Design Aspects of Membrane Reactors for Dry Reforming of Methane

still prohibitively expensive and the cost of the laboratory membranes is unknown. Comparison of the energy costs of the ... Published on Web 12/14/20...
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Ind. Eng. Chem. Res. 2001, 40, 242-251

Design Aspects of Membrane Reactors for Dry Reforming of Methane for the Production of Hydrogen W. J. Onstot, R. G. Minet, and T. T. Tsotsis* Department of Chemical Engineering, University of Southern California, Los Angeles, California 90089-1211

This paper presents an evaluation of various plant designs incorporating high-temperature membranes for the dry-reforming reaction of methane for hydrogen production. Two different types of membranes are evaluated: a high-temperature ceramic membrane to be used in the catalytic reformer and a carbon-based molecular sieve membrane to be used in the plant’s purification section. Membranes evaluated in this work either are commercially available or have been demonstrated at the laboratory scale. Only operating energy costs for the various designs are presented, because at this stage of their development, commercial membranes are still prohibitively expensive and the cost of the laboratory membranes is unknown. Comparison of the energy costs of the various designs suggests, however, where future development work may be warranted. Among the plant designs considered, those utilizing the carbon-based membranes had the lowest energy requirements. Designs using commercially available hightemperature ceramic membranes exhibiting Knudsen diffusion separation factors were generally found to be inefficient. A unique design utilizing a yet-to-be-developed microporous ceramic membrane is also presented. A plant utilizing this membrane is shown to be more energy efficient and simpler to operate than conventional hydrogen plants. Introduction High-temperature membrane reactors (MRs) have attracted considerable attention in recent years.1-3 Among the potential virtues of these reactors is the fact that in such systems reaction and separation are performed in the same vessel. In addition, conversions greater than equilibrium may be obtained in these MRs, while in conventional plug-flow reactors (PFRs), conversion is limited by thermodynamic equilibrium. For endothermic reactions, this potentially allows the MRs to achieve the same conversion as that attained in a PFR at significantly lower temperatures. One potential use of MRs is steam/CO2 reforming of methane for the production of either synthesis gas or hydrogen. Early investigations4,5 of the application of MRs to this reaction have demonstrated increased conversions over those which could be obtained in a conventional PFR. Most investigations have centered on the use of dense Pd or Pd/Ag membranes. These membranes, while exhibiting excellent selectivity toward hydrogen, suffer from a number of drawbacks: they are expensive, exhibit relatively low H2 permeability, are susceptible to poisoning by sulfur and coke (both typically encountered in catalytic reforming), and are prone to pinhole/crack formation as a result of hydrogen embrittlement and fatigue. To avoid these difficulties, other investigators6,7 have used porous ceramic membranes. These membranes exhibit, in general, better mechanical and thermal stability than dense film membranes but allow gases other than hydrogen to permeate through the membrane. This negatively impacts product yield and process economics. Commercially available ceramic membranes are mesoporous, with pore sizes such that they exhibit Knudsen separation factors (i.e., fluxes inversely proportional to the square root of molecular weight).

A technical and economic evaluation of the use of a dense Pd membrane in methane steam reforming has been presented by Aasberg-Petersen et al.8 They assumed a thin (2-µm-thick) Pd membrane which exhibited perfect separation, and as a result, the pure hydrogen product was taken from the permeate side of the membrane. No sweep gas was used on the permeate side of the reactor. This necessitated compression of the low-pressure hydrogen product. The authors concluded that membrane-based reforming using a dense film membrane became attractive only in the cases where electrical costs were low. This paper is distinct from the paper of AasbergPetersen et al. in that it considers the application of both mesoporous and microporous membranes in the reforming of methane to produce hydrogen. The emphasis, furthermore, is not on the production of pure hydrogen but on a hydrogen product of a certain purity to be used for power generation in either mobile or stationary applications. The driver of our own MR research is small- to medium-scale applications involving the reforming of landfill gas or other types of biogases to produce fuel hydrogen. Landfill gas or biogas is the end product of in situ biodegradation of municipal solid waste placed in the landfills that dot the nation’s landscape. For the purpose of this study, landfill gas was assumed to consist of equimolar amounts of CH4 and CO2, a typical composition for such a gas. The results of this study, however, are also of relevance to dry reforming of natural gas (i.e., steam reforming in the presence of substantial CO2 amounts in the feed), which is attracting increased attention today with the newly found emphasis on CO2 and its potential as a global-warming agent. In fact, several recent studies have focused on the application of MRs to dry reforming.9-11

10.1021/ie0003685 CCC: $20.00 © 2001 American Chemical Society Published on Web 12/14/2000

Ind. Eng. Chem. Res., Vol. 40, No. 1, 2001 243 Table 1. First Law Energy Analysis of Proposed Hydrogen Plant Designsa

feed, GJ/h fuel steam 40 barg 14 barg 3.5 barg electricity energy usage, GJ/h MJ/kgmol HP H2 product a

case 2: reformer w/Al2O3 membrane

case 2a: reformer w/Al2O3 & C membrane

case 3: conventional H2 plant w/fuel cell

case 3a: conventional H2 plant w/fuel cell & C membrane

case 1: conventional H2 plant

case 1a: conventional H2 plant w/C membrane

106.00 82.40

110.80 86.50

225.50 155.80

237.90 164.40

208.00 161.50

217.10 169.50

128.8 78.07

-20.74 -9.47 72.21

-22.12 -10.49 9.50

230.40 464.68

174.19 349.75

-22.34 -15.01 140.97 -114.05 370.87 744.65

-23.57 -15.86 5.01 -119.81 248.07 499.37

-41.31 -18.58 141.61 -114.05 337.17 677.01

-43.36 -20.55 18.61 -114.05 227.25 456.45

206.87 415.44

case 4: autoreformer design

Negative numbers denote energy credits.

Table 2. Estimated Operation Costs for Various Hydrogen Plant Designs

feed, $/h fuel steam 40 barg 14 barg 3.5 barg electricity oxygen operating cost, $/h $/kgmol H2 product

case 2: reformer w/Al2O3 membrane

case 3a: conventional H2 plant w/fuel cell & C membrane

case 1: conventional H2 plant 212.00 164.80

221.60 173.00

451.00 311.60

475.80 328.60

416.00 323.00

434.20 339.00

257.60 156.14

-68.44 -28.41 180.53 0.00 0.00 460.47 0.88

-73.00 -31.47 23.75 0.00 0.00 313.88 0.59

-73.72 -45.03 352.43 -2004.98 1690.60 681.89 1.47

-77.78 -47.58 12.53 -2106.24 1783.60 369.12 0.91

-136.32 -55.74 354.03 -2004.98 1690.60 586.58 1.30

-143.09 -61.65 46.53 -2004.98 1690.60 300.61 0.75

413.74 0.79

The emphasis in this study is on evaluating the various design options in terms of their energy operating costs. This is because we believe that given the current state of affairs in the area of high-temperature membranes (the carbon-based and high-temperature ceramic membranes discussed here are still in an early stage of development), it is rather difficult to estimate the capital costs of a MR-based plant. Our view is that it is unlikely that any such plant will be adopted commercially in the short run, unless its energy operating costs promise to be competitive with (and most likely significantly better than) those of competing conventional plants. Seven different plant designs are presented here: a conventional hydrogen plant for the production of hydrogen; a conventional hydrogen plant equipped with a microporous carbon membrane-based CO2 removal unit; two designs utilizing a high-temperature, mesoporous ceramic MR in place of a conventional plug-flow reformer; and a unique hydrogen plant design, which utilizes a yet-to-be-developed high-temperature microporous membrane. In addition, two cases in which excess highpressure hydrogen is made and utilized in a fuel cell are presented. Energy usage and operating economics are presented for all of the designs. The membrane separation characteristics are shown to be critical, not only in terms of improving the reactor yield but also in terms of impacting product purity and downstream purification process requirements. Other parameters that have a significant impact on the product hydrogen cost include the price of electricity, the oxygen cost that may be required for optimal fuel cell operation, and the value of the landfill gas. Our calculations indicate that variation in the price of one of these quantities is

case 2a: reformer w/Al2O3 & C membrane

case 3: conventional H2 plant w/fuel cell

case 1a: conventional H2 plant w/C membrane

case 4: autoreformer design

sufficient to swing the economics in favor of the conventional or the MR-based plant designs. Design Basis The design of seven different hydrogen plants is presented. The basis of design for all of the plants is the production of 10 million standard ft3 of hydrogen product/day (10 MMSCFD) of 97% purity. The product hydrogen is to be supplied at 38 °C and 7 bara. This hydrogen product is suitable for use in a high-temperature fuel cell such as a molten carbonate (MCFC) or solid oxide fuel cell (SOFC). Trace amounts of CO and CO2 in the hydrogen product exiting the methanator would make it unsuitable for use in low-temperature fuel cells (polymer electrolyte or alkaline fuel cell) without additional cleanup in a pressure swing adsorption (PSA) unit. In this analysis, it has been assumed that the fuel cell would need to operate with pure oxygen instead of air. The final selection of fuel cell technology would dictate, however, the choice between air and oxygen. The feedstock for the plants is landfill gas consisting of an equimolar mixture of CH4/CO2 and available at 20 °C and 10 bara. A small amount of the hydrogen product is recycled back to the reformer feed to prevent deactivation of the reforming catalyst. This has been accounted for in the energy balance. The energy usage and operating economics of the various designs are presented in Tables 1 and 2. Steam was produced from boiler feedwater (BFW) at 104 °C, and credit or debit was taken as appropriate. The overall fuel cell efficiency was assumed to be 80%. The landfill gas has a lower heating value (LHV) of 13 370 kJ/kg.

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The energy operating costs for the various plant designs have been determined based on the following prices of the commodities involved: landfill gas $1.90/GJ LHV steam 40 barg $7.28/1000 kg 14 barg $6.61/1000 kg 3.5 barg $5.51/1000 kg note: the above values reflect the cost of the boiler feedwater electricity $0.06/kWh oxygen $0.22/kg

A detailed description of the various designs is presented in the next section. Process Description Case 1: Conventional Hydrogen Plant. The process flowsheet for this case is shown Figure 1a. The case 1 design is that of a conventional plant typically used for the production of hydrogen, utilizing a catalytic packed-bed reformer, high- and low-temperature shift reactors, a monoethanolamine (MEA) CO2 removal unit, and a methanator. Steam is generated from the boiler feedwater in the convection section of the furnace and mixed with the preheated landfill gas feedstock such that the steam-to-methane ratio becomes 5:1. The mixture is then fed to the radiant section of the primary reformer (H-1), where it is reacted over a nickel-based supported catalyst. The basic reactions that occur are

reforming CH4 + H2O T CO + 3H2

∆H°298K ) 206 kJ/mol

water gas shift CO + H2O T CO2 + H2

∆H°298K ) -41 kJ/mol

For this study, the need to desulfurize the feed gas has been ignored, because the level of cleanup required is probably the same for all plant designs. The primary reformer is assumed to operate with a typical commercial methane space velocity of 1000 h-1. The conversion of methane in an industrial reformer is typically limited by thermodynamic equilibrium. For the reforming and shift reactions, the approaches to equilibrium, ∆TR and ∆TS, are defined as

KR(TP + ∆TR) )

KS(TP + ∆TS) )

PCOPH23 PCH4PH2O PCO2PH2 PCOPH2O

where TP is the reformer outlet temperature, KR and KS are the equilibrium constants for the reforming and shift reactions, respectively, and Pi represents the partial pressure of species i in the reformer effluent. In conventional plants, a -25 °C approach to equilibrium is typically found for the reforming reaction, while the faster shift reaction reaches equilibrium. The kinetics of the nickel reforming catalyst were fitted to the Xu and Froment model,19 with parameters fitted so as to produce the above approaches to equilibrium. These were then further used with the MR models.

For the purity requirement of 97%, the required primary reformer effluent outlet temperature was 800 °C, which, at 10 bara pressure, corresponded to a 95.5% conversion of methane. Because the water gas shift reaction is exothermic, the reformer effluent was cooled to 400 °C by generation of 40 barg steam in exchanger E-1 prior to its introduction to the high-temperature shift (HTS) converter R-1. The effluent from the HTS was then cooled to 240 °C by generation of 14 barg steam in exchanger E-2 prior to introduction to the lowtemperature shift (LTS) converter R-2. The composition of the LTS effluent was 43% H2, 35% H2O, 21% CO2, 0.5% CH4, and 0.3% CO. Both the HTS and LTS reactors were assumed to operate adiabatically and achieve equilibrium. To meet the hydrogen purity specification, the CO2 in the LTS effluent was removed via a MEA absorber/ regenerator system. The choice of CO2 removal process between an amine- or solvent-based process and a PSA unit is dictated by the purity requirement of the hydrogen product. For hydrogen product purities below 99%, amine-based processes still find widespread use. The LTS effluent was cooled to 38 °C by first exchanging heat with the methanator feed stream in E-3 and then by exchanging heat with cooling water in exchanger E-4. After removal of the condensed water in flash drum F-1, the hydrogen-rich mixture was fed to the MEA absorber T-1. In the absorber, the CO2 concentration was reduced to 1200 ppm by absorption with MEA. The rich amine from the MEA absorber was heated by exchange with the lean amine from the MEA regenerator T-2 prior to its introduction to the MEA regenerator. The rich amine was then stripped to produce the CO2 product overhead in the regenerator, with the lean amine from the bottom of the regenerator cooled in exchangers E-7 and E-8 prior to its introduction to the MEA absorber. The overhead stream from the MEA absorber was heated in exchanger E-3 and fed to the methanator R-3, where the remaining carbon oxides were converted to methane. The methanator was assumed to operate adiabatically and achieve equilibrium. The hydrogen product from the methanator was cooled to 38 °C before being sent off-plot. As shown in Table 1, the amount of feed used was equivalent to 106.0 GJ/h (the landfill gas has an LHV of 13 370 kJ/kg), while the fuel to the reformer was 82.4 GJ/h. After credit is taken for steam generation at the 40 and 14 barg levels (20.74 and 9.47 GJ/h, respectively) and accounting for the usage of 3.5 barg steam (72.21 GJ/h, almost all of which is due to the MEA regenerator reboiler), the energy required to make 1 kgmol of hydrogen product was 462.68 MJ. Case 1a: Conventional Hydrogen Plant with a CO2 Removal Module. The process flowsheet for this case is shown in Figure 1b. The design is essentially the same as that of case 1 with the exception that a CO2 removal module R-4 has been added upstream of the MEA absorber. This CO2 removal module is equipped with carbon membranes that have a CO2/CH4 selectivity of 35:1. The preparation and separation characteristics of these membranes are discussed elsewhere.12-18 These membranes, which are typically formed by the pyrolysis of polyimide or similar polymeric precursor layers on a porous substrate, offer a combination of high permeability and high selectivity. The module may be configured with the membranes in cylindrical tubes in the fashion of a shell-and-tube

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Figure 1. (a) Conventional hydrogen plant. (b) Conventional hydrogen plant with a CO2 removal unit.

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Table 3. Transport Characteristics of High-Temperature Ceramic Membranes seperation factor with respect to methane Pi/PCH4 CH4 permeance, H2 H2O CO CO2 cm3/cm2 min psi case 2 and 2a designs 2.82 0.94 0.76 case 4 autoreformer 30 0.94 0.76 design

0.6 0.6

0.8 0.75

heat exchanger or as a series of flat, parallel plates. The permeate side of the vessel is swept with air, and the permeate is assumed to be disposed of as part of the combustion air fed to the primary reformer H-1. The CO2 removal module is designed to reduce the CO2 partial pressure of the feed gas to 0.35 bar, which translates to 86.9% removal of the inlet CO2. This level of reduction was chosen so as to provide a reasonable sweep air flow rate and membrane area. As shown in Figure 1b, this dramatically reduces the required duties and utilities of the MEA section of the plant versus the case 1 design. Because a small amount of hydrogen was lost in the permeate stream, some additional landfill gas feedstock (with respect to case 1) is required in order to produce the 10 MMSCFD of high-pressure hydrogen product. As a result of the additional feedstock required, slightly more feed and fuel is required for this case versus case 1 (feed, 110.8 vs 106.0 GJ/h; fuel, 86.5 vs 82.4 GJ/h). Steam generation credits at the 40 and 14 barg levels remain about the same, but the usage of 3.5 barg steam has been reduced dramatically by the CO2 removal unit (9.50 vs 72.21 GJ/h). As a result, the energy required to produce 1 kgmol of hydrogen product is reduced from 462.68 to 349.75 MJ. This dramatic improvement is directly attributable to the carbon membrane CO2 removal unit. Case 2: Primary Reformer Equipped with an Alumina Membrane. The process flowsheet for this case is shown in Figure 2a. This process design differs from the case 1 design by the addition of a MR using a commercial mesoporous alumina membrane in place of the conventional plug-flow primary reformer. While these membranes exhibit good mechanical and thermal stability, they also allow gases other than hydrogen to permeate. The permeabilities and separation factors for this membrane are summarized in Table 3. These membranes, with pore sizes in the mesoporous range, exhibit Knudsen separation factors (i.e., fluxes inversely proportional to the square root of their molecular weight). For this study, the MR is operated with the same methane space velocity as the other cases, 1000 h-1. The kinetics of the nickel reforming catalyst used were the same as those in case 1. The transport of the various species across the membrane was calculated using a modified Fick diffusion model as proposed by Veldsink et al.20 The flux of species i through the membrane may be written as

Ni )

(

)

-1 eff ∂(xiP) Bo ∂P + xiP D RT i ∂r µ ∂r

where Ni is the flux of species i [mol/(m2 s)], xi is the mole fraction of species i, r is the radial distance through the membrane, P is the pressure, µ is the viscosity, T is the temperature, and R is the universal gas constant [8.314 J/(mol K)]. The parameter Bo is a membrane-

specific parameter fitted from single and mixed gas diffusivity experiments. The effective diffusivity, Deff i , may be measured experimentally or may be calculated in the transition region from the Bosanquet formula

Deff i )

1 1 1 + eff D Di,m i,K

where the mixture and Knudsen diffusivities are calculated from

Di,K )

eff Di,m

)

( )x

4 rp 3τ 2

8RT πMi

n

1

Deff ∑ ij xj (1 - x )j)l,j*i i

 o Deff ij ) Dij τ where  is the membrane porosity, τ is the membrane tortuosity, rp is the effective membrane pore diameter, and Doij are the binary diffusivities, which may be determined experimentally or calculated from the Chapman-Enskog correlation. Integration of this flux equation together with the chemical reaction along the length of the membrane is carried out in a stepwise fashion. The specific area of the membrane used in the model was 17.95 cm2 of membrane/cm3 of catalyst. This design produces two separate hydrogen product streams: the high-pressure hydrogen product of the previous designs and a lowpressure hydrogen product produced on the permeate side of the MR. The permeate side of the MR is swept with low-pressure (3.5 barg) steam in the ratio of 2:1 mol of sweep steam/mol of CH4 fed. The steam in the permeate mixture is condensed using cooling water, and the low-pressure hydrogen is used in a high-temperature fuel cell. The feed gas side effluent is processed as in case 1 except that the steam-to-methane ratio was 3:1. Thus, the total steam to the reformer is the same as that in case 1. Because of the loss of hydrogen to the permeate side of the membrane, approximately 96% more landfill gas feedstock is used than for case 1. This also required additional fired duty for the primary reformer (energy usage: feed, 225.5 vs 106.0 GJ/h; fuel, 155.8 vs 82.4 GJ/h in the case 1 design). In addition, the increased concentration of carbon oxides on the feed gas side places greater duties on the MEA section of the plant. The increased carbon oxides on the feed gas side necessitate greater methane conversion than that for case 1; thus, the outlet temperature of the primary reformer does not change, although the membrane provides slightly greater conversion than in case 1. The increased flow of carbon oxides also adds a large load on the MEA section of the plant (3.5 barg of steam usage: 141.0 vs 72.21 GJ/h in case 1). As with the previous cases, the HTS, LTS, and methanator were assumed to operate adiabatically and achieve equilibrium. The low-pressure hydrogen, when reacted with oxygen in a high-temperature fuel cell, produces 114.1 GJ/h worth of electricity (assuming 80% overall fuel cell

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Figure 2. (a) Reformer equipped with an alumina membrane. (b) Reformer equipped with an alumina membrane and a CO2 removal unit.

efficiency). However, this electrical generation does not offset the increased energy requirements elsewhere in

the plant, and the energy cost of the hydrogen product is 744.65 MJ/kgmol. This inefficiency is a consequence

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Figure 4. Cost of product hydrogen as a function of membrane permeance and selectivity for the case 4 autoreformer design. Figure 3. Autoreformer design.

of the poor selectivity of the presently available porous ceramic membranes. Case 2a: Primary Reformer Equipped with an Alumina Membrane and a CO2 Removal Module. The process flowsheet for this case is shown in Figure 2b. This process design is essentially the same as that for case 2 except for the addition of the CO2 removal module described in the discussion of case 1a. Again, because of some hydrogen losses in the CO2 removal module, slightly more feed and fuel are required relative to case 2. Because the ceramic membrane reformer produces a tube-side effluent with a greater concentration of CO2, the changes in the MEA section utility requirements between cases 2 and 2a are even greater than those between cases 1 and 1a. Designing the CO2 removal module to reduce the CO2 partial pressure to 0.35 bar prior to the MEA unit translates to a 96.8% removal of CO2. As a result, the energy requirement is reduced to 499.35 MJ/kgmol, almost the same as that of the conventional hydrogen plant. To compare the effect of the ceramic membrane reformer directly, cases 3 and 3a are those of a conventional hydrogen plant and hydrogen plant equipped with a CO2 removal module scaled-up to produce enough excess hydrogen to produce 11.41 GJ/h of electricity in a fuel cell. Thus, the direct comparison of the energy requirements of case 2 with case 3 (744.65 MJ/kgmol of H2 product vs 677.01 GJ/kgmol) and case 2a with case 3a (499.4 vs 456.45 MJ/mol) shows that the membranes presently available for high-temperature reforming do not provide an improvement over conventional plants or conventional plants equipped with a CO2 removal module. Case 4: Autoreformer Case. The process flowsheet for this case is shown in Figure 3. The comparison of cases 2 and 2a with cases 3 and 3a shows that the hightemperature ceramic membranes presently available commercially (which achieve only Knudsen separation factors) do not permit the design of more efficient hydrogen plants (at least not with a space velocity of 1000 h-1). However, sol-gel, CVD, or other surface modification techniques currently under study for producing smaller pore size membranes may eventually result in materials with high H2/CH4 separation factors. A membrane with a large separation factor would then permit the design of a unique hydrogen plant. The relationship between the membrane’s selectivity, as expressed by the separation factor of the various

species, and the permeance of the membrane is critical to the autoreformer design. Membrane permeance is defined as (assuming a tubular membrane)

Pj )

Vr 2πR1L(Pt - Ps)

where Pj is the species permeance [cm3 of gas at STP/ (cm2 min psi)], Vr is the volumetric flow of gas across the membrane (cm3/min), R1 is the membrane radius (cm), L is the membrane length (cm), and Pt and Ps are the tube- and shell-side partial pressures of species j, respectively. Dense metallic membranes, while producing a pure hydrogen product, have very low permeances (roughly 1/ 1000th of the commercially available membranes discussed here). Thus, to achieve significant improvements in methane conversion, either a high ratio of membrane area-to-catalyst volume or very low space velocities (∼10 h-1) must be used. Both of these options would make a commercially sized plant prohibitively expensive. At commercial space velocities (∼1000 h-1), a 100-fold increase in membrane permeance would be required either to produce significant increases in methane conversion or to permit operation at a lower temperature for a given methane conversion. Producing a dense metallic membrane that is sufficiently thin, thermally stable, and pinhole-free capable of this permeance still remains a challenge. Commercially available ceramic membranes have permeances that would not require excessive membrane area or operation at lower space velocities, if their selectivity toward hydrogen could be improved to reduce methane slip across the membrane to acceptable levels. The commercially available membranes have pore sizes on the order of 50 Å. For such large pores the selectivity between species is limited to the Knudsen ratios (i.e., inversely proportional to the square root of the molecular weights). However, sol-gel and CVD techniques may in the future produce membranes with permeances comparable to the current commercial membranes but sufficiently small pore sizes to selectively permit only hydrogen to pass through. In the design shown in Figure 3, a membrane with a H2/CH4 separation factor of 30 is utilized (separation factors for other species are summarized in Table 3). This membrane is assumed to consist primarily of small micropores, with a small number of mesopores that

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Figure 5. Cost of product hydrogen as a function of the oxygen cost.

Figure 6. Cost of product hydrogen as a function of the price of electricity.

permit the transport of species other than hydrogen. A high-pressure sweep steam is used on the permeate side as a sweep gas in the ratio of 3:1 mol of sweep steam/ mol of CH4 fed. The landfill gas is mixed with steam such that the tube-side feed to the reactor has a H2Oto-CH4 ratio of 2. The high separation factor permits a 97% pure H2 product to be produced by condensation of the sweep steam on the permeate side. The tube-side

effluent, from which roughly 85% of the hydrogen has been recovered, may be used as fuel to the reformer itself. This effluent stream provides about half of the fired duty of the reformer, with the rest of the required duty provided by landfill gas. As shown in Table 1, this process has a low energy requirement (415.44 MJ/ kgmol), has low capital requirements, and is extremely simple to operate because the plant is in steam balance.

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Figure 7. Cost of product hydrogen as a function of the landfill gas price.

Note that for the autoreformer design not only is it necessary that the membrane have sufficient selectivity to meet the hydrogen purity requirement but the membrane must also have sufficient permeance to allow operation at commercially acceptable space velocities. In addition to the autoreformer becoming excessively large, without sufficient permeance the tube-side effluent from the reactor would contain too much hydrogen, providing excess fuel to the autoreformer. This excess energy would be either lost or used to produce export steam. Membrane permeance in excess of this minimum would allow operation of the autoreformer at a lower temperature, reducing the required fired duty. Improvement of the membrane’s selectivity toward hydrogen also reduces the energy requirement of the process, though to a lesser extent than an increase in permeance. This is shown in Figure 4 for an autoreformer for membranes with hydrogen-to-methane selectivities of 30, 200, and perfect hydrogen selectivity, correspondingly, operating with a space velocity of 1000 h-1. Using the same geometry as the commercially available ceramic membranes, a minimum hydrogen permeance of 1.0 cm3/(cm2 min psi) was required; lesser permeances translated into excessive export steam production. One of the advantages of the autoreformer design is that the plant may be operated efficiently in steam balance (i.e., without import or export of steam), provided the membrane has both sufficient selectivity and permeance. Overall Evaluation of the Various Designs The energy requirements for all of the different designs are summarized in Table 1. The cost requirements for the various designs are shown in Table 2.

Comparing the various designs, one concludes that the lowest cost product is produced by case 1a, the conventional hydrogen plant with the CO2 removal module, while the highest cost product is produced by the case 2 design, utilizing a MR equipped with a mesoporous membrane as the primary reformer. The values in these tables and the conclusion of which process design is most efficient are strongly dependent on the assumption for the prices of the various commodities involved (the values for the base case are typical for the design of a grassroots plant), which may vary significantly among various locations, and for more integrated facilities. A $0.01/KWH increase in the price of electricity, for example, greatly favors membrane reformer plants, while a comparable increase in the cost of oxygen makes all of the fuel cell designs noncompetitive. The effects of different oxygen, electrical, and fuel costs are shown in Figures 5-7. Figure 5 shows the cost of the product hydrogen as a function of oxygen price (the price of all other utilities are shown above). Because the case 1, 2, and 4 designs do not generate electricity, the cost of oxygen has no effect on the cost of the product hydrogen. As shown in the Figure 5, for oxygen prices above $0.215/kg, the case 1a design provides the lowest cost hydrogen, while for oxygen costs less than $0.215/ kg, the case 3a design provides the lowest cost design. Interestingly, for oxygen costs below $0.176/kg, the conventional hydrogen plant (case 1) is the most expensive design and all of the fuel-cell-equipped plants (cases 2, 2a, 3, and 3a) produce a cheaper hydrogen product than the nonelectrical-generating plants. For oxygen costs above $0.22/kg, the electrical-generating designs are less attractive than the nonelectrical-generating designs.

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Similar effects for the sale price of electricity on hydrogen product cost are shown in Figure 6. This is not surprising, because it is the relative cost of electricity versus oxygen that determines the profitability of the fuel-cell-equipped plants. The effect of landfill gas price is shown in Figure 7. As shown in the figure, the case 2 (membrane-equipped reformer) and case 3 designs (conventional plant equipped with a fuel cell) provide the highest hydrogen costs, while the case 1a design provided the lowest cost hydrogen over the landfill gas price range of $1.60-3.16/GJ LHV. Below $1.64/GJ LHV, the case 3a design provides the lowest cost product. Obviously, the landfill gas price and the cost of oxygen and electricity greatly affect the product hydrogen cost. As a result, the most efficient design depends on the local conditions. It should be noted that capital costs have not been included in this analysis, nor has operation of the plants at space velocities significantly different from those typically found in the industry been considered. As a result, plant designs for small amounts of hydrogen, such as those for homes equipped with fuel cells to produce electricity for electrical cars, may operate at much lower space times and hence favor the ceramic membrane-equipped plants. The case 4 (autoreformer) design never produced the lowest cost hydrogen product in this analysis, but its energy operating cost was always less than a conventional hydrogen plant and has the prospect of large capital savings and ease of operation. What may be said without hesitation is that the carbon membrane-based CO2 removal unit plants provided significant cost savings over conventional amine-based units. Literature Cited (1) Sun, Y.; Khang, S. Catalytic Membrane for Simultaneous Chemical Reaction and Separation Applied to a Dehydrogenation Reaction. Ind. Eng. Chem. Res. 1988, 27, 7. (2) Tellez, C.; Menendez, M.; Santamaria, J. Oxidative Dehydrogenation of Butane Using Membrane Reactors. AIChE J. 1997, 43, 3. (3) Sloot, H. J.; Smolders, C. A.; van Swaaij, W. P. M.; Versteeg, G. F. High-Temperature Membrane Reactor for Catalytic GasSolid Reaction. AIChE J. 1992, 38, 6. (4) Shu, J.; Grandjean, B. P. A.; Kaliaguine, S. Methane Steam Reforming in Asymmetric Pd- and Pd-Ag/Porous SS Membrane Reactors. Appl. Catal. 1994, 119, 305. (5) Jorgensen, S.; Hojlund-Nielsen, P. E.; Lehrmann, P. Steam Reforming of Methane in a Membrane Reactor. Catal. Today 1995, 25, 303.

(6) Chai, M.; Machida, M.; Eguchi, K.; Arai, H. Promotion of Hydrogen Permeation on Metal-Dispersed Alumina Membranes and its Application to a Membrane Reactor for Methane Steam Reforming. Appl. Catal. 1993, 110, 239. (7) Vasileiadis, S. P. Ph.D. Dissertation, University of Southern California, Los Angeles, CA, 1994. (8) Aasberg-Petersen, K.; Nielsen, C.; Jorgensen, S. Membrane Reforming for Hydrogen. Catal. Today 1998, 46, 193. (9) Prabhu, A. K.; Radhakrshnan, R.; Oyama, S. T. Support Nickel Catalysts for CO2 Reforming of Methane in Plug Flow and Membrane Reactors. Appl. Catal. A 1999, 183, 2. (10) Galuszka, J.; Pandey, R. N.; Ahmed, S. Methane Conversion to Syngas in a Palladium Membrane Reactor. Catal. Today 1998, 46, 2-3. (11) Ostrowski, T.; Giroir-Fendler, A.; Mirodates, C.; Mleczko, L. Comparative Study of the Catalytic Partial Oxidation of Methane to Synthesis Gas in Fixed Bed and Fluidized Bed Membrane Reactors: Part I: A Modelling Approach. Catal. Today 1998, 40, 2-3. (12) Jones, C. W.; Koros, W. J. Carbon Molecular Sieve Gas Separation Membranes-I. Preparation and Characterization Based on Polyimide Precursors. Carbon 1994, 32, 8. (13) Geiszler, V. C.; Koros, W. J. Effects of Polyimide Pyrolysis Conditions on Carbon Molecular Sieve Membrane Properties. Ind. Eng. Chem. Res. 1996, 35, 2999. (14) Rao, M. B.; Sircar, S. Nanoporous Carbon Membranes for Separation of Gas Mixtures by Selective Surface Flow. J. Membr. Sci. 1993, 85, 253. (15) Hayashi, J.; Yamamoto, M.; Kusakabe, K.; Morooka, S. Simultaneous Improvement of Permeance and Permselectivity of 3′,3′,4,4′-Biphenyltetracarboxylic Dianhydride-4,4′-Oxydianiline Polyimide Membrane by Carbonization. Ind. Eng. Chem. Res. 1995, 34, 12. (16) Chen, Y. D.; Yang, R. T. Preparation of Carbon Molecular Sieve Membranes and Diffusion of Binary Mixtures. Ind. Eng. Chem. Res. 1994, 33, 3146. (17) Sedigh, M. G.; Xu, L.; Tsotsis, T. T.; Sahimi, M. Transport and Morphological Characteristics of Polyetherimide-Based Carbon Molecular Sieve Membranes. Ind. Eng. Chem. Res. 1999, 38, 3367. (18) Sedigh, M. G.; Onstot, W. J.; Xu, L.; Peng, W. L.; Tsotsis, T. T.; Sahimi, M. Experiments and Simulations of Transport and Separation of Gas Mixtures in Carbon Molecular Sieve Membranes. J. Phys. Chem. 1998, 102 (49), 8580. (19) Xu, J.; Froment, G. F. Methane Steam Reforming, Methanation, and Water-Gas Shift: I. Intrinsic Kinetics. AIChE J. 1989, 35, 1. (20) Champagnie, A. M. Ph.D. Dissertation, University of Southern California, Los Angeles, CA, 1991.

Received for review March 31, 2000 Accepted September 21, 2000 IE0003685