein the tail gases are fed back into the FT reactor, allows increasing efficiency of technology for all methods of of syngas production. The analysis of efficiency dependencies on recirculation rate shows that there is an optimum value of the SC recirculation rate. The optimum value of the SC recirculation rate is around 1.3 for all three methods of syngas production and provides the flowsheet carbon efficiency of ca. 56%, which is higher than the no-circulation value by 3.85% for the SDR case, by 12.11% for the ATR case, and by 11.83% for the POX case. A further increase of the recirculation rates beyond the optimum value does not lead to significant increase in process efficiency. So the efficiency jump provided by the introduction of the SC loop is more significant for ATR and POX cases. The reason is related with the main effect of the SC loop, that is, with the decrease of circulation purge directed to burners. In the case of the SDR-based flowsheet it means that the tubular furnace requires more fresh natural gas, which leads to the suppression of efficiency growth. Although the best carbon efficiency is achieved in the case of the ATR with combined SC-SBC loops, the technology choice is burdened by a much higher energy expense of ATR- and POX-based processes in comparison with the SDR-based process. So high energy consumption is determined by the use of ASU, which consumes almost half of the necessary power. The POX-including GTL requires more energy than the ATRincluding one just because of higher oxygen consumption and hence more powerful ASU. Also, in terms of technology choice, the energy consumption matters as it influences the economic performance of the GTL technology through capital and operational expenses related to more powerful pumps, turbines, and compressors. The relation between energy consumption and thermal efficiency is of interest. Comparison of the natural gas LHV

Figure 3. Effect of different tail gases circulation loops on the energy consumption of GTL process using SDR, ATR or POX. In this figure designations are the same as in Figure 2

As shown in Figure 3, in the case of SDR the energy consumption without recirculation is 1.31 MJ/m3 of natural gas, which is 1.6 times lower than that for the ATR case (2.13 MJ/m3) and 1.75 times lower than that for the POX case (2.29 MJ/m3). E

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value, which is ca. 43.46 MJ/m3, with full energy consumption shows that the flowsheets with more efficient circulation loops have an energy consumption of 1.19 MJ/m3 for the SDR case and 2.41 MJ/m3 for the ATR case. It means that the introduction of circulation loops adds ca. 0.1 MJ/m3, which corresponds to the loss of about 0.25% of the gas stock LHV. However, the productivity jumps may be as high as 40 g/m3, that is, thermal efficiency gain is up to 16% for ATR-based flowsheets and up to 4% for SDR-based flowsheets.

