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Experimental and simulation study on a catalyst packed tubular dense membrane reactor for partial oxidation of methane to syngas
Chemical Engineering Science 55 (2000) 2617}2625
Experimental and simulation study on a catalyst packed tubular dense membrane reactor for partial oxidation of methane to syngas W. Jin, X. Gu, S. Li, P. Huang, N. Xu*, J. Shi Membrane Science and Technology Research Center, Nanjing University of Chemical Technology, Nanjing 210009, People's Republic of China Received 10 March 1999; received in revised form 18 October 1999; accepted 20 October 1999
Abstract A tubular perovskite-type La Sr Co Fe O membrane reactor packed with a Ni/Al O catalyst and an isothermal \d model for this reactor have been developed for the partial oxidation of methane to syngas. CH conversion decreased and CO selectivity remained essentially constant with increasing of carbon space velocity. In 1098}1158 K, the CH conversion was larger than 96% with the CO selectivity larger than 97% at low carbon space velocity. The simulated results showed good agreement with the experimental data. The model simulated the e!ects of air #ow rate and membrane reactor length on the performance of membrane reaction, which suggested optimal operating conditions for the syngas production in the membrane reactor. 2000 Elsevier Science Ltd. All rights reserved. Keywords: Membrane reactor; Tubular dense perovskite-type membrane; Methane partial oxidation; Syngas; Reactor model; Simulation
1. Introduction The partial oxidation of methane to syngas (CH #1/2O "CO#2H , DH "!36 kJ/mol) as a promising process for the production of syngas has been extensively studied (Prettre, Eichner & Perrin, 1946; Ashcroft et al., 1990; Dissanayake, Rosynek, Kharas & Luneford, 1991; Choudhary, Rajaput & Prabhakar, 1992; Hickman & Schmidt, 1993a,b) due to the more favorable H to CO ratio in the product gas as well as the mild exothermicity of the reaction. In this process, although air is usually used as the oxygen source, downstream processing requirements can not tolerate nitrogen, so either nitrogen must be separated from the syngas product mixture, or oxygen must separated from the air (usually cryogenically) before being fed to the syngas reactor. Since upstream nitrogen separation from air would be more favorable than costly downstream puri"cation, an oxygen separation plant would be neces-
* Corresponding author. Tel.: #81-025-3316755; fax: #81-02533}345. E-mail address: [email protected] (N. Xu).
sary. Accordingly, the most signi"cant cost associated with conventional partial oxidation of methane to syngas is that of the oxygen separation plant. Recently, inorganic membrane reactors have been drawing considerable attention due to their ability to carry out simultaneous reaction and separation (Armor, 1989; Bhave, 1991; Zaman & Chakma, 1994). There are basically two types of inorganic membranes, which can be used for membrane reactors: dense and porous membranes. One kind of dense membrane reactors utilizes oxygen semi-permeable ion (or mixed) conducting dense ceramic membranes (Nozaki, Yamazaki & Omata, 1992; Eng & Stoukides, 1991; Wang & Lin, 1995; Balachandran et al., 1995,1997; Tsai, Dixon, Moser & Ma, 1997; Xu & Thomson, 1997; Zeng, Lin & Swartz, 1998; Li, Jin, Huang, Xu, Shi & Lin, 2000a,b; Li et al., 1999c). Since the concentration of oxygen can be controlled by adjusting the feed pressure of oxygen, this kind of oxygen-selective dense membrane reactor has a signi"cant advantage over the "xed-bed reactor. It is possible to maintain the oxygen concentration at a uniform value. For partial oxidation reaction of methane, the supply of oxygen can have a great e!ect on product selectivities (Otsuka, Yokoyama & Morikawa, 1985; Kao, Lei & Lin, 1997). More recent
0009-2509/00/$ - see front matter 2000 Elsevier Science Ltd. All rights reserved. PII: S 0 0 0 9 - 2 5 0 9 ( 9 9 ) 0 0 5 4 2 - 4
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Nomenclature
r
A G E G F H
r r ¹ o v HG
J H K G l ¸ P H P Q P R Q I r R R H
frequency factor of reaction i, mol/g -s-Pa activation energy of reaction i, kJ/mol molar #ow rate of species i in tube side, mol/s permeation #ux of component j, mol/m-s equilibrium constant of reaction i axial distance of dense membrane reactor, m total length of dense membrane reactor, m partial pressure of component j, atm total pressure in the shell side of membrane reactor, Pa total pressure in the tube side of membrane reactor, Pa molar #ow rate of species k in shell side, mol/s radial distance of dense membrane reactor, cm universal gas constant, 8.314 J/mol K rate of reaction j, mol/s
reports (Balachandran et al., 1995,1997; Tsai et al., 1997) suggested that oxygen-permeating dense membrane reactors have potential application in partial oxidation of methane to syngas. The modeling and simulation of catalytic membrane reactors have also attracted the interest of many investigators over the last two decades (Hsieh, 1991; Saracco & Specchia, 1994; Sanchez & Tsotsis, 1996; Yang, Xu & Shi, 1998). Most studies have focused on particular membrane reactor systems aiming to simulate their performance in terms of attainable yield and selectivities. In a review article, a general mathematical model was developed by Tsotsis, Minet, Champagine and Liu (1993) for the packed-bed catalytic membrane reactor. However, only special cases of this model have been solved and discussed due to the mathematical complexity. For the dense membrane reactor, a model for a packed-bed membrane reactor using a palladium membrane was "rst presented by Itoh (1987) and examined with experimental data. The model assumed isothermal conditions and the catalyst is packed in the tube side only. With the emergence of solid oxide membranes and their use in membrane reactor application, a number of models have appeared recently to model such reactors. Some modeling studies on oxygen-selective dense membrane reactors have been put forward for partial oxidation of methane to syngas and oxidative coupling of methane (Dixon, Moser & Ma, 1994; Wang & Lin, 1995; Nozaki & Fujimoto, 1994; Tsai, 1996; Kao et al., 1997). Among them, only the model developed by Tsai (1996) dealt with the partial oxidation of methane to syngas. The model
inner radius of perovskite dense membrane tube, cm external radius of dense membrane tube, cm inner radius of reactor shell, cm temperature, K bulk density of catalytic bed, kg/m stoichiometric coe$cient of component j of reaction i
Subscripts i j, k, m
ordinal number of reactions ordinal number of gas species
Superscripts 0 ' ''
inlet values upstream of membrane interfaces downstream of membrane interfaces
was set up according to a two-dimensional non-isothermal membrane reactor utilizing a multi-layered tubular membrane. The membrane tube was considered to consist of three layers: a macroporous support layer, an O semi-permeable perovskite dense layer and a porous catalytic layer. The model simulated the operating parameters for the methane oxidation to syngas reaction in bench- and industrial-scale membrane reactors. However, there are no experimental data reported on partial oxidation of methane to syngas in a supported perovskite dense membrane reactor, which was examined in the model of Tsai. This is partly because there is no successful experimental study on perovskite dense membranes on porous substrates up to date, although some e!orts (Ng, Reichert, Schwartz & Collins, 1996; Xia, Ward, Atanasova & Schwartz, 1998) had been made in this "eld. In the foreseeable future, commercial membrane reactors would most likely be based on straight tubular ceramic membranes. Balachandran et al. (1995) is the "rst, and probably the only group who studied the conversion of methane to syngas on a tubular dense ceramic membrane reactor. They reported a methane conversion larger than 98% with a CO selectivity of 90% for partial oxidation of methane to syngas in a tubular non-perovskite SrCo FeO membrane reactor at about 9003C in V the presence of Rh-based reforming catalyst, which is very expensive. The objective of this paper is to study the performance of the partial oxidation of methane to syngas in a tubular dense mixed-conducting perovskite-type membrane reactor packed with a Ni/c-Al O catalyst by experiment
W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625
and simulation. In this study, the perovskite-type oxide La Sr Co Fe O was selected for the tubular \B membrane because of its better chemical stability and high oxygen permeability as reported by Tai, Nasrallan and Anderson (1995) and Xu and Thomson (1998), and also found recently in our laboratory (Li et al., 2000a). This study is focused on identifying the optimum operation conditions for the partial oxidation of methane to syngas in an unsupported dense perovskite-type membrane reactor.
2. Experimental 2.1. Membrane preparation The La Sr Co Fe O powder was syn \B thesized by solid-state reaction of appropriate amounts of La O , Sr(CO ) , Co O , and Fe O (the Second Chemical Industry of Shanghai, purity of 99.9%). The details of powder preparation conditions were outlined in our previous work (Li et al., 2000a). Membrane tubes were prepared by isostatic pressing at a pressure of 25 MPa, which included loading, pressing and ejecting. The green tubes were sintered in air at 1523 K for 5 h in a MoSi furnace (at a heating and cooling rate of, respec tively, 3 K/min and 2 K/min). The "nal tubular membranes were 8 mm in outer diameter, 15 cm in length, and 1.5 mm in wall thickness. The phase development of
Fig. 1. The module of membrane reactor.
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calcined powder and sintered membranes was studied by X-ray di!raction (Rigaku D/MAX-RB di!ractometer, with CuKa}radiation). The relative densities of the sintered membranes were determined by Archimedes method. The prepared membranes had densities at least 90% of the theoretical density with a single perovskite phase. 2.2. Membrane reactor setup and procedures Fig. 1 shows the membrane reactor module used in this study for the partial oxidation of methane to syngas. As shown in Fig. 1, two dense alumina tubes ( 8.5 mm, ID) supported the tubular membrane (17.3 mm long) with a ceramic binder to seal both ends of the membrane with the walls of the alumina tubes. A quartz tube ( 16 mm, ID) surrounding the two alumina tubes formed the shell side of the reactor. The catalyst was packed in the middle of the tubular membrane. The temperature of the reactor module (surrounded by a tubular furnace) was measured by a type-K thermocouple encased in an alumina tube. A microprocessor temperature controller (Model 708PA) was used to control the temperature to within $1 K of the set points. The schematic diagram of the complete reactor setup is shown in Fig. 2. The #ow rates of inlet gases were controlled through mass #ow controllers (Models D077A/ZM). The shell side of reactor was exposed to air (100 cm (STP)/min), while the tube side was exposed to diluted methane (with helium). Before the reaction experiment was carried out, the tube side of reactor was "rst purged with helium until the air was completely replaced, and then methane was introduced into the system. Both shell side and tube side of the reactor were maintained at the atmospheric pressure. The e%uent streams were analyzed by an on-line gas chromatograph (Model Shimadzu GC-7A). A molecular sieve 5A column (60/80 mesh, 2 m;3 mm, Shimadza) was used for separation of H , O , N , CH and CO, and a GDX-103 column (60/80 mesh, 2 m;3 mm, Tianjin 2nd Chemical Reagent Plant, China) was used for separation of CO ,
Fig. 2. Schematic diagram of membrance reaction apparatus: (1) gas cylinder; (2) gas dryer; (3) valve; (4) mass #ow controller; (5) mixer; (6) pressure gauge; (7) pressure sensor; (8) computer; (9) electric furnace; (10) membrane reactor; (11) chromatogram, (12) recorder; (13) bubble #ow meter.
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hydrocarbons and H O. Gas leakage through the sea lant, if present, could be detected by monitoring the nitrogen concentration in the e%uent from the tube side. The conversion of CH and selectivity of CO were de "ned, respectively, as follows:
Steam reforming: CH #H O & CO#3H P P !- & R "A P P exp(!E /R¹) 1! . !& & K P P !& & (4)
F !F !& , X " !& !& F !&
(1)
F !. S " !- F !F !& !&
(2)
Carbon dioxide reforming: CH #CO & 2CO#2H : P P !- & . R "A P P exp(!E /R¹) 1! !& !K P P !& !(5)
Blue}green 17.5% NiO/c-Al O catalyst was pur chased from Lanzhou Institute of Chemical Physics of Chinese Academy of Sciences. The catalyst used in the membrane reactor was pretreated in 1 : 1 mixture of H : He for 3 h at 7003C. The XRD analysis con"rmed that the NiO was completely reduced to the metal Ni by the H treatment. The Ni/c-Al O catalysts have the BET surface area of 97.62 m/g.
The equilibrium constants K , which are temperature G dependent, were calculated from thermodynamic properties (Shi, Wang, Yu & Chen, 1996) for i"2, 3. The values of activation energy E (i"1, 2, 3) were chosen from the G literature for each of the reactions (Arai, Yamada & Eguchi, 1986; Ross & Steel, 1973; Richardson & Paripatydar, 1990). The frequency factors A are obG tained by "xed-bed reactor experiment, which was conducted in our laboratory using reduced Ni/Al O (typically, CH : O "1.78 : 1). The kinetic constants are given in Table 1.
3. Kinetics and reactor mathematical model
3.2. Oxygen permeation rate in La0 6 Sr0 4 Co0 2 Fe0 8 O3
\d tubular membrane
3.1. Kinetic equations of methane oxidation to syngas over a Ni/Al2 /O3 catalyst Two main mechanisms of the partial oxidation of methane to syngas, depending on the catalyst, have been proposed. One is the mechanism of methane combustion followed by both steam and carbon dioxide reforming (Ashcroft et al., 1990), and the other is the catalytic CH pyrolysis followed by H desorption and carbon oxida tion (Hickman & Schmidt, 1993a,b). Dissanayake et al. (1991) examined the oxidation state and phase composition of the Ni/Al O catalyst as a function of axial position in the catalyst bed. Their results strongly supported the mechanism of methane combustion. Blank, Wittrig and Peterson (1990) also assumed this mechanism in their work by using a Ni/Al O catalyst. In contrast, Hickman and Schmidt (1993a,b) proposed the mechanism of the catalytic CH pyrolysis to explain the high selectivity of syngas formation using Pt and Rh as catalysts. They concluded that, under CH -rich atmo sphere and at high temperatures, H and CO were the primary products of direct oxidation of methane via a pyrolysis mechanism. The mechanism of methane combustion was assumed in the present simulation since a Ni/Al O catalyst was applied in this study, The kinetic equations of Blank et al. (1990) and Tsai (1996) were adapted as follows: Methane combustion: CH #2O PCO #2H O, R "A P P exp(!E /R¹). !&
(3)
The oxygen permeation through a tubular La Sr Co Fe O membrane under air/He con \B dition has been reported in our recent paper (Li et al., 2000a). Based on the experimental data of oxygen permeation, the oxygen permeation #ux can be approximately correlated to oxygen partial pressures and temperature by the following empirical equation: J "(1.37;10\¹#9.77;10\);¹ ;[(P ;10\)\ "\> "\2 !(P ;10\)\ "\> "\2!0.044], (6) where J is the #ux of oxygen permeation based on inner surface area of membrane tube, in cm (STP)/cm/s; ¹ is the temperature, in K; P is the oxygen partial pressure, in Pa; and the superscripts &&and'' represent the upstream and downstream membrane interfaces, respectively. Eq. (6) can be used in the temperature range of 1093}1173 K.
Table 1 Kinetic parameters for the partial oxidation of methane to syngas Reaction i
Activation energy E (J/mol) G
Frequency factor A G (mol/g -s-Pa)
1 2 3
166;10 29;10 23.7;10
1.10 4.19;10\ 2.42;10\
W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625
3.3. Model for dense membrane reactor The tube-and-shell reactor arrangement used in simulation is the same as the experimental arrangement shown in Fig. 3. The diluted methane is fed to the reactor tube side packed with the catalyst, and air is fed concurrently to the shell side of the reactor. Since the partial oxidation of methane to syngas is a mildly exothermic reaction and diluted feed is used, the assumption of an isothermal reactor may be reasonable. The mathematical model in this work is based on the following assumptions: (1) The membrane reactor is operated at steady state under isothermal and isobaric conditions. (2) Homogeneous gas phase oxidation reactions and carbon deposition are negligible. (3) The internal mass transfer resistance of the packed catalytic granules is neglected due to the small catalyst particles (40}60 mesh) and the external mass transfer resistance was also neglected because the Da (Damkoller) number was much less than the one in this work. (4) Ideal gas law is used to describe the gas behavior of single component and gas mixture. According to the principle of mass conservation, the design model equations for the perovskite membrane reactor are derived in each of the three regions as given below. In the tube side, i.e., the catalytic layer: 0(r(r dF L H #2pr J " !nr o l R "0, (7) H PP HG G dl G where the subscript j represents the reaction gas species, i.e., CH , O , CO, CO , H and H O. In the dense membrane: r (r(r r J " J " , (8) r - PP J "0, K
(9)
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where the gas species m includes CH , CO, CO , H , H O and N : In the shell side: r (r(r dQ I !2pr J " "0 (10) I PP dl in which the subscript k denotes O or N , respectively. The initial conditions, l"0, for the di!erent regions are F "F , (11) !& !& F "0, (12) H H$!& Q "Q , (13) Q "Q . (14) , , This set of di!erential equations is coupled together by their boundary conditions as follows: At the catalytic layer/dense membrane layer interface: r"r J "J " (15) - PP for other gas species m (including CH , CO, CO , H and H O): J "0, (16) K F H P, P" (17) H + F R G H where M"6, indicating CH , O , CO, CO , H and H O. At the dense membrane layer/shell side interface: r"r r J " J , (18) r PP J "0, (19) , Q I P. P" (20) I + Q Q G I in which M"2, indicating O and N .
Fig. 3. Schematic diagram of dense membrane reactor.
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These di!erential equations were reduced to a set of dimensionless equations by introducing dimensionless quantities, such as FM "F /F , QM "Q /Q and H H !& I I lM "l/¸. Then the model equations were solved using the numerical Runge}Kutta method.
4. Results and discussion In order to compare the simulating results with the experimental data, in most cases the membrane reaction experiments and simulation were carried out under the same conditions given in Table 2, other parameters will be given in speci"c cases. 4.1. Ewect of carbon space velocity In this case, the helium rate is "xed at 58 ml (STP)/min and the inlet methane rate is changed in the tube side. Fig. 4 shows the e!ect of carbon space velocity in the tube side on the reaction performance at 1158 K. The CH conversion decreases with increasing carbon space veloTable 2 Geometric properties of the dense membrane reactor and operation conditions Pressure for both tube and shell side Bulk density of catalytic bed Virtual length of reactor External diameter of membrane tube Inner diameter of membrane tube Air #ow rate Helium #ow rate
0.1013 MPa 0.726 g/cm 1.73 cm 0.8 cm 0.5 cm 100 ml(STP)/min 58 ml(STP)/min
Fig. 4. E!ect of carbon space velocity on the syngas production at 1158 K Experimental value: (䉫) CH conversion; (䉭) CO selectivity; Simulation value: (*) CH conversion; (- - - -) CO selectivity; (} ) } ) }) H /CO.
city, and the CO selectivity is larger than 90% in all cases. As the carbon space velocity increases, this in turn causes the CH /O ratio to increase. Since the permeation rate of oxygen is constant, the conversion decreases due to insu$cient oxygen for the reaction. Moreover, the carbon deposition on the surface of catalyst occurred due to higher CH /O ratio at the higher carbon space velocity. Dissanayake et al. (1991) found that in the "xed-bed reactor at temperatures *1023 K, the amount of surface carbon formed on the Ni/c-Al O catalyst was primarily in#uenced by the CH /O feed ratio. A high ratio of CH /O could cause deposition of large amounts of carbon that "lled the catalyst pores. In this work, EDS analysis con"rmed the deposition of carbon on the catalyst after use in the membrane reactor for the partial oxidation of methane at a higher carbon space velocity. The deposition of carbon reduced the activity of the catalyst. In addition, the decrease in methane conversion with increasing carbon space velocity is due to the fact that the syngas production reaction order with respect to methane was not "rst order which would give a methane concentration-independent conversion (Blank et al., 1990). All this suggests that the operation of the membrane reaction should be at a low carbon space velocity to obtain high methane conversion. As shown in Fig. 4, the simulating selectivity of CO shows good agreement with the experimental data, and the methane conversion of simulation is a little larger than that of experiment, especially at higher carbon space velocity. This may be attributed to the following reasons. First, the model in this work did not involve the e!ect of radial di!usion of gases on the performance of reaction, and assumed that at the same section of reactor the concentration of each gas is the same. This will result in the calculated rate of reactions being larger than the virtual value. Second, the model did not involve the e!ect of carbon deposition on the performance of reaction. This will lead to simulated methane conversion larger than the experimental data in higher carbon space velocity. Third, in fact, oxygen partial pressure at the wall of membrane tube in this model was mean value, which was lower than realistic oxygen partial pressure, the oxygen #ux calculated from Eq. (6) could be larger than the real oxygen #ux. The increase in oxygen #ux would lead to the increase in methane conversion. In addition, at higher carbon space velocity, the reaction would become radical, and the resistance of di!usion would in#uence the performance of reaction to some extent. Therefore, the di!erence between simulating conversion of methane and experimental data became larger at higher carbon space velocity. 4.2. Ewect of temperature Since the oxygen #ux of the La Sr Co Fe O membrane became usable above 1073 K (Li \B
W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625
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Table 3 E!ect of temperature on the syngas reaction Temp. (3C)
825 850 885
Carbon space velocity (min\)
8.24 8.24 8.24
X (%) !&
S (%) !-
H /CO
Exp.
Model
Exp.
Model
97.3 100 96.7
97.2 99.7 100
97 100 98
98.9 97.9 97.8
et al., 2000a) and the top temperature for the sealing was 1173 K, the experiment was carried out in the temperature range of 1073}1173 K. Table 3 gives the results of the experiment and the simulation at three di!erent temperatures. The experimental data cannot show the trends of changes in CH conversion and CO selectivity with temperatures due to the experimental error, however, the simulating results show that the CH conversion in creases and CO selectivity decreases with increase of temperatures. But, it can be seen that the reaction temperature in the range of 1098}1158 K has no signi"cant e!ect on both the CH conversion and CO selectivity when a lower carbon space velocity is applied. The CH conversion and the CO selectivity are, respectively, larger than 96 and 97% in all cases. These results are similar to the results of the "xed-bed reactor obtained by Dissanayake et al. (1991). This indicates that the temperature range of 1098}1158 K was suitable for the operation of the membrane reactor. From the above comparison between the experimental and simulating results, the model developed in this work was considered to be reasonable. Therefore, the model was used to simulate the e!ects of other parameters on the performance of the syngas reaction. 4.3. Ewect of air yow rate Fig. 5 shows the e!ect of shell side air #ow rate on the performance of the reaction under di!erent carbon space velocities at 1123 K. As shown, the in#uences of air #ow rate on both CH conversion and CO selectivity are insigni"cant when air #ow rate is greater than 100 ml (STP)/min in the shell side. In our previous experimental and modeling study (Li et al. 2000a,c), it was found that under an air/helium gradient the oxygen permeation #ux remained essentially unchanged with varying air #ow rate in the shell side when the air #ow rate was larger than 100 ml (STP)/min. In this work, the oxygen permeation took place under reaction conditions. The oxygen permeation rate under reaction conditions, which could be calculated from the #ow rate and the mole fractions of oxygen-containing product gases (such as O , CO, and H O in the tube side e%uent, was much larger than that under the air/helium gradient at the
Exp. 1.69 1.68 1.96
Model 2 2 2
Fig. 5. E!ect of air #ow rate on CH conversion and CO selectivity: (*) CH conversion; (- - - -) CO selectivity; (䉬䉫) He : CH 20 : 1; (䢇*) He : CH 15 : 1; (䉱䉭) He : CH 10 : 1.
same temperature. This is because the oxygen pressure near the membrane surface in the reducing side is extremely low when a reaction is taking place. However, it can be derived from Fig. 4 that for a certain carbon space velocity oxygen permeation #ux also remains essentially unchanged with varying air #ow rate when air #ow rate is no less than 100 ml (STP)/min. This indicates that the e!ect of air #ow rate on the oxygen permeation under reaction conditions is the same as that under air/helium gradient. This suggests that in the operation of the membrane reactor, it is unnecessary to further increase the air #ow rate if the air supplied is enough for providing stable oxygen permeation #ux. 4.4. Ewect of membrane tube length Simulations in this case were done according to the following parameters: a helium to methane ratio of 10/1, carbon space velocity of 6.614 min\, air #ow rate of 200 ml (STP)/min in the shell side, the membrane tube length of 0.07 m, temperature of 1123 K. Fig. 6 shows the simulation results of CH conversion and CO selectivity
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Fig. 6. E!ect of membrane reactor length on syngas production: (*) CH conversion; (......) CO selectivity; (} ) } ) }) H /CO.
but the CO selectivity remained essentially unchanged with an increased carbon space velocity. In the range of 1098}1158 K, the CH conversion was larger than 96% with the CO selectivity larger than 97% at a low carbon space velocity. An isothermal model developed based on the experiment was also presented to simulate the performance of the membrane reactor under various operating conditions. Good agreement was obtained between the experimental data and simulating results. The simulation results indicate that the temperature range of 1098}1158 K is feasible for the operation of the membrane reactor. The air #ow rate ('100 ml (STP)/min) in shell side shows little in#uence on the performance of syngas reaction, and the CH conversion increases grad ually along the axial position of the membrane reactor and CO selectivity does not decrease (from 100%) until the methane conversion is close to 100%.
Acknowledgements along axial position of the reactor. As shown, the CH conversion increases gradually along the axial position and "nally reaches 100%, remains unchanged, and CO selectivity does not decrease (from 100%) until the methane conversion is close to 100%. The position at which methane conversion reaches 100% corresponds to around 71% of the membrane tube length. Beyond this position, CO selectivity decreases gradually due to the deep oxidation of products, which is caused by excess oxygen near the outlet of the reactor. Fig. 6 also re#ects the pro"les of methane and CO at di!erent axial positions. Compared with the co-feed reactor, oxygen is more uniformly distributed along the reactor axial direction in the membrane reactor. In "xed-bed reactor, the oxygen concentration, with the maximum at the entrance, decreases to zero along the reactor length and the reaction became the sharpest in the front part of the reactor. This could lead to a large adiabatic temperature rise, which may cause reactor runaway. However, in dense membrane reactors the oxygen is fed to the entire reactor length by permeation and the reaction became mild. The methane conversion and CO selectivity depend on the length of the dense membrane reactor for a certain carbon space velocity. In other words, for a "xed length of a membrane reactor, it is necessary to choose a suitable carbon space velocity to obtain high methane conversion and CO selectivity.
5. Conclusions The partial oxidation of methane to syngas was performed in a tubular La Sr Co Fe O mem \B brane reactor in the presence of a pre-reduced NiO/ c-Al O catalyst. The conversion of methane decreased
This work is supported by the National Advanced Materials Committee of China (NAMCC, No. 715-0060120), and the National Natural Science Foundation of China (NNSFC, No. 59789201).
References Arai, H., Yamada, T., & Eguchi, K. (1986). Catalytic combustion of methane over various perovskite-type oxides. Applied Catalysis, 26, 265}276. Armor, J. N. (1989). Catalysis with permselective inorganic membranes. Applied Catalysis, 49, 1}25. Ashcroft, A. T., Cheetham, A. K., Foord, J. S., Green, M. L. H., Grey, C. P. J., Murrell, A., & Vernon, P. D. F. (1990). Selective oxidation of methane to synthesis gas using transition metal catalysis. Nature, 344, 319}321. Balachandran, U., Dusek, J. T., Mieville, R. L., Poeppel, R. B., Klee"sh, M. S., Pei, S., Kobylinski, T. P., Udovich, C. A., & Bose, A. C. (1995). Dense ceramic membranes for partial oxidation of methane to syngas. Applied Catalysis A: General, 133, 19}29. Balachandran, U., Dusk, J. T., Maiya, P. S., Ma, B., Mieville, R. L., Klee"sch, M. S., & Udovich, C. A. (1997). Ceramic membrane reactor for converting methane to syngas. Catalysis Today, 36, 265}272. Bhave, R. R. (1991). Inorganic membranes synthesis, characteristics and applications. New York: Van Nostrand, Reinhold. Blank, R. F., Wittrig, T. S., & Peterson, D. A. (1990). Bidirection adiabatic synthesis gas generator. Chemical Engineering Science, 45, 2407}2413. Choudhary, V. R., Rajput, A. M., & Prabhakar, B. (1992). Low temperature oxidative conversion of methane to syngas over NiO}CaO catalyst. Catalysis Letters, 15, 363}370. Dissanayake, D., Rosynek, M. P., Kharas, K. C. C., & Lunsford, J. H. (1991). Partial oxidation of methane to carbon monoxide and hydrogen over a Ni/Al O catalyst. Journal of Catalysis, 132, 117}127. Dixon, A. G., Moser, W. R., & Ma, Y. H. (1994). Waste reduction and recovery using O permeable membrane reactors. Industrial and Engineering Chemistry Research, 33, 3015}3024.
W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625 Eng, D., & Stoukides, M. (1991). Catalytic and electrocatalytic methane oxidation with solid oxide membranes. Catalysis Reviews Science and Engineering, 33, 375}412. Hickman, D. A., & Schmidt, L. D. (1993a). Steps in CH oxidation on Pt and Rh surfaces: high-temperature reactor simulations. A.I.Ch.E. Journal, 39, 1164}1177. Hickman, D. A., & Schmidt, L. D. (1993b). Production of syngas by direct catalytic oxidation of methane. Science, 259, 343}346. Hsieh, H. P. (1991). Inorganic membrane reactors. Catalysis Reviews Science and Engineering, 33, 1}70. Itoh, N. Y. (1987). A membrane reactor using palladium. A.I.Ch.E. Journal, 33, 1576}1585. Kao, Y. K., Lei, L., & Lin, Y. S. (1997). A comparative simulation study on the oxidative coupling of methane in "xed-bed and membrane reactors. Industrial and Engineering Chemistry Research, 36, 3583}3593. Li, S., Jin, W., Huang, P., Xu, N., Shi, J., and Lin, Y. S. (2000a). Tubular La Sr Co Fe O perovskite membrane for oxygen per \ meation. Journal of Membrane Science, in press. Li, S., Jin, W., Huang, P., Xu, N., Shi, J., and Lin, Y. S. (2000b). Experimental and modeling study on the tubular perovskite type dense membranes for oxygen permeation. A.I.Ch.E. Journal, in press. Li, S., Jin, W., Huang, P., Xu, N., Shi, J., Hu, Michael Z.-C., Payzant, E. A., and Ma, Y. H. (1999c). Perovskite-related ZrO -doped SrCo Fe O membrane for oxygen permeation. A.I.Ch.E. \ Journal, 45, 276}284. Ng, M. F., Reichert, T. L., Schwartz, R. W., and Collins, J. P. (1996). Fabrication of SrCo FeO oxygen separation membranes on \B porous supports. Proceedings of fourth International Conference on Inorganic Membranes, Gatlinburg, TN. Nozaki, T., & Fujimoto, K. (1994). Oxide ion transport for selective oxidative coupling of methane with new membrane reactor. A.I.Ch.E. Journal, 40, 870}881. Nozaki, T., Yamazaki, O., & Omata, K. (1992). Selective oxidative coupling of methane with membrane reactor. Chemical Engineering Science, 47, 2945}2953. Otsuka, K., Yokoyama, S., & Morikawa, A. (1985). Catalytic activity and seletivity control for oxidation coupling of methane by oxygen pumping through yttria-stabilized zircona. Chemistry Letters, 3, 319}322. Prettre, M., Eichner, C. H., & Perrin, M. (1946). The catalytic oxidation of methane to carbon monoxide and hydrogen. Faraday Society Transactions, 42, 335}340. Richardson, J. T., & Paripatydar, S. A. (1990). Carbon dioxide reforming of methane with supported Rhodium. Applied Catalysis, 61, 293}309.
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Ross, J. R. H., & Steel, M. C. F. J. (1973). Mechanism of steam reforming of methane over a coprecipitated Nickel-alumina catalyst. Chemical Society of Faraday Transactions, 69, 10}21. Sanchez, J., & Tsotsis, T. T. (1996). Current developments and future research in catalytic membrane reactors. In A. F. Burggraaf, & L. Cot, Fundamental of inorganic membrane science and technology (pp. 529}568). Amsterdam: Elsevier Science B.V. Saracco, G., & Specchia, V. (1994). Catalytic inorganic membrane reactors: Present experience and future opportunities. Catalysis Reviews Science and Engineering, 36, 305. Shi, J., Wang, G. D., Yu, G. Z., & Chen, M. H. (1996). Chemical engineering handbook (pp. 2}99). Beijing, China: Chemical Industry Press. Tai, L. W., Nasrallan, M. M., & Anderson, H. U. (1995). Thermochemical stability, electrical conductivity, and seebeck coe$cient of Sr-doped LaCo Fe O . Journal of Solid State Chemistry, 118, 117}127. \B Tsai, C. Y. (1996). Perovskite dense membrane reactors for the partial oxidation of methane to synthesis gas. Ph.D. dissertation, Worcester Polytechnic Institute, Worcester, MA, USA. Tsai, C. -Y., Dixon, A. G., Moser, W. R., & Ma, Y. H. (1997). Dense perovskite membrane reactors for the partial oxidation of methane to syngas. A.I.Ch.E. Journal, 43, 2741}2750. Tsotsis, T. T., Minet, R. G., Champagine, A. M., & Liu, P. K. T. (1993). Catalytic membrane reactors. In E. R. Becker, & C. J. Pereira, Computer aided design of catalysts (pp. 471}551). New York: Dekler. Wang, W., & Lin, Y. S. (1995). Analysis of oxidative coupling of methane in dense membrane reactors. Jounal of Membrane Science, 103, 219}233. Xia, C., Ward, T. L., Atanasova, P., & Schwartz, R. W. (1998). Metalorganic chemical vapor deposition of Sr}Co}Fe}O "lms on porous substrates. Journal of Materials Research, 13, 173}183. Xu, S., & Thomson, W. J. (1997). Perovskite-type oxide membranes for the oxidative coupling of methane. A.I.Ch.E. Journal, 43, 2731}2740. Xu, S., & Thomson, W. J. (1998). Stability of La Sr Co Fe O perovskite membrane in reducing and nonreduc \B ing environments. Industrial and Engineering Chemistry Research, 37, 1290}1299. Yang, C., Xu, N., & Shi, J. (1998). Experimental and modeling study on a packed-bed membrane reactor for partial oxidation of methane to Formaldehyde. Industrial and Engineering Chemistry Research, 37, 2601}2610. Zaman, J., & Chakma, A. (1994). Inorganic membrane reactors. Journal of Membrane Science, 92, 1}28. Zeng, Y., Lin, Y. S., & Swartz, S. L. (1998). Perovskite type ceramic membrane: Synthesis, oxygen permeation and membrane reactor performance for oxidative coupling of methane. Journal of Membrane Science, 150, 87}96.
Experimental and simulation study on a catalyst packed tubular dense membrane reactor for partial oxidation of methane to syngas W. Jin, X. Gu, S. Li, P. Huang, N. Xu*, J. Shi Membrane Science and Technology Research Center, Nanjing University of Chemical Technology, Nanjing 210009, People's Republic of China Received 10 March 1999; received in revised form 18 October 1999; accepted 20 October 1999
Abstract A tubular perovskite-type La Sr Co Fe O membrane reactor packed with a Ni/Al O catalyst and an isothermal \d model for this reactor have been developed for the partial oxidation of methane to syngas. CH conversion decreased and CO selectivity remained essentially constant with increasing of carbon space velocity. In 1098}1158 K, the CH conversion was larger than 96% with the CO selectivity larger than 97% at low carbon space velocity. The simulated results showed good agreement with the experimental data. The model simulated the e!ects of air #ow rate and membrane reactor length on the performance of membrane reaction, which suggested optimal operating conditions for the syngas production in the membrane reactor. 2000 Elsevier Science Ltd. All rights reserved. Keywords: Membrane reactor; Tubular dense perovskite-type membrane; Methane partial oxidation; Syngas; Reactor model; Simulation
1. Introduction The partial oxidation of methane to syngas (CH #1/2O "CO#2H , DH "!36 kJ/mol) as a promising process for the production of syngas has been extensively studied (Prettre, Eichner & Perrin, 1946; Ashcroft et al., 1990; Dissanayake, Rosynek, Kharas & Luneford, 1991; Choudhary, Rajaput & Prabhakar, 1992; Hickman & Schmidt, 1993a,b) due to the more favorable H to CO ratio in the product gas as well as the mild exothermicity of the reaction. In this process, although air is usually used as the oxygen source, downstream processing requirements can not tolerate nitrogen, so either nitrogen must be separated from the syngas product mixture, or oxygen must separated from the air (usually cryogenically) before being fed to the syngas reactor. Since upstream nitrogen separation from air would be more favorable than costly downstream puri"cation, an oxygen separation plant would be neces-
* Corresponding author. Tel.: #81-025-3316755; fax: #81-02533}345. E-mail address: [email protected] (N. Xu).
sary. Accordingly, the most signi"cant cost associated with conventional partial oxidation of methane to syngas is that of the oxygen separation plant. Recently, inorganic membrane reactors have been drawing considerable attention due to their ability to carry out simultaneous reaction and separation (Armor, 1989; Bhave, 1991; Zaman & Chakma, 1994). There are basically two types of inorganic membranes, which can be used for membrane reactors: dense and porous membranes. One kind of dense membrane reactors utilizes oxygen semi-permeable ion (or mixed) conducting dense ceramic membranes (Nozaki, Yamazaki & Omata, 1992; Eng & Stoukides, 1991; Wang & Lin, 1995; Balachandran et al., 1995,1997; Tsai, Dixon, Moser & Ma, 1997; Xu & Thomson, 1997; Zeng, Lin & Swartz, 1998; Li, Jin, Huang, Xu, Shi & Lin, 2000a,b; Li et al., 1999c). Since the concentration of oxygen can be controlled by adjusting the feed pressure of oxygen, this kind of oxygen-selective dense membrane reactor has a signi"cant advantage over the "xed-bed reactor. It is possible to maintain the oxygen concentration at a uniform value. For partial oxidation reaction of methane, the supply of oxygen can have a great e!ect on product selectivities (Otsuka, Yokoyama & Morikawa, 1985; Kao, Lei & Lin, 1997). More recent
0009-2509/00/$ - see front matter 2000 Elsevier Science Ltd. All rights reserved. PII: S 0 0 0 9 - 2 5 0 9 ( 9 9 ) 0 0 5 4 2 - 4
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Nomenclature
r
A G E G F H
r r ¹ o v HG
J H K G l ¸ P H P Q P R Q I r R R H
frequency factor of reaction i, mol/g -s-Pa activation energy of reaction i, kJ/mol molar #ow rate of species i in tube side, mol/s permeation #ux of component j, mol/m-s equilibrium constant of reaction i axial distance of dense membrane reactor, m total length of dense membrane reactor, m partial pressure of component j, atm total pressure in the shell side of membrane reactor, Pa total pressure in the tube side of membrane reactor, Pa molar #ow rate of species k in shell side, mol/s radial distance of dense membrane reactor, cm universal gas constant, 8.314 J/mol K rate of reaction j, mol/s
reports (Balachandran et al., 1995,1997; Tsai et al., 1997) suggested that oxygen-permeating dense membrane reactors have potential application in partial oxidation of methane to syngas. The modeling and simulation of catalytic membrane reactors have also attracted the interest of many investigators over the last two decades (Hsieh, 1991; Saracco & Specchia, 1994; Sanchez & Tsotsis, 1996; Yang, Xu & Shi, 1998). Most studies have focused on particular membrane reactor systems aiming to simulate their performance in terms of attainable yield and selectivities. In a review article, a general mathematical model was developed by Tsotsis, Minet, Champagine and Liu (1993) for the packed-bed catalytic membrane reactor. However, only special cases of this model have been solved and discussed due to the mathematical complexity. For the dense membrane reactor, a model for a packed-bed membrane reactor using a palladium membrane was "rst presented by Itoh (1987) and examined with experimental data. The model assumed isothermal conditions and the catalyst is packed in the tube side only. With the emergence of solid oxide membranes and their use in membrane reactor application, a number of models have appeared recently to model such reactors. Some modeling studies on oxygen-selective dense membrane reactors have been put forward for partial oxidation of methane to syngas and oxidative coupling of methane (Dixon, Moser & Ma, 1994; Wang & Lin, 1995; Nozaki & Fujimoto, 1994; Tsai, 1996; Kao et al., 1997). Among them, only the model developed by Tsai (1996) dealt with the partial oxidation of methane to syngas. The model
inner radius of perovskite dense membrane tube, cm external radius of dense membrane tube, cm inner radius of reactor shell, cm temperature, K bulk density of catalytic bed, kg/m stoichiometric coe$cient of component j of reaction i
Subscripts i j, k, m
ordinal number of reactions ordinal number of gas species
Superscripts 0 ' ''
inlet values upstream of membrane interfaces downstream of membrane interfaces
was set up according to a two-dimensional non-isothermal membrane reactor utilizing a multi-layered tubular membrane. The membrane tube was considered to consist of three layers: a macroporous support layer, an O semi-permeable perovskite dense layer and a porous catalytic layer. The model simulated the operating parameters for the methane oxidation to syngas reaction in bench- and industrial-scale membrane reactors. However, there are no experimental data reported on partial oxidation of methane to syngas in a supported perovskite dense membrane reactor, which was examined in the model of Tsai. This is partly because there is no successful experimental study on perovskite dense membranes on porous substrates up to date, although some e!orts (Ng, Reichert, Schwartz & Collins, 1996; Xia, Ward, Atanasova & Schwartz, 1998) had been made in this "eld. In the foreseeable future, commercial membrane reactors would most likely be based on straight tubular ceramic membranes. Balachandran et al. (1995) is the "rst, and probably the only group who studied the conversion of methane to syngas on a tubular dense ceramic membrane reactor. They reported a methane conversion larger than 98% with a CO selectivity of 90% for partial oxidation of methane to syngas in a tubular non-perovskite SrCo FeO membrane reactor at about 9003C in V the presence of Rh-based reforming catalyst, which is very expensive. The objective of this paper is to study the performance of the partial oxidation of methane to syngas in a tubular dense mixed-conducting perovskite-type membrane reactor packed with a Ni/c-Al O catalyst by experiment
W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625
and simulation. In this study, the perovskite-type oxide La Sr Co Fe O was selected for the tubular \B membrane because of its better chemical stability and high oxygen permeability as reported by Tai, Nasrallan and Anderson (1995) and Xu and Thomson (1998), and also found recently in our laboratory (Li et al., 2000a). This study is focused on identifying the optimum operation conditions for the partial oxidation of methane to syngas in an unsupported dense perovskite-type membrane reactor.
2. Experimental 2.1. Membrane preparation The La Sr Co Fe O powder was syn \B thesized by solid-state reaction of appropriate amounts of La O , Sr(CO ) , Co O , and Fe O (the Second Chemical Industry of Shanghai, purity of 99.9%). The details of powder preparation conditions were outlined in our previous work (Li et al., 2000a). Membrane tubes were prepared by isostatic pressing at a pressure of 25 MPa, which included loading, pressing and ejecting. The green tubes were sintered in air at 1523 K for 5 h in a MoSi furnace (at a heating and cooling rate of, respec tively, 3 K/min and 2 K/min). The "nal tubular membranes were 8 mm in outer diameter, 15 cm in length, and 1.5 mm in wall thickness. The phase development of
Fig. 1. The module of membrane reactor.
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calcined powder and sintered membranes was studied by X-ray di!raction (Rigaku D/MAX-RB di!ractometer, with CuKa}radiation). The relative densities of the sintered membranes were determined by Archimedes method. The prepared membranes had densities at least 90% of the theoretical density with a single perovskite phase. 2.2. Membrane reactor setup and procedures Fig. 1 shows the membrane reactor module used in this study for the partial oxidation of methane to syngas. As shown in Fig. 1, two dense alumina tubes ( 8.5 mm, ID) supported the tubular membrane (17.3 mm long) with a ceramic binder to seal both ends of the membrane with the walls of the alumina tubes. A quartz tube ( 16 mm, ID) surrounding the two alumina tubes formed the shell side of the reactor. The catalyst was packed in the middle of the tubular membrane. The temperature of the reactor module (surrounded by a tubular furnace) was measured by a type-K thermocouple encased in an alumina tube. A microprocessor temperature controller (Model 708PA) was used to control the temperature to within $1 K of the set points. The schematic diagram of the complete reactor setup is shown in Fig. 2. The #ow rates of inlet gases were controlled through mass #ow controllers (Models D077A/ZM). The shell side of reactor was exposed to air (100 cm (STP)/min), while the tube side was exposed to diluted methane (with helium). Before the reaction experiment was carried out, the tube side of reactor was "rst purged with helium until the air was completely replaced, and then methane was introduced into the system. Both shell side and tube side of the reactor were maintained at the atmospheric pressure. The e%uent streams were analyzed by an on-line gas chromatograph (Model Shimadzu GC-7A). A molecular sieve 5A column (60/80 mesh, 2 m;3 mm, Shimadza) was used for separation of H , O , N , CH and CO, and a GDX-103 column (60/80 mesh, 2 m;3 mm, Tianjin 2nd Chemical Reagent Plant, China) was used for separation of CO ,
Fig. 2. Schematic diagram of membrance reaction apparatus: (1) gas cylinder; (2) gas dryer; (3) valve; (4) mass #ow controller; (5) mixer; (6) pressure gauge; (7) pressure sensor; (8) computer; (9) electric furnace; (10) membrane reactor; (11) chromatogram, (12) recorder; (13) bubble #ow meter.
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hydrocarbons and H O. Gas leakage through the sea lant, if present, could be detected by monitoring the nitrogen concentration in the e%uent from the tube side. The conversion of CH and selectivity of CO were de "ned, respectively, as follows:
Steam reforming: CH #H O & CO#3H P P !- & R "A P P exp(!E /R¹) 1! . !& & K P P !& & (4)
F !F !& , X " !& !& F !&
(1)
F !. S " !- F !F !& !&
(2)
Carbon dioxide reforming: CH #CO & 2CO#2H : P P !- & . R "A P P exp(!E /R¹) 1! !& !K P P !& !(5)
Blue}green 17.5% NiO/c-Al O catalyst was pur chased from Lanzhou Institute of Chemical Physics of Chinese Academy of Sciences. The catalyst used in the membrane reactor was pretreated in 1 : 1 mixture of H : He for 3 h at 7003C. The XRD analysis con"rmed that the NiO was completely reduced to the metal Ni by the H treatment. The Ni/c-Al O catalysts have the BET surface area of 97.62 m/g.
The equilibrium constants K , which are temperature G dependent, were calculated from thermodynamic properties (Shi, Wang, Yu & Chen, 1996) for i"2, 3. The values of activation energy E (i"1, 2, 3) were chosen from the G literature for each of the reactions (Arai, Yamada & Eguchi, 1986; Ross & Steel, 1973; Richardson & Paripatydar, 1990). The frequency factors A are obG tained by "xed-bed reactor experiment, which was conducted in our laboratory using reduced Ni/Al O (typically, CH : O "1.78 : 1). The kinetic constants are given in Table 1.
3. Kinetics and reactor mathematical model
3.2. Oxygen permeation rate in La0 6 Sr0 4 Co0 2 Fe0 8 O3
\d tubular membrane
3.1. Kinetic equations of methane oxidation to syngas over a Ni/Al2 /O3 catalyst Two main mechanisms of the partial oxidation of methane to syngas, depending on the catalyst, have been proposed. One is the mechanism of methane combustion followed by both steam and carbon dioxide reforming (Ashcroft et al., 1990), and the other is the catalytic CH pyrolysis followed by H desorption and carbon oxida tion (Hickman & Schmidt, 1993a,b). Dissanayake et al. (1991) examined the oxidation state and phase composition of the Ni/Al O catalyst as a function of axial position in the catalyst bed. Their results strongly supported the mechanism of methane combustion. Blank, Wittrig and Peterson (1990) also assumed this mechanism in their work by using a Ni/Al O catalyst. In contrast, Hickman and Schmidt (1993a,b) proposed the mechanism of the catalytic CH pyrolysis to explain the high selectivity of syngas formation using Pt and Rh as catalysts. They concluded that, under CH -rich atmo sphere and at high temperatures, H and CO were the primary products of direct oxidation of methane via a pyrolysis mechanism. The mechanism of methane combustion was assumed in the present simulation since a Ni/Al O catalyst was applied in this study, The kinetic equations of Blank et al. (1990) and Tsai (1996) were adapted as follows: Methane combustion: CH #2O PCO #2H O, R "A P P exp(!E /R¹). !&
(3)
The oxygen permeation through a tubular La Sr Co Fe O membrane under air/He con \B dition has been reported in our recent paper (Li et al., 2000a). Based on the experimental data of oxygen permeation, the oxygen permeation #ux can be approximately correlated to oxygen partial pressures and temperature by the following empirical equation: J "(1.37;10\¹#9.77;10\);¹ ;[(P ;10\)\ "\> "\2 !(P ;10\)\ "\> "\2!0.044], (6) where J is the #ux of oxygen permeation based on inner surface area of membrane tube, in cm (STP)/cm/s; ¹ is the temperature, in K; P is the oxygen partial pressure, in Pa; and the superscripts &&and'' represent the upstream and downstream membrane interfaces, respectively. Eq. (6) can be used in the temperature range of 1093}1173 K.
Table 1 Kinetic parameters for the partial oxidation of methane to syngas Reaction i
Activation energy E (J/mol) G
Frequency factor A G (mol/g -s-Pa)
1 2 3
166;10 29;10 23.7;10
1.10 4.19;10\ 2.42;10\
W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625
3.3. Model for dense membrane reactor The tube-and-shell reactor arrangement used in simulation is the same as the experimental arrangement shown in Fig. 3. The diluted methane is fed to the reactor tube side packed with the catalyst, and air is fed concurrently to the shell side of the reactor. Since the partial oxidation of methane to syngas is a mildly exothermic reaction and diluted feed is used, the assumption of an isothermal reactor may be reasonable. The mathematical model in this work is based on the following assumptions: (1) The membrane reactor is operated at steady state under isothermal and isobaric conditions. (2) Homogeneous gas phase oxidation reactions and carbon deposition are negligible. (3) The internal mass transfer resistance of the packed catalytic granules is neglected due to the small catalyst particles (40}60 mesh) and the external mass transfer resistance was also neglected because the Da (Damkoller) number was much less than the one in this work. (4) Ideal gas law is used to describe the gas behavior of single component and gas mixture. According to the principle of mass conservation, the design model equations for the perovskite membrane reactor are derived in each of the three regions as given below. In the tube side, i.e., the catalytic layer: 0(r(r dF L H #2pr J " !nr o l R "0, (7) H PP HG G dl G where the subscript j represents the reaction gas species, i.e., CH , O , CO, CO , H and H O. In the dense membrane: r (r(r r J " J " , (8) r - PP J "0, K
(9)
2621
where the gas species m includes CH , CO, CO , H , H O and N : In the shell side: r (r(r dQ I !2pr J " "0 (10) I PP dl in which the subscript k denotes O or N , respectively. The initial conditions, l"0, for the di!erent regions are F "F , (11) !& !& F "0, (12) H H$!& Q "Q , (13) Q "Q . (14) , , This set of di!erential equations is coupled together by their boundary conditions as follows: At the catalytic layer/dense membrane layer interface: r"r J "J " (15) - PP for other gas species m (including CH , CO, CO , H and H O): J "0, (16) K F H P, P" (17) H + F R G H where M"6, indicating CH , O , CO, CO , H and H O. At the dense membrane layer/shell side interface: r"r r J " J , (18) r PP J "0, (19) , Q I P. P" (20) I + Q Q G I in which M"2, indicating O and N .
Fig. 3. Schematic diagram of dense membrane reactor.
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W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625
These di!erential equations were reduced to a set of dimensionless equations by introducing dimensionless quantities, such as FM "F /F , QM "Q /Q and H H !& I I lM "l/¸. Then the model equations were solved using the numerical Runge}Kutta method.
4. Results and discussion In order to compare the simulating results with the experimental data, in most cases the membrane reaction experiments and simulation were carried out under the same conditions given in Table 2, other parameters will be given in speci"c cases. 4.1. Ewect of carbon space velocity In this case, the helium rate is "xed at 58 ml (STP)/min and the inlet methane rate is changed in the tube side. Fig. 4 shows the e!ect of carbon space velocity in the tube side on the reaction performance at 1158 K. The CH conversion decreases with increasing carbon space veloTable 2 Geometric properties of the dense membrane reactor and operation conditions Pressure for both tube and shell side Bulk density of catalytic bed Virtual length of reactor External diameter of membrane tube Inner diameter of membrane tube Air #ow rate Helium #ow rate
0.1013 MPa 0.726 g/cm 1.73 cm 0.8 cm 0.5 cm 100 ml(STP)/min 58 ml(STP)/min
Fig. 4. E!ect of carbon space velocity on the syngas production at 1158 K Experimental value: (䉫) CH conversion; (䉭) CO selectivity; Simulation value: (*) CH conversion; (- - - -) CO selectivity; (} ) } ) }) H /CO.
city, and the CO selectivity is larger than 90% in all cases. As the carbon space velocity increases, this in turn causes the CH /O ratio to increase. Since the permeation rate of oxygen is constant, the conversion decreases due to insu$cient oxygen for the reaction. Moreover, the carbon deposition on the surface of catalyst occurred due to higher CH /O ratio at the higher carbon space velocity. Dissanayake et al. (1991) found that in the "xed-bed reactor at temperatures *1023 K, the amount of surface carbon formed on the Ni/c-Al O catalyst was primarily in#uenced by the CH /O feed ratio. A high ratio of CH /O could cause deposition of large amounts of carbon that "lled the catalyst pores. In this work, EDS analysis con"rmed the deposition of carbon on the catalyst after use in the membrane reactor for the partial oxidation of methane at a higher carbon space velocity. The deposition of carbon reduced the activity of the catalyst. In addition, the decrease in methane conversion with increasing carbon space velocity is due to the fact that the syngas production reaction order with respect to methane was not "rst order which would give a methane concentration-independent conversion (Blank et al., 1990). All this suggests that the operation of the membrane reaction should be at a low carbon space velocity to obtain high methane conversion. As shown in Fig. 4, the simulating selectivity of CO shows good agreement with the experimental data, and the methane conversion of simulation is a little larger than that of experiment, especially at higher carbon space velocity. This may be attributed to the following reasons. First, the model in this work did not involve the e!ect of radial di!usion of gases on the performance of reaction, and assumed that at the same section of reactor the concentration of each gas is the same. This will result in the calculated rate of reactions being larger than the virtual value. Second, the model did not involve the e!ect of carbon deposition on the performance of reaction. This will lead to simulated methane conversion larger than the experimental data in higher carbon space velocity. Third, in fact, oxygen partial pressure at the wall of membrane tube in this model was mean value, which was lower than realistic oxygen partial pressure, the oxygen #ux calculated from Eq. (6) could be larger than the real oxygen #ux. The increase in oxygen #ux would lead to the increase in methane conversion. In addition, at higher carbon space velocity, the reaction would become radical, and the resistance of di!usion would in#uence the performance of reaction to some extent. Therefore, the di!erence between simulating conversion of methane and experimental data became larger at higher carbon space velocity. 4.2. Ewect of temperature Since the oxygen #ux of the La Sr Co Fe O membrane became usable above 1073 K (Li \B
W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625
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Table 3 E!ect of temperature on the syngas reaction Temp. (3C)
825 850 885
Carbon space velocity (min\)
8.24 8.24 8.24
X (%) !&
S (%) !-
H /CO
Exp.
Model
Exp.
Model
97.3 100 96.7
97.2 99.7 100
97 100 98
98.9 97.9 97.8
et al., 2000a) and the top temperature for the sealing was 1173 K, the experiment was carried out in the temperature range of 1073}1173 K. Table 3 gives the results of the experiment and the simulation at three di!erent temperatures. The experimental data cannot show the trends of changes in CH conversion and CO selectivity with temperatures due to the experimental error, however, the simulating results show that the CH conversion in creases and CO selectivity decreases with increase of temperatures. But, it can be seen that the reaction temperature in the range of 1098}1158 K has no signi"cant e!ect on both the CH conversion and CO selectivity when a lower carbon space velocity is applied. The CH conversion and the CO selectivity are, respectively, larger than 96 and 97% in all cases. These results are similar to the results of the "xed-bed reactor obtained by Dissanayake et al. (1991). This indicates that the temperature range of 1098}1158 K was suitable for the operation of the membrane reactor. From the above comparison between the experimental and simulating results, the model developed in this work was considered to be reasonable. Therefore, the model was used to simulate the e!ects of other parameters on the performance of the syngas reaction. 4.3. Ewect of air yow rate Fig. 5 shows the e!ect of shell side air #ow rate on the performance of the reaction under di!erent carbon space velocities at 1123 K. As shown, the in#uences of air #ow rate on both CH conversion and CO selectivity are insigni"cant when air #ow rate is greater than 100 ml (STP)/min in the shell side. In our previous experimental and modeling study (Li et al. 2000a,c), it was found that under an air/helium gradient the oxygen permeation #ux remained essentially unchanged with varying air #ow rate in the shell side when the air #ow rate was larger than 100 ml (STP)/min. In this work, the oxygen permeation took place under reaction conditions. The oxygen permeation rate under reaction conditions, which could be calculated from the #ow rate and the mole fractions of oxygen-containing product gases (such as O , CO, and H O in the tube side e%uent, was much larger than that under the air/helium gradient at the
Exp. 1.69 1.68 1.96
Model 2 2 2
Fig. 5. E!ect of air #ow rate on CH conversion and CO selectivity: (*) CH conversion; (- - - -) CO selectivity; (䉬䉫) He : CH 20 : 1; (䢇*) He : CH 15 : 1; (䉱䉭) He : CH 10 : 1.
same temperature. This is because the oxygen pressure near the membrane surface in the reducing side is extremely low when a reaction is taking place. However, it can be derived from Fig. 4 that for a certain carbon space velocity oxygen permeation #ux also remains essentially unchanged with varying air #ow rate when air #ow rate is no less than 100 ml (STP)/min. This indicates that the e!ect of air #ow rate on the oxygen permeation under reaction conditions is the same as that under air/helium gradient. This suggests that in the operation of the membrane reactor, it is unnecessary to further increase the air #ow rate if the air supplied is enough for providing stable oxygen permeation #ux. 4.4. Ewect of membrane tube length Simulations in this case were done according to the following parameters: a helium to methane ratio of 10/1, carbon space velocity of 6.614 min\, air #ow rate of 200 ml (STP)/min in the shell side, the membrane tube length of 0.07 m, temperature of 1123 K. Fig. 6 shows the simulation results of CH conversion and CO selectivity
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W. Jin et al. / Chemical Engineering Science 55 (2000) 2617}2625
Fig. 6. E!ect of membrane reactor length on syngas production: (*) CH conversion; (......) CO selectivity; (} ) } ) }) H /CO.
but the CO selectivity remained essentially unchanged with an increased carbon space velocity. In the range of 1098}1158 K, the CH conversion was larger than 96% with the CO selectivity larger than 97% at a low carbon space velocity. An isothermal model developed based on the experiment was also presented to simulate the performance of the membrane reactor under various operating conditions. Good agreement was obtained between the experimental data and simulating results. The simulation results indicate that the temperature range of 1098}1158 K is feasible for the operation of the membrane reactor. The air #ow rate ('100 ml (STP)/min) in shell side shows little in#uence on the performance of syngas reaction, and the CH conversion increases grad ually along the axial position of the membrane reactor and CO selectivity does not decrease (from 100%) until the methane conversion is close to 100%.
Acknowledgements along axial position of the reactor. As shown, the CH conversion increases gradually along the axial position and "nally reaches 100%, remains unchanged, and CO selectivity does not decrease (from 100%) until the methane conversion is close to 100%. The position at which methane conversion reaches 100% corresponds to around 71% of the membrane tube length. Beyond this position, CO selectivity decreases gradually due to the deep oxidation of products, which is caused by excess oxygen near the outlet of the reactor. Fig. 6 also re#ects the pro"les of methane and CO at di!erent axial positions. Compared with the co-feed reactor, oxygen is more uniformly distributed along the reactor axial direction in the membrane reactor. In "xed-bed reactor, the oxygen concentration, with the maximum at the entrance, decreases to zero along the reactor length and the reaction became the sharpest in the front part of the reactor. This could lead to a large adiabatic temperature rise, which may cause reactor runaway. However, in dense membrane reactors the oxygen is fed to the entire reactor length by permeation and the reaction became mild. The methane conversion and CO selectivity depend on the length of the dense membrane reactor for a certain carbon space velocity. In other words, for a "xed length of a membrane reactor, it is necessary to choose a suitable carbon space velocity to obtain high methane conversion and CO selectivity.
5. Conclusions The partial oxidation of methane to syngas was performed in a tubular La Sr Co Fe O mem \B brane reactor in the presence of a pre-reduced NiO/ c-Al O catalyst. The conversion of methane decreased
This work is supported by the National Advanced Materials Committee of China (NAMCC, No. 715-0060120), and the National Natural Science Foundation of China (NNSFC, No. 59789201).
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