Novel Syngas Production Techniques for GTL-FT Synthesis of

Nov 10, 2010 - Post Box 217, 7500 AE Enschede, The Netherlands. Syngas production plays a key role in gas-to-liquid (GTL) technology. In this feasibil...
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Ind. Eng. Chem. Res. 2010, 49, 12529–12537

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Novel Syngas Production Techniques for GTL-FT Synthesis of Gasoline Using Reverse Flow Catalytic Membrane Reactors C. Dillerop, H. van den Berg,* and A. G. J. van der Ham Faculty of Science and Technology, Process Plant Design, UniVersity of Twente, Post Box 217, 7500 AE Enschede, The Netherlands

Syngas production plays a key role in gas-to-liquid (GTL) technology. In this feasibility study, existing and new syngas production techniques are compared. As the conventional autothermal reforming reactors (ATR) always consist of extensive heat exchange equipment and air separation units, alternatives are based on integrating these within the reactor design. Reverse flow catalytic membrane reactors (RFCMR) make it possible to perform heat integration within the reactor; integrated air separation is accomplished by incorporating O2 selective membranes. The alternatives show a higher economic potential than the conventional technique. The use of an O2 selective membrane still experiences reliability problems and was therefore not considered. The second part of the study consists of an extended conceptual design of both the conventional design as well as the reactor with integrated heat transfer. The design involved a GTL plant producing 100 000 bbl/day of gasoline. The feasibility of the reversed flow catalytic membrane reactor is shown by reduced methane and oxygen consumption. 1. Introduction Gas-to-liquids (GTL) technology is an increasingly lucrative opportunity to exploit stranded and/or abundant natural gas for conversion to liquid fuels such as gasoline. A GTL plant consists of a syngas unit that produces a gas mixture consisting largely of CO and H2, which is commonly known as synthesis gas or syngas, a Fischer-Tropsch (FT) unit, and product upgrading facilities (see also Figure 1). The production of syngas can only be performed with oxygen. The oxygen can either be supplied as air or pure O2. As air contains (21% O2 (vol %), the use of air for syngas production would require large process equipment, which could be infeasible. For this reason, an air separation unit often provides the syngas section with pure O2. However, air separation is an expensive technology due to the cryogenic conditions required to separate N2 and O2. Studies have confirmed that air separation units cost as much as 23% of the total installed CAPital EXpenditure (CAPEX) of GTL plants (see ref 1). With a proven estimate of a GTL CAPEX cost of U.S. $25 000/daily bbl (see ref 2) of liquid product and given that current projects are planning plant capacities on the order of 100 000 bbl/day of liquid product (typical refinery size), a rough estimate indicates that a CAPEX investment of (U.S. $575 million would be required for the air separation alone. This is a considerable investment and a major driver for the introduction of other possible technologies, which could reduce such investment costs. The most important conventional techniques available for syngas production are (non)catalytic partial oxidation, steam reforming, and autothermal reforming. Steam reforming is mainly used for producing products with a high hydrogen content such as ammonia. Autothermal reforming (ATR) combines partial oxidation with steam reforming. Both autothermal reforming and partial oxidation are being used for syngas production with gasoline as an end product. These processes require oxygen in the feed and extensive heat exchange equipment, as the reaction takes place at an elevated * To whom correspondence should be addressed. E-mail: [email protected].

temperature of 1000 °C or more (refs 2 and 3). Also, the syngas reactor product eventually needs to be cooled down to lower temperatures (e.g., 340 °C, ref 4) for the FT process, again requiring heat transfer equipment. Novel syngas production techniques, as discussed in this paper, are based on three potential improvements: 1. The re-evaluation of the necessity of extensive heat exchange equipment by introducing recuperative heat integration, which could be done by the introduction of a reverse flow reactor (Energy Integration Optimization). 2. The re-evaluation of the necessity of air separation by using a selective membrane to separate the oxygen from the air within the reactor (Air Separation Optimization). 3. The re-evaluation of the necessity of air separation and extensive heat exchange equipment by combining the two potential improvements as described above (Combined Energy Integration and Air Separation Optimization). The purpose of this study was to investigate the technical and economical feasibility of the combinations of energy integration and air separation reactor concepts and a subsequent comparison with the more conventional ATR technology for syngas production for the GTL synthesis of gasoline. 2. Design Process Basically, the design process consists of three parts: (1) literature survey, (2) first technical and economical evalua-

Figure 1. Overview of typical GTL-FT plant.

10.1021/ie1007568  2010 American Chemical Society Published on Web 11/10/2010

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1 partial oxidation (POX, CPO): CH4(g) + O2(g) T CO(g) + 2 2H2(g) (2) carbon dioxide reforming: CH4(g) + CO2(g) T 2CO(g) + 2H2(g) (3) water gas shift reaction (WGS): CO(g) + H2O(g) T CO2(g) + H2(g) (4) total oxidation: CH4(g) + 2O2(g) T CO2(g) + 2H2O(g)

(5)

Figure 2. Working methodology for feasibility study.

tion, and (3) conceptual design. A systematic method is applied, going from black box to conceptual design to index flowsheet and, finally, to the PFD level. It includes all tools such as working diagrams, alternatives, and decision trees and appeared to be robust and flexible at the same time (see also Figure 2). 3. Syngas Production Techniques A detailed literature study has been conducted presenting various technical aspects of GTL (syngas production in particular) as well as a market study and an economic analysis. The most suitable syngas production technique depends on the FT process it will feed. The selected boundary conditions will be discussed later. Also, the type of syngas feed can also have an impact. In this report, only methane (lean natural gas) is considered. 3.1. Syngas Production Chemistry. First of all, the major reactions involved in syngas production are shown: steam reforming (SMR): CH4(g) + H2O(g) T CO(g) + 3H2(g) (1)

In essence, the same reactions take place in all of the syngas production processes. However, a few reactions are usually more favorable in comparison to others. This varies for the different technology types. 3.2. Conventional Syngas Production Techniques. Steam methane reforming (SMR) refers to a reactor in which methane and steam are catalytically and endothermically converted to hydrogen and carbon monoxide. The dominant reactions are steam reforming and the water gas shift reaction. The reactor is usually a tubular reformer (operating at about 1000 °C and 25-40 barg) where the heat is supplied externally from the furnace. Partial oxidation (POX) is an exothermic, noncatalytic reaction of methane and oxygen to produce a syngas mixture. Usually, steam is injected as well to increase the H2 content in the syngas (ref 2). POX is a high-temperature process with reactor operating temperatures of around 1400 °C and pressures up to 80 barg. Catalytic partial oxidation (CPO, or also CPOX) is different from noncatalytic partial oxidation (POX) in that chemical conversion takes place over a catalyst bed, but it does not use a burner. A disadvantage of CPO is the presence of a highly flammable mixture upstream of the reactor, and low feed temperatures could be required (autoignition temperature is approximately 250 °C). Therefore, the feed temperature should be lowered to 200 °C, leading to higher oxygen and methane consumption (ref 3). Typical operating conditions are similar to the ATR described below (see also ref 3). As described before, (non)catalytic partial oxidation, steam reforming, and autothermal reforming are the most common syngas production techniques. Autothermal reforming is of most interest for this report, as its syngas is suitable for gasoline production. Autothermal reforming (ATR) is an energy integrated syngas production process. The process includes partial combustion of the feed in order to generate heat for thermal and catalytic steam reforming. A refractory-lined pressure vessel is used for the ATR process. The hydrocarbon and oxygen are fed to the entrance of the reactor as two separate feed streams. Here, the two streams are mixed by a specially designed burner, hence enabling combustion of the hydrocarbon, and as such, is termed the combustion zone, which could reach temperatures up to 2000 °C. The lower section of the reactor is loaded with a hightemperature reforming catalyst operating at about 1000 °C. The operating pressure ranges from 20 to 100 barg. The energy required for the endothermic reforming is provided by the exothermic combustion of the hydrocarbon at the reactor inlet. A typical process diagram for the application of an ATR is given in Figure 3a. 3.3. Novel Syngas Production Techniques. As described before, novel syngas production techniques are based on three potential improvements in energy integration, air separation or

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Figure 3. Schematic overview of syngas reactors: (a) ATR reactor, (b) RFCMR reactor with porous membranes, (c) CMR reactor with O2 selective membranes, (d) RFCMR reactor with O2 selective membranes. Table 1. Comparison of Syngas Production Reactors

process

ASU required (O2 prod.)

recuperative heat exchange integrated within reactor

O2/CH4 (volume ratio)

ATR RFCMR (porous) CMR (O2 sel.) RFCMR (O2 sel.)

yes yes no no

no yes no yes

0.6 0.5 0.6 0.5

combined energy integration, and air separation. Figure 3b, c, and d and Table 1 show the schematics of these novel techniques and are described below in more detail. In all cases, methane is the reactor feed and syngas is the reactor product. Depending on the syngas production technique, either pure oxygen or air (assumed to be oxygen and nitrogen) is a reactor feed, and nitrogen is a separate reactor effluent. Optimal energy integration is achieved with the so-called reverse flow catalytic membrane reactor with porous membranes (RFCMR porous, see Figure 3b). There are two separate reactor feeds for methane and oxygen. The RFCMR with porous membranes is able to alternate feed flow directions during specific cycles in order to keep a constant temperature profile within the reactor. The switching of feed directions allows for the storage of heat from the slightly exothermic CPO reaction in special heat capacity storage regions and subsequent recovery thereof in the next cycle, where the feed direction is switched (see also Figure 4). This storage and recovery of heat presents another possible advantage of RFCMR. Less O2 is required to generate heat by combusting part of the CH4 feed, and less heat exchange is required compared to ATR. This is also shown in Table 1 by the O2/CH4 ratio, which is lower for the reverse flow reactors. For more detailed information on RFCMR, refer to ref 5. Air separation optimization is obtained in the catalytic membrane reactor (CMR) with selective membranes. The separation of the oxygen from the nitrogen is achieved by perovskite membranes, which selectively permeate oxygen. The operating conditions for the perovskite membrane are approximately 20 bar and 1100 °C, which is consistent with those required for the catalytic partial oxidation (CPO) of natural gas to produce syngas. The obvious benefit of this reactor type is the negation of the air separation unit (ASU; ref 6).

Figure 4. Schematic overview syngas RFCMR reactor with porous membranes, including the desired temperature profile within the reactor.

Combining both energy integration and air separation is possible with the reverse flow catalytic membrane reactor (RFCMR) with O2 selective membranes. Both the reverse flow reactor concept as well as the O2 selective perovskite membranes are integrated within one reactor design. Both conventional and novel syngas production techniques have been outlined and can be subject to further analysis by a feasibility study. 4. Feasibility Study A feasibility study is performed on the above-described syngas production techniques by performing a technical and economical evaluation. The feasibility study is performed according to the schedule shown in Figure 2. 4.1. Functional Blocks and Black Boxes. The first part of the evaluation is to determine the functional blocks of the GTL plant: air separation unit (ASU), syngas production unit, and

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Figure 5. Black box for process based on oxygen (A) or air (B) feed.

Fischer-Tropsch unit (FT). Two black boxes are made on the basis of these functional blocks: a GTL plant with or without an ASU. As can be seen in Figure 5, the end products are the same. However, for this study, the boundaries have been placed around the syngas production and FT unit only. This requires boundary conditions for the oxygen feed (Figure 5a with ASU) or the air feed (Figure 5b without ASU). The first comparison of the different techniques is based on simplified models and related calculations. For example, the reactions above are considered to be stoichiometric. Thus, reaction equilibria are not considered at this stage, and some components are left out of the equation (e.g., CO2). In a later stage, more detailed calculations shall be performed with a conceptual design. 4.2. Boundary Conditions. The design capacity and boundary conditions of the GTL-FT processes with different syngas production techniques are as follows: • The production capacity is 100 000 bbl/day of gasoline (assumed to be C8H18). • The process feedstock is either pure oxygen or air. • The methane is pure: This is intended to simulate a natural gas feedstock. The use of just methane to represent natural gas is not expected to have a huge impact on the process analysis, as natural gas contains up to 95% methane. • Table 2 provides the prices used for the economic evaluation. • Feed ratios for the novel reactor types are already presented in Table 1. Only partial oxidation of CH4 is assumed with the syngas production unit. • In the case that heat integration is applied (e.g., with RFCMR), the applicable stoichiometric reaction used as a function of the O2/CH4 ratio is 0.5: 1 CH4 + O2 f CO + 2H2 2 • In the case that no heat integration is applied, the applicable stoichiometric reactions used as a function of the O2/CH4 ratio is 0.6, as additional oxygen is required for the combustion of methane: 10CH4 + 6O2 f 10CO + 18H2 + 2H2O • Complete conversion of the limiting reagent (H2) is assumed across the FT unit. • Gasoline density was taken to be approximately 724 kg m-3. • Byproducts and recycles (such as CO2; see also Figure 5) are not considered. The FT unit is described with the following stochiometric reaction: 8CO + 17H2 f C8H18 + 8H2O

4.3. Mass Balances and Economics. The results for the feasibility study are provided in Table 3. The total flow describes the product flows at the exit of the FT process, including byproducts such as CO. The carbon and hydrogen efficiency describe the percentages of carbon and hydrogen atoms that end up in the end product. These are basically determined by the stoichiometric reactions as described above. The efficiencies already give an indication that the novel reactor types (RFCMR with porous membranes, CMR with O2 selective membranes, and RFCMR with O2 selective membranes) are economically more feasible compared to the ATR by comparing the economic potentials (EP0): • RFCMR (porous): reduced CH4 and O2 consumption. • CMR (O2 selective): no oxygen cost. • RFCMR (O2 selective): reduced CH4 consumption and no oxygen costs. EP0 is defined as the economical potential (products - raw materials) for transition to either of the novel reactor concepts related to the maximum attainable profit for the ATR. In the case that the reduced O2 consumption for RFCMR (porous) would not be included, the sum of EP0 for RFCMR (porous) and CMR (O2 selective) would be the same as for the EP0 of RFCMR (O2 selective). The breakeven point between win and loss for this investment is also of importance. The calculated range of the maximum allowed CAPEX, for varying interest rates (8-12% p/a) and payback periods (4-10 years), was estimated to be U.S. $22 000-55 000/daily barrel of liquid product. This was found to be consistent with values reported in the literature (ref 2) and served as a validation of the results of the feasibility study. 4.4. Discussion. Upon liaison with process engineers active in the field of RFCMR/CMR technology, it became clear that the perovskite membranes exhibited reliability problems at the elevated pressures required. This, coupled with the fact that the RFCMR should allow for switching feed directions at specific intervals, raised further reliability concerns and shifted the focus of the project to RFCMR, which uses porous tubes and already showed improved reliability. Hence, the options with O2 selective membranes were not considered in the conceptual design, as RFCMR with porous membranes was considered to be the most practical candidate to compete with ATR. The decision process is also shown in Figure 6 with a decision tree. 5. Conceptual Design ATR and RFCMR with Porous Membranes The first part of the conceptual design is similar to the feasibility study described above (see also Figure 2), however, with more accurate calculations. For example, more and better defined boundary conditions are applied, and a Gibbs reaction model is used instead of stoichiometric reactions. After defining boundary conditions and operating conditions, detailed mass and energy balances are made for both ATR and RFCMR with porous membranes. The end products are index flow sheets. 5.1. Boundary Conditions. The boundaries of the GTL-FT plant are already described in Figure 5A. The feeds CH4, O2, and H2O are set to 20 bar and 27 °C. The intermediate syngas is set to 20 bar at 340 °C, and FT products exit the unit with 1 bar at 25 °C. Instead of using one simple stoichiometric reaction for describing the FT product, a more sophisticated model was used that was based on high temperature Fischer-Tropsch (HTFT) (ref 1). HTFT gives a wide product distribution ranging from the light gases up to the heavy waxes. However, for HTFT, the lighter fractions, especially the gasoline, are the largest in size.

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Table 2. Prices of Components for Feasibility Study component

price

basis/source

CH4 air O2

U.S. $0.08/kg U.S. $0.000/kg U.S. $0.04/kg

steam H2O gasoline (C8H18) H2 CO CO2

U.S. U.S. U.S. U.S. U.S. U.S.

location: Qatar (ref 7) air is assumed to be free on site source: Ullmann Encyclopedia price: U.S. $5-100/100 m3, depending on location of O2 plant; assumed O2 plant is on site w U.S. $5 per 100 m3; density of O2: 1.326 kg/m3 http://www.oit.doe.gov/bestpractices/steam rough assumption ref 7 rough assumption not necessarysonly in product stream not necessarysonly in product stream

$0.02/kg $0.002/kg $0.63/kg $3.00/kg $0.00/kg $0.00/kg

Table 3. Results of Feasibility Study of Different Syngas Production Techniques process

total flow (t/d)

carbon efficiency

hydrogen efficiency

max profit ($ MM/day)

economic potential

ATR RFCMR (porous) CMR (O2 sel.) RFCMR (O2 sel.)

33567 27464 33567 27464

85% 94% 85% 94%

48% 53% 48% 53%

5.39 5.69 6.12 6.21

basis 6% 14% 15%

Again, the aim is to produce 100.000 bbl/day of gasoline equivalent to 280.000 bbl/day of FT product. This resulted in a HTFT molar feed ratio H2/CO of 2.14. Table 4 provides an overview of the mass balance around the FT unit. However, to include the effect of the water-gas-shift ratio, the module M shall be included as well. As the module M is 2 for HTFT (ref 8), the feeds shall be adjusted accordingly. M)

H2 - CO2 CO + CO2

Other important assumptions were as follows: the feed is assumed to be pure O2 and CH4 (to represent natural gas) and syngas is directly routed to the HTFT unit, without any composition alteration. 5.2. Operating Conditions ATR. The following operating conditions were applied for the ATR conceptual design:

Figure 6. Decision tree for feasibility study.

• The molar O2/CH4 ratio is assumed to be 0.6 (see also Table 1). • The effect of steam injection into an ATR on H2 production is limited, as it would also produce extra CO2. This would not improve the module M, as the water-gas-shift reaction would convert the H2 (see also Figure 10). Steam reforming is assumed to be a possible technology that can be utilized to increase the module (by improving the H2 content), should the need arise. The operating conditions for the steam reforming unit were fixed at 900 °C and 25 bar (ref 9). The elevated pressure level of 25 bar in comparison to the ATR, and the feed for the FT has not been further evaluated in the design. • The operating pressure of the ATR reactor was fixed at 20 bar, coinciding with typical operating pressures for RFCMR

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Table 4. Results of Mass Balance on HTFT Unita feed required for group group/average composition lights/CH4 C2-C4 olefins/C3H6 C2-C4 paraffins/C3H8 gasoline/C8H18 middle distillates/C5H10 heavy cuts and waxes/C30H62 oxygenates (e.g., MTBE)/C5H12O Subtotal [ton/day] [kmol/s] H2/CO ratio total feed/product [ton/day] a

product

CO [tons/day]

H2[tons/day]

organic [tons/day]

H2O [tons/day]

5026 15078 3770 22617 10052 3141 3141 62826 26

1077 2154 628 3433 1436 456 449 9633 55.7

2872 7539 1975 11511b 5026 1578 1975 32475

3231 9693 2423 14540 6462 2019 1616 39984

2.14 72460

72460 b

Product upgrading of other fractions to produce gasoline blend not considered. Translates to 100 000 bbl/day gasoline.

Figure 7. Syngas composition versus reactor temperature (isothermal Gibbs reactor (ATR) at 20 bar with O2/CH4 ) 0.6 and H2O/CH4 ) 0.0).

Figure 8. Graph of CH4 conversion and “CO + H2 selectivity” versus reactor temperature (isothermal Gibbs reactor at 20 bar with O2/CH4 ) 0.6 and H2O/CH4 ) 0.0).

(see ref 5), thereby allowing comparison. This also falls within the range of commercially available ATRs. • A sensitivity analysis was done to investigate suitable operating temperatures for the ATR. The results are shown in Figures 7, 8, and 9. • Figure 7 shows that in terms of the CO and H2 produced there is little benefit to operating at temperatures in excess of 1100 °C. Similarly, it can be seen from Figure 8 that there is also little benefit in terms of CH4 conversion and syngas (CO + H2) selectivity to operate at temperatures in excess of 1100 °C. Hence, the temperature of the ATR unit was fixed at approximately 1100 °C. • The feed temperature to the ATR was estimated by doing a sensitivity analysis using an adiabatic GIBBS reactor (Figure 9). The feed temperature was varied, and the reactor temperature was monitored. It can be seen that for a reactor feed temperature of approximately 500 °C, energy gener-

Figure 9. ATR temperature vs feed temperature (adiabatic Gibbs reactor at 20 bar with O2/CH4 ) 0.6 and H2O/CH4 ) 0.5).

ated by the combustion of methane is sufficient to maintain a reactor temperature of approximately 1100 °C. 5.3. Operating Conditions RFCMR with Porous Membranes. The following operating conditions were applied for the ATR conceptual design: • The molar O2/CH4 feed ratio is fixed at 0.44. It was decided to change the previously used O2/CH4 feed ratio of 0.5 (see Table 1) so that a direct comparison of published simulated and experimental data could be executed (see ref 5). • The molar H2O/CH4 feed ratio for RFCMR was fixed at 0.12 (see ref 5). • The RFCMR temperature and pressure were fixed at 1100 °C and 20 bar, respectively (see ref 5). • It is noted once more that these are values from the literature. A check has been made with Aspen showing that these conditions result in a module M of about 2, as was expected (see also Figure 10). 5.4. Material and Energy Balances. Detailed simulations for both concepts were done in Aspen (with Peng-Robinson EOS), and an overview of the results is shown in Table 5. More details can be found in the applicable PFDs (Figure 11 and Table 6; Table 7 and Figure 12). The stoichiometric module (M) of 2 could easily be achieved for RFCMR with porous membranes. This was not the case for the ATR where a maximum M of 1.79 was found, also when including steam injection (see Figure 10). This is attributed to the larger CO2 content in ATR effluent, which is a result if there is combustion of part of the ATR feed to generate heat. The larger CO2 content reduces the module M. For this reason, additional steam methane reformers (SMR) were required to increase the M value by virtue of increased hydrogen production.

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Figure 10. Check operating conditions RFCMR with porous membranes (oxidation of methane for different H2O/CH4 ratios at 20 bar and 1100 °C for Gibbs reactor according to PR).

The required SMR capacity was found to be approximately 26% that of ATR, on the basis of syngas mass flow rates. The ATR/SMR process also consisted of considerably more auxiliary process equipment, i.e., heat exchangers, pumps, and compressors compared to the RFCMR, which is a stand-alone unit. The RFCMR additionally presented improved H2 and CO selectivities (SCO ) 0.97, SH2) 0.95) compared to ATR/SMR (SCO ) 0.86, SH2) 0.79). Once again, this is attributed to the additional CO2 produced by ATR during the combustion of

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natural gas for heat generation. In terms of component efficiencies, the carbon and hydrogen efficiencies are 0.85 and 0.84 for RFCMR and 0.77 and 0.72 for ATR/SMR, respectively. This could imply larger process equipment for the downstream FT unit for an ATR syngas production facility, but this was not considered in this study. 5.5. Number of Reactor Units. In order to get more information on the capital costs for the ATR combined with SMR, the number of units is determined on the basis of the material balances (see also Table 6) and typical unit sizes. The number of required ATR units is calculated to be 13, which is based on an overall syngas production rate of 1106 kg/s and a maximum unit size of 93 kg/s (ref 11) capacity with 95% availability. The number of required SMR units is calculated to be 4, which is based on an overall hydrogen production rate of 31 kg/s and a maximum unit size of 8 kg/s (ref 12) capacity with 95% availability. There were no actual data available of typical unit sizes for a RFCMR unit with porous membranes. However, a feasible reactor capacity was determined on the basis of a shell and tube configuration. Simulations, as available (see ref 5), were used to determine the capacity of a single segment of such a reactor type (one tube and one shell). Furthermore, on the basis of a reactor diameter of 3 m, the maximum number of tubes was calculated. The resulting syngas production capacity was 10 m3/s per reactor. The total number of units is calculated to be 53 based on an overall syngas production of 483 m3/s and 95%

Figure 11. Simulation schematic of ATR/SMR syngas production. Table 5. Results of Conceptual Design variable

ATR

RFCMR porous

M feed CH4 [kmol/s] feed O2 [kmol/s] feed H2O [kmol/s] total molar flow [kmol/s] ηC ) [n(CO)]/[n(CH4)feed] (carbon efficiency) ηH ) [2n(H2)]/[4n(CH4) + 2n(H2O)] (hydrogen efficiency) SCO ) CO/(CO + CO2) (CO selectivity) SH2 ) H2/(H2 + H2O) (hydrogen selectivity) XCH4 (overall conversion) total hot utilities [MW]

2 33.5 15 21 69.5 0.77 0.72 0.86 0.79 0.88 3618

2 30.5 13.4 3.7 47.6 0.85 0.84 0.97 0.95 0.89

total cold utilities [MW]

2748

potential impact specified: boundary condition larger (and possibly more) process equipment needed for ATR-GTL FT; could lead to increase in CAPEX

increased CO2 levels to HTFT unit for ATR-GTL-FT increased H2O levels to HTFT unit for ATR-GTL-FT no real advantage to either technology additional heat exchange equipment and/or increased utility requirements (energy integration not considered)

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Table 6. Heat and Material Balance for Syngas Unit of ATR GTL-FT Process (Numbering As Per Figure 11) stream no.

temperature pressure vapor fraction total mole flow mass flow volume flow enthalpy mole flow CO H2 H2O CO2 O2 CH4

°C bar kmol/s kg/s m3/s MW kmol/s

1

2

3

4

5

6

27 20 1 15 480 18.5 -1.6

27 20 0.668 37.5 627 30.4 -5478

500 20 1 37.5 627 120.7 -4100

27 20 0.501 17 289 10.4 -3090

500 20 1 17 289 54.5 -2351

1096 20 1 87 1106 497.3 -4102

12.5

12.5

8.5

8.5

25

25

8.5

8.5

900 25 1 27 289 106 -851

8

9

10

340 20 1 87 1106 223.4 -6326

340 20 1 27 289 69.4 -1408

340 20 1 114 1396 291 -7733

21.6 47.6 14.5 3.2

4.4 15.7 2.9 0.6

21.6 47.6 14.5 3.2

4.4 15.7 2.9 0.6

26.0 63.3 17.3 3.9

0.2

3.5

0.2

3.5

3.7

15

Table 7. Heat and Material Balance for Syngas Unit of RFCMR GTL-FT Process (Numbering As Per Figure 12) stream no. 1 temperature pressure vapor fraction total mole flow mass flow volume flow enthalpy mole flow CO H2 H2O CO2 O2 CH4

7

2

3

4

°C bar

27 834 1047 340 20 20 20 20 0.925 1 1 1 kmol/s 47.58 47.58 87.66 87.66 kg/s 985 985 985 985 m3/s 54 220 483 225 MW -3336 -1273 -1273 -3336 kmol/s 26.0 26.0 54.2 54.2 3.7 3.7 3.0 3.0 0.8 0.8 13.4 13.4 30.5 30.5 3.8 3.8

availability. Lower availability and reliability for the RFCMR can be expected as with the ATR and SMR process but has not been considered for the moment. More research and pilot plants will be required in order to achieve an acceptable level of availability and reliability of the current conceptual design of the RFCMR. 5.6. Other Process Issues. In terms of safety, it can be stated that the RFCMR with porous membranes could potentially be

a safer technology, as feeds do not have to be preheated (entrance temperature of approximately 27 °C), thereby reducing the possibility of the autoignition of natural gas ((580 °C) under system malfunctioning. However, it is also fair to note that switching of the feed directions at operating pressures of 20 bar could also result in unforeseen malfunctions or failure of the RFCMR or auxiliary process equipment. 5.7. Capital Expenditure. It was already shown in section 5.4 that there is economic potential for both RFCMR and ATR. A short economic evaluation showed that RFCMR is indeed a more economical technology with an estimated CAPEX and OPEX of U.S. $0.88 billion and U.S. $1.67 billion p/a, respectively. The corresponding CAPEX and OPEX for ATR/ SMR syngas is U.S. $2.78 billion and U.S. $2.31 billion p/a. It should be noted however that the CAPEX for RFCMR was based on vectorial estimation based on the RFCMR reactor. Also, the CAPEX for ATR/SMR was based on extrapolated values from the literature (ref 10). 5.8. Conclusions and Recommendations. From the above discussion, it can be concluded that the RFCMR with porous membranes is a promising technology compared to the combination of ATR and SMR for syngas production for GTL synthesis of gasoline. The main advantage is the integration of reaction and heat exchange within one piece of equipment, thereby reducing the required CH4 and O2 feed. This provides

Figure 12. Simulation schematic of RFCMR (with porous membranes) syngas production.

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an economic driver for the RFCMR concept: capital cost reduction is estimated to be about 70%, and operating cost reduction is estimated to be about 30%. However, it should be remembered that RFCMR is still a lab-scale design reactor and needs to be proven on an industrial scale. Also, some boundary conditions specified above could also have an impact on the overall result. For example, here, high temperature FT is applied as a requirement at the back end. However, other process conditions and syngas compositions are required in the case of low temperature FT, methanol production, or hydrogen production. Also, in this report it was assumed that hydrogen is 100% converted in the FT unit. However, if this is lower, more hydrogen would be required, and module M needs to be increased further, favoring the RFCMR concept even further. Feedstock is another example, if naphtha is used instead of natural gas. From this study, it is envisaged that key issues that still require attention in the pilot/industrial-scale RFCMR will be • Cycle time optimization. • Equipment reliability under high pressure feed direction changes. • Maintaining a desired plateau temperature in the reaction zone. Nonetheless, the results of this study do suggest that more should be invested to improve the understanding of an industrial RFCMR, as this is undoubtedly a promising technology. Acknowledgment Joris Smit and Umesh Ramdhani are acknowledged for their contributions to this paper. Literature Cited (1) Wilhelm, D. J.; Simbeck, D. R.; Karp, A. D.; Dickenson, R. L. ‘Syngas Production for gas-to-liquids applications: technologies, issues and outlook’. Fuel Process. Technol. 2001, 71, 139–148.

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(2) Vosloo, A. C. ‘Fischer-Tropsch: A Futuristic View’. Fuel Process. Technol. 2001, 71, Sasol Technology Research and Development, Sasolburg, South Africa. (3) Aasberg-Petersen, K.; Christensen, S. T.; Nielsen, C. S.; Dybkjaer, I. ‘Recent developments in autothermal reforming and pre-reforming for synthesis gas production in GTL applications’. Fuel Process. Technol. 2003, 83, Haldor Topsoe A/S, Nymollevej 55, DK-2800 Lyngby, Denmark. (4) Steynberg, A. P.; Espinoza, R. L.; Jager, B.; Vosloo, A. C. ‘High temperature Fischer-Tropsch in commercial practice’. Appl. Catal., A 1999, 186, 41–54, Sasol Technology Research and Development Division, P.O. Box 1, Sasolburg, South Africa. (5) Smit, J.; Van Sint Annaland, M.; Kuipers, J. A. M. ‘Feasibility Study of a Reverse Flow Catalytic Membrane Reactor with Porous membranes for the Production of Syngas’. Chem. Eng. Sci. 2005, 60, 6971–6982. (6) Smit J., Van Sint Annaland M., Kuipers J. A. M.; ‘Modelling of a Reverse Flow Catalytic Membrane Reactor for the Partial Oxidation of Methane’; International Journal of Chemical Reactor Engineering, Vol. 1, 2003, Article A54. (7) Economides M. J., ‘The Economics of Gas to Liquids Compared to Liquefied Natural Gas’, World Energy, Volume 8, Nr 1, 2005. (8) Aasberg-Peterson, K.; Bak Hansen, J. H.; Christensen, T. S.; Dybkjaer, I.; Seier Christensen, P.; Stib Nielson, C.; Winter Madsen, S. E. L.; Rostrup-Nielsen, J. R. ‘Technologies for large-scale gas conversion’. Appl. Catal., A 2001, 221, 379–387, Haldor Topsoe A/S, Nymoellevej 55, Lyngby 2800, Denmark. (9) Haldor Topsoe, Products and Services - Technologies - Synthesis Gas - Oxogas, www.topsoe.com, Cited: October 2005. (10) Mintz, M., Molburg, J., Folga, S., Gillette, J., ‘Hydrogen Distribution Infrastructure’, Argonne National Laboratory, 2003. (11) Peter, Brook ‘GTL-The Technical Challenges’, Foster Wheeler Oil and Gas Division, www.fwc.com, 2005. (12) Mintz, M., Molburg, J., Folga, S., Gilette, J., ‘Hydrogen Dictribution Infrastructure’, Argonne National Laboratory, Center for Transportation and Research and Decision and Information Sciences Division, Argonne, IL 60439, USA.

ReceiVed for reView March 30, 2010 ReVised manuscript receiVed October 13, 2010 Accepted October 19, 2010 IE1007568