(4) Gradassi, M. J.; Green, N. W. Economics of Natural Gas Conversion Processes. Fuel Process. Technol. 1995, 42, 65−83. (5) Mordkovich, V. Z. Higher One-Pass Conversion and Productivity in a Scaled-up Pilot GTL Unit. The 10th Natural Gas Conversion Symposium, Doha, Qatar, March 2−7, 2013. Abstracts. No. 687, pp 1−2. (6) Ermolaev, I. S.; Ermolaev, V. S.; Mordkovich, V. Z. Substantiating the Selection of Recirculation Circuits in Technology for Synthesizing Liquid Hydrocarbons from Natural Gas. Theor. Found. Chem. Eng., 2013, Vol. 47, No. 2, 153−158. (7) de Klerk, A. Fischer−Tropsch Refining; Wiley: Weinheim, Germany, 2011. (8) Panahi, M.; Rafiee, A.; Skogestad, S.; Hillestad, M. A natural gas to liquid process model for optimal operation. Ind. Eng. Chem. Res. 2012, 51, 425−433. (9) Iandoli, C. L.; Kjelstrup, S. Exergy analysys of a GTL process based on low-temperature slurry F-T reactor technology with cobalt catalyst. Energy Fuels 2007, 21, 2317−2324. (10) Rostrup-Nielsen, J.; Christiansen, L. J. Concepts in Syngas Manufacture; Catalytic Science Series, Vol. 10; Imperial College Press: London, England, 2011. (11) Aasberg-Petersen, K.; Dybkjær, I.; Ovesen, C. V.; Schjødt, N. C.; Sehested, J.; Thomsen, S. G. Natural gas to synthesis gasCatalysts and catalytic processes. J. Nat. Gas Sci. Eng. 2011, 3, 423−459. (12) Semenov, V. P. Proizvodstvo Ammiaka; Himiya: Moskva, Rossiya, 1985. (13) Arutyunov, V. S.; Krylov, A. V. Oxidative Conversion of Methane; Nauka: Moscow, Russia, 1998. (14) Zare Aliabadi, H. Thermodynamic modeling of the high temperature shift converter reactor using minimization of Gibbs free energy. World Acad. Sci., Eng. Technol. 2009, 25, 189−193. (15) Kohl, A. L.; Nielsen, R. B. Gas Purification, 5th ed.; Gulf Publishing Company: Houston, TX, 1997. (16) Khodakov, A. Y.; Chu, W.; Fongarland, P. Advances in the development of novel cobalt Fischer−Tropsch catalysts for synthesis of long-chain hydrocarbons and clean fuels. Chem. Rev. 2007, 107, 1692−1744. (17) Sheng, M.; Yang, H.; Cahela, D. R.; Yantz, W. R.; Gonzales, C. F.; Tatarchuk, B. J. High conductivity catalyst structures for applications in exothermic reaction. Appl. Catal. 2012, A 445−446, 143−152. (18) Mordkovich, V. Z.; Sineva, L. V.; Solomonik, I. G.; Ermolaev, V. S.; Mitberg E. B. WO Pat. 147513, 2010. (19) De Falco, M.; Marrelli, L.; Laquaniello, G. Membrane Reactors for Hydrogen Production Processes; Springer: London, England, 2011. (20) Kerry, F. G. Industrial Gas Handbook: Gas Separation and Purification; CRC Press: Boca Raton, 2007. (21) Leytes, I. L.; Sosna, M. H.; Semenov, V. P. Theory and Practice of Chemical Energy Technology; Himiya: Moscow, Russia, 1988.

5. CONCLUSION A simulation of GTL flowsheets with different syngasproducing methods has been done. Pressure variation and various circulation loops have been accounted with the purpose of identifying primary factors influencing the process efficiency. The results of the simulation suggest that the GTL with combined steam and dry reforming (SDR) should be carried out at higher pressures, close to the pressure of Fischer− Tropsch synthesis. In the case of SDR, the efficiency of the GTL process increases with the application of the circulation loop of the tail gas into Fischer−Tropsch synthesis, while in the case of ATR or POX better efficiency is achieved through the use of a combined loop on the conversion stage. Other types of recirculation do not result in further optimization of the process. The highest carbon efficiency can be reached with ATR- or POX-based GTL. However, it is reasonable to suggest that a better way of producing syngas in GTL technology is combined steam and dry reforming (SDR) due to much lower power consumption. The practical choice of syngas production method can be different due to factors which lie outside the process energy efficiency, such as cost of reformers, or operating costs, or access to electricity and water, etc. Ways for further improvement of GTL process efficiency, as can be seen from the simulation results, can be found in the optimization of the heat exchange in the main reformer of SDR or in the reduction of energy consumption of air separation units in the cases of ATR and POX.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel.: +7 (499) 272 23 15 int. 374. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This work was partly supported by the Ministry of Education and Science of Russian Federation, Contract No. 16.523.11.3002.

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REFERENCES

(1) Steynberg, A. P.; Dry, M. E. Fischer−Tropsch Technology; Elsevier: Amsterdam, The Netherlands, 2004; Vol. 152. (2) Lange, J.-P. Economics of Alkane Conversion. Sustainable Strategies for the Upgrading of Natural Gas: Fundamentals, Challenges, and Opportunities. NATO Sci. Ser. II: Math., Phys. Chem. 2005, 191, 51−83. (3) Maitlis, P. M.; de Klerk, A. Greener Fischer−Tropsch Processes for Fuels and Feedstocks; Wiley: Weinheim, Germany, 2013. F

dx.doi.org/10.1021/ie402284q | Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX