Ind. Eng. Chem. Res. 1997, 36, 3369-3374
3369
Methane Steam Reforming Analysis in a Palladium-Based Catalytic Membrane Reactor Giuseppe Barbieri,†,‡ Vittorio Violante,§ Francesco P. Di Maio,| Alessandra Criscuoli,| and Enrico Drioli*,†,| CNR, Institute of Research on Membranes and Chemical Reactors Modelling, C.da Arcavacata, 87036 Rende (CS), Italy, CR ENEA Frascati, Frascati (RM), Italy, and Department of Chemical Engineering and Materials, University of Calabria, C.da Arcavacata, 87036 Rende (CS), Italy
The methane steam reforming in a catalytic membrane reactor has been studied. A previous theoretical study of this reaction has been carried out. In the model a global kinetic rate as function of three reactions over nickel catalyst as proposed by Xu and Froment (AIChE J. 1989, 35 (1), 88-96, 97-103) has been considered. It has been shown that the counterflow configuration has, at high temperature (500 °C), a marginal advantage on parallel flow and, also, that the space velocity cannot be considered a design variable for membrane reactors. A laboratory plant was realized utilizing membranes of Pd and Pd/Ag supported on Al2O3. The Pd membranes utilized have been prepared using the co-condensation technique and the electroless plating method. A comparison of the overall membrane performance has also been carried out. The experiments were aimed to study the effects of several parameters such as temperature, feed flow rate, and feed molar ratio on the methane conversion. The experimental results have been compared with the data predicted by the previously developed theoretical model. Introduction In the last few years the interest for innovative solutions, in industrial and chemical processes, has been focused on membrane reactors (Falconer et al., 1993; Zaman and Chakma, 1994; Saracco and Specchia, 1994). One of the most important advantages of these kinds of reactors is the possibility of overcoming the limitation imposed by the thermodynamic equilibrium. To remove hydrogen from the reaction volume, it is possible to use palladium and its alloys, permeable only by hydrogen. Techniques have been developed aimed at depositing a small amount of Pd over porous supports: chemical vapor deposition (Morooka et al., 1995), electroless plating (Shu et al., 1993), sputtering, and solvated metal atom deposition (Capannelli et al., 1993; Basile et al., 1996). In this way, it is possible to reduce the thickness of the membrane with the advantages of lower costs and of higher permeability. The durability of the Pd membranes, which are highly fragile after thermal excursions in the presence of hydrogen, is, however, a significant problem. In order to overcome this drawback, Pd/Ag (Gobina et al., 1995) and other alloys have been used. The hydrogen and syngas productions are two of the most important processes for chemical industry. About 50% of hydrogen demand is actually satisfied by means of the methane steam reforming (MSR) reaction, due also to increased natural gas production (mainly methane). Methane steam reforming involves five species in three reversible reactions.
CH4 + H2O ) CO + 3H2
∆H298 ) 206 kJ mol-1 (1)
CO + H2O ) CO2 + H2
∆H298 ) -41 kJ mol-1 (2)
* Author to whom correspondence should be addressed. † CNR. ‡ Phone: +39 984 492039. Fax: +39 984 402103. E-mail:
[email protected]. § CR ENEA Frascati. | University of Calabria. S0888-5885(97)00033-X CCC: $14.00
The reforming reaction (1) is thermodynamically favored by high temperature and low pressure. The second reaction, water-gas shift (WGS), does not depend on pressure and is favored by low temperature. Heat developed by reaction (2) is not sufficient for the thermal needs of the process, as can been seen from the overall reaction:
CH4 + 2H2O ) CO2 + 4H2
∆H298 ) 165 kJ mol-1 (3)
The overall reaction (3) is endothermic and limited by equilibrium; it needs high temperature to obtain reasonable industrial conversion. Traditionally the industrial converters are made of hundreds of parallel tubes contained in furnaces and operating at 850 °C to obtain a degree of methane conversion of around 80%. The catalyst used is Ni/alumina and is packed inner tubes. This reaction has been widely studied in the traditional systems. The catalyst that showed more efficiency for MSR reaction is the Ni/alumina, as reported by Xu and Froment (1989a,b). Their kinetic analysis, in a traditional reactor, has been made in the temperature range 573-823 K and for pressure varying from 300 to 1000 kPa. This kinetic model is based on the LangmuirHinshelwood mechanism and involves 13 steps that lead to the three global reactions reported above. Recently various papers on MSR in a catalytic membrane reactor (CMR) appeared in the literature. Tsotsis et al. (1992) have studied MSR in a porous catalytic membrane reactor in the temperature range 200-600 °C. Hughes (Adris et al., 1991) studied the behavior of MSR reaction in a fluidized bed. Uemiya et al. (1990) utilized a Pd-membrane reactor to study the same reaction. Laegsgaard et al. (1995) have analyzed the carbon deposition over catalyst, in a Pd-Ag membrane reactor, for the MSR reaction; in membrane reactors a higher steam/methane ratio than those in traditional ones has to be used to avoid the carbon formation due hydrogen removal from the reaction volume. Shu et al. (1993) showed that using Pd and Pd-Ag membranes for MSR reaction, it is possible to enhance the conver© 1997 American Chemical Society
3370 Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997
sion of methane by removal of hydrogen, also at mild operating conditions. Bredesen and Sogge (1996), studying the production of hydrogen and syngas by means of natural gas transformations, showed that membrane reactors give promising technical and economical performance. The experimental results, obtained using Pd membrane and reported in the literature, refer to infinite selectivity of these membranes for hydrogen. In this work we tried to make clear the discrepancy which exists when comparing the theoretical results from a model based on ideal membrane and experimental results which can be reached by the Pd membrane characterized by some defects. The fact these defects grow with time can also have effects on the agreement between the experimental results and the model predictions. Experimental Apparatus An experimental plant to study methane steam reforming reaction in a catalytic membrane reactor was realized. It consists of a CMR packed with catalyst (about 7 g), analytical units, and control systems for flows and temperature. The CMR consists of two concentric pipes: the inner one is the membrane, whereas the outer is a SS shell. The sealing between these two parts is realized by means of graphite O-rings. The reactants are supplied to the inner tube where catalyst is packed. The commercial catalyst 30% Ni over alumina has been furnished by Hengelhard de Meern BV. In the annular space, the sweep gas flows in parallel flow mode with the reactant stream. Both retentate and permeate streams are analyzed by means of a GC. Mass flow controllers have been used to control the feed and sweep gas flow rates. The heating system consists of a temperature controller (PID) connected to a heating tape that winds up the SS shell. Two thermocouples were located at the feed side of the reactor at the beginning of the catalyst bed and at the exit of the reaction volume. Temperature oscillations (up to 10 °C) have been observed in the system. An average value of the temperature has been considered in the elaboration of the data.
well applied to have uniform metal deposition on relatively complex geometry. The ceramic support is activated, e.g., by stannous and palladium chloride (SnCl2 and PdCl2), to have the metallic layer formation on the external surface of the ceramic support. The PdAg alloy is obtained by diffusion at high temperature. The amount of each deposited metal is consistent with the required alloys composition. An intermediate ceramic layer (having 0.1 µm of average pore size) is required, between the metallic film and support, if the support average pore size is larger than 0.1 µm. Solvated Metal Atom Deposition (SMAD). Recently, a new technique to deposit a very thin film on a commercial ceramic membrane has been suggested by Capannelli et al. (1993): Pt vapor and mesitylene have been co-condensed at 77 K in a glass reactor. The solid matrix produced has been heated up to 233 K, yielding a solution stable at low temperature which has been filled in a commercial ceramic tube made of four layers (three R-alumina layers of different of pore sizes, as macroporous support, and a thin γ-alumina layer). The same technique has been used to deposit a thin layer of Pd on the same type of ceramic tube; in this case Pd vapor, mesitylene, and hexene-1 have been co-condensed at 77 K in a glass reactor. This technique allows one to deposit the metallic film (about 0.1 µm) on the inner surface of the ceramic support. Mathematical Model For membrane reactors, besides species generation and axial fluxes (Ni), permeation fluxes through the membrane have been considered in the development of a mathematical model for steady-state operations (Barbieri and Di Maio, 1997). The basic assumptions are plug flow on reaction and permeation side, isothermal and isobaric conditions; MSR and WGS kinetics are described by means of the global reactions (1-3) based on reaction rates (rj) proposed by Xu and Froment. The set of equations written in terms of fluxes consists of the following ordinary differential equations:
Reaction side 1000Wcat
dNi )
Membrane Preparation Membranes used in this work are based on a very thin dense film of Pd or Pd-Ag alloys deposited on a tubular support of alumina, with two different techniques: electroless plating (EP) (Shu et al., 1993) and solvated metal atom deposition (SMAD) (Capannelli et al., 1993; Basile et al., 1996). The first technique allows the deposition of the metallic film on the external surface of the support. With the second one the deposition can be carried out also on the inner surface of the tubular support. Electroless Plating (EP). The thin metallic film has been obtained with a modified electroless technique. The film is a Pd-Ag (79/21%) alloy, 10 µm thick, with a low defects concentration and homogeneous surface, ensuring a significant separation factor between hydrogen and the other gaseous components before the reaction tests. Such a composite structure shows a remarkable capacity to operate at high temperature. The film has been obtained by means of the electroless plating technique because of the applicability of such a technique to obtain layers of different metals having different thicknesses. The electroless deposition can be
dz
60Vr
3
νi, jrj - H Pd ∑ i j)1
πφ Ar
Permeation side dNi πφ ) (H Pd i dz Ar where Wcat is the catalyst weight, Vr is the reaction volume, Ar is the cross section of the reaction side, φ is the membrane diameter, νi,j is the reaction coefficient, and z is the reactor axial position; the + sign is used for the parallel-flow configuration and the - sign must be used for the counterflow configuration. Only the hydrogen balance equations (i ) H2) (reaction and permeation side) have the permeation term Pd HPd is equal to i . For the other species the term Hi zero. The traditional reactor model is obtained from the present model by assuming no permeation even for hydrogen. The set of differential equations has been solved numerically by a fourth-order variable step RungeKutta method; for counterflow simulation the shooting method has been used.
Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997 3371 Table 1. Selectivity of the Products with Respect to the Methane for EP and SMAD Membranes, Measured during Reaction Tests SMAD EP
Figure 1. Degree of methane conversion versus space velocity/ reference space velocity (S/S0). S0 ) 11 g of catalyst/162 SCCM of feed, T ) 500 °C, H2O/CH4 ) 3, Pr ) 103 kPa, Pp ) 101 kPa, Qp ) 40 SCCM. Membrane thickness ) 20 µm. (a) Degree of conversion of the membrane reactor, when the amount of catalyst is reduced from the initial value of 11 g at a fixed feed flow rate of 162 SCCM. (b) Obtained for membrane reactor by increasing the feed flow rate from the initial value of 162 SCCM at a fixed catalyst amount of 11 g. (c) Calculated results obtained for a traditional reactor with a change (one at time) in the amount of catalyst and flow rate. The solid square represents the equilibrium value.
The driving force for hydrogen permeation is the difference of the square root of the partial hydrogen pressures between the reaction and permeate sides. Simulation results showed that the counterflow configuration is not always better than the parallel one. Depending on operating conditions (e.g., temperature, pressure, feed and sweep flow rate, etc.), the conversion degree for parallel flow is higher than that obtained by means of the counterflow. The model shows that the space velocity cannot be considered a design variable due to the complex interaction between flow rate and permeation through the membrane. In traditional reactor models, if space velocity (S) is changed by increasing (or reducing) the reactant flow rate or decreasing (or increasing) catalyst weight, the same results are obtained. The model (Barbieri and Di Maio, 1997) shows (Figure 1) that a different conversion can be achieved with the same (S/S0) value for the membrane reactor. It can be observed that methane conversion is always higher than that of the traditional reactor and that the difference in the values depends on the way the S/S0 ratio is changed. If the variation is imposed by increasing the reactant flow rate, curve b is obtained. Instead, if the variation is made by reducing catalyst weight, the results obtained (curve a) show a significant difference with respect to those of the traditional reactor. Results and Discussion In all experiments the analysis of the relative concentrations of the exit streams was made by measurement with a GC system. The conversion values have been based on both shell and tube concentrations. The experimental results clearly put in evidence the nonideal behavior of the Pd membranes utilized; its selectivity, moreover, is reduced with the progress of experimental runs. During the reaction tests, together with the H2, a certain percentage of the reagents and products were also permeated. This kind of phenomenon will be considered in the analysis of the experi-
H2/CH4
CO/CH4
CO2/CH4
2.61 0.87
3.07 1.85
1.79 0.58
Figure 2. Degree of methane conversion versus feed flow rate. Experimental data for SMAD (solid circles) and EP (solid diamonds) membranes; fitting of experimental data (a) and model results (b). The solid square represents the equilibrium value. T ) 536 °C, H2O/CH4 ) 2.5, Pr ) Pp ) 101 kPa, and Qp ) 53.2 SCCM.
ments and in the comparison with the analytical prediction. In Table 1 experimental selectivities are reported. These selectivities have been calculated on the basis of the GC analysis of the relative concentrations of the permeate and retentate streams. The presence of all species in the permeate stream indicates the nonideality of the membrane due to pinholes and cracks on the metal layers. The two membranes show different values of selectivities; in particular, the SMAD membrane presents higher selectivities with respect to the EP one. The difference in selectivity values can be explained assuming that two transport mechanisms, in addition to the hydrogen diffusion in palladium bulk, occur: Knudsen and surface diffusion. Being that the oxygenated species preferentially adsorbed on the membrane surface (Xu and Froment, 1989a), both the SMAD and the EP membrane present a H2/CH4 selectivity lower than the CO/CH4 and CO2/ CH4 selectivities. In addition, in the EP membrane the Pd-Ag layer (10 µm) is located on the external surface of the support and is thus far from the reaction region; here the gaseous phase containing more reaction products than the reaction zone could react (Capannelli et al., 1996). So, CO and CO2 could react with H2 in the Pd layer to give CH4 and H2O. These reactions do not occur for the SMAD membrane, since the very thin Pd layer is deposited on the internal surface of the support. In this way it is possible to justify the lower selectivities for the EP membrane with respect to the SMAD one. Figure 2 reports the degree of methane conversion against the reactant feed flow rate for a H2O/CH4 molar feed ratio equal to 2.5. Solid circles in Figure 2 represent the experimental data obtained with the SMAD membrane, while the solid diamonds represent the experimental data for the EP membrane.
3372 Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997
Figure 3. Degree of methane conversion versus feed flow rate for the EP membrane at two different feed molar ratios. Experimental data: (a; solid diamonds) m ) 2.5; (b; solid circles) m ) 1.18. Model results: (c) m ) 2.5; (d) m ) 1.18. T ) 536 °C, Pr ) Pp ) 101 kPa, and Qp ) 53.2 SCCM. The solid square represents the equilibrium value at m ) 1.18, and the open square represents the equilibrium value at m ) 2.5.
Figure 4. Degree of methane conversion versus feed molar ratio; (solid circle) obtained with the SMAD membrane; (solid diamonds) relative to the EP membrane; (dashed line) model results; (solid line) equilibrium values. T ) 536 °C, Pr ) Pp ) 101 kPa, Qr ) 88.5 SCCM, and Qp ) 53.2 SCCM.
The degree of conversion decreases, increasing the feed flow rate as shown by fitting (curve a) because of the reduced time factor. This is in agreement with the model (curve b), that predicts another small decrease due to the reduction of residence time. It can be noted that the two membranes show the same trend and that the methane conversion values are very close to each other. However, the EP membrane presents lower conversion values with respect to the SMAD one due to its lower selectivity. In Figure 3 the experimental degree of conversion against the feed flow rate for two feed molar ratios is reported. The membrane used is the EP one. The data obtained with a feed molar ratio equal to 2.5 (curve a) showed the same trend as those obtained at H2O/CH4 ) 1.18 (curve b). The conversion values obtained by increasing the feed molar ratio are, of course, higher. Figure 4 shows the degree of methane conversion versus the feed molar ratio. Experimental data, performed with the EP membrane, show the same trend
Figure 5. Methane conversion versus temperature: (a) experimental data for the SMAD membrane; (b) model results; (c) equilibrium. Pr ) Pp ) 101 kPa, Qr ) 98 SCCM, Qp ) 53.2 SCCM, and m ) 4.
Figure 6. Experimental degree of methane conversion versus Qr for the SMAD membrane. T ) 536 °C, Qp ) 53.2 SCCM, m ) 4, and Pr ) Pp ) 101 kPa. The solid square represents the equilibrium value.
as predicted by the model. On the same figure a point referring to the SMAD membrane is also reported. The effect of temperature on the degree of methane conversion is shown in Figure 5. The conversion increases, as predicted by the model, with the temperature. The agreement between experiments and the model is good in the low-temperature range, but at higher temperature values a significant difference is observed. Figure 6 reports the degree of methane conversion against the reactant feed flow rate for a H2O/CH4 molar feed ratio equal to 4. The membrane used is the SMAD one. The results reported in Figures 2-5 have been compared to the theoretical predictions derived by the model (Barbieri and Di Maio, 1967). The observed discrepancy between the theoretical and the experimental data can be explained as follows. In the simulation program only hydrogen permeates through the membrane, whereas the two membranes used were found to be selectively permeable to all species. Particularly, the
Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997 3373
Acknowledgment The authors are grateful to Prof. G. Vitulli of the University of Pisa, Pisa, Italy, for the helpful discussion and preparation of the SMAD membrane and to Dr. Carlo Cavenaghi of the Hengelhard de Meern BV, Milano, Italy, for the catalyst supply. Nomenclature
Figure 7. Model results for several sweep gas flow rates: (a) Qp ) 500 SCCM; (b) Qp ) 200 SCCM; (c) Qp ) 53.2 SCCM. The experimental results (circles) for a SMAD membrane are also reported (see Figure 2). T ) 536 °C, Pr ) Pp ) 101 kPa, and m ) 2.5. The solid square represents the equilibrium value.
SMAD membranes showed a higher selectivity for products than for reactants (see Table 1); thus, the shift of equilibrium in the experimental cases has been higher than that in the theoretical one, with subsequent increase in conversion. The increase in conversion has to be also related, for both membranes, to a “dilution” effect due to the presence of the sweep gas in the reaction side. In Figure 7 the experimental degree of methane conversion as a function of Qr is compared with the model predictions for different values of sweep gas flow rates. The same values obtained in the experimental runs might also be achieved theoretically for a sweep gas flow rate on the order of 500 SCCM. The discrepancy observed might be the results of a combination of more effects: the higher presence of products in the permeate stream, a certain increasing of sweep flow rate (due to the permeation of all species), and the presence of sweep gas in the reaction side. Conclusions The data presented in this work confirm the interest of catalytic membrane reactors in methane steam reforming reaction in terms of a possible increase of conversion and the possibility of operating at mild operating conditions. However, the fact that most of the available Pd membranes will not behave as ideal ones has to be considered. The existing models, in general, assumed infinite selectivities for hydrogen. The experimental results obtained in our work have shown that also with noninfinite selectivities of membrane, if appropriate operating conditions are introduced, interesting benefits can be obtained. For example, an increase of the conversion value on the order of 20%, with respect to the equilibrium one, has been observed for H2O/CH4 equal to 2. However, the possibility of operating at lower temperatures than those used in traditional systems appears to be promising. The experimental results show also that the two different techniques used for the preparation do not change the overall performance of the system, inducing a good reproducibility in the obtained data. Nevertheless, the SMAD membrane gave certain higher conversion values than the EP one.
A ) area [cm2] HPd ) molar flux through a Pd membrane [mmol cm-2 min-1] m ) H2O/CH4 molar feed ratio N ) molar flux [mmol cm-2 min-1] P ) pressure [Pa] Q ) flow rate [SCCM ) cm3 (STP) min-1] r ) reaction rate [mol (g of catalyst)-1 h-1] S ) space velocity [SCCM (g of catalyst)-1] T ) temperature [°C] V ) volume [cm3] Wcat ) catalyst weight [g] z ) reactor axial position [cm] ∆H298 ) reaction heat [kJ mol-1] φ ) Pd-membrane outer diameter [cm] νi, j ) reaction coefficient Superscripts and Subscripts 0 ) initial conditions or reference value p ) permeation side r ) reaction side i ) species i j ) reaction j
Literature Cited Adris, A. M.; Elnashaie, S. S. E. H.; Hughes, R. A fluidized bed membrane reactor for the steam reforming of methane. Can. J. Chem. Eng. 1991, 69, 1061-1070. Barbieri, G.; Di Maio, F. P. Modelling of methane steam reforming process in catalytic Pd-membrane reactor. Ind. Eng. Chem. Res. 1997, in press. Basile, A.; Drioli, E.; Santella, F.; Violante, V.; Capannelli, G.; Vitulli, G. A study on catalytic membrane reactors for water gas shift reaction. Gas Sep. Purif. 1996, 10 (1), 53-61. Bredesen, R.; Sogge, J. A technical and economic assessment of membrane reactors for hydrogen and syngas production. Seminar on Ecological Applications of Innovative Membrane Technology in the Chemical Industry, Cetraro (CS), Italy, May 1996. Capannelli, G.; Bottino, A.; Gao, G.; Grosso, A.; Servida, A.; Vitulli, G.; Mastantuono, A.; Lazzaroni, R.; Salvadori, P. Porous Pt/γAl2O3 catalytic membrane reactors prepared using mesitylene solvated Pt atoms. Catal. Lett. 1993, 20, 287-297. Capannelli, G.; Carosini, E.; Cavani, F.; Monticelli, O.; Trifiro`, F. Comparison of the catalytic performance of V2O5/γ-Al2O3 in the oxidehydrogenation of propane to propylene in different reactor configurations: i) packed-bed reactor, ii) monolith-like reactor and iii) catalytic membrane reactor. Chem. Eng. Sci. 1996, 51 (10), 1817-1826. Falconer, J. L.; Noble, R. D.; Sperry, D. P. Catalytic membrane reactors. The Handbook of Membrane Separations; Marcel Dekker: New York, 1993. Gobina, E. N.; Oklany, J. S.; Hughes, R. Elimination of ammonia from coal gasification streams by using a catalytic membrane reactor. Ind. Eng. Chem. Res. 1995, 34, 3777-3783. Laegsgaard Jorgensen, S.; Hojlund Nielsen, P. E.; Lehrmann, P. Steam reforming of methane in a membrane reactor. Catal. Today 1995, 25 (3-4), 303-307. Morooka, S.; Yan, S.; Yokoyama, S.; Kusakabe, K. Palladium membrane formed in macropores of support tube by chemical vapor deposition with crossflow through a porous wall. Sep. Sci. Technol. 1995, 30 (14), 2877-2889. Saracco, G.; Specchia, V. Catalytic inorganic-membrane reactors: present experience and future opportunities. Catal. Rev.-Sci. Eng. 1994, 36 (2), 305-384.
3374 Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997 Scholz, W. H. Processes for industrial of hydrogen and associated environmental effects. Gas Sep. Purif. 1993, 7 (3), 131-139. Shu, J.; Grandjean, B. P. A.; Ghali, E.; Kaliaguine, S. Simultaneous deposition of Pd and Ag on porous stainless steel by electroless plating. J. Membr. Sci. 1993, 77, 181-195. Tsotsis, T. T.; Champagnie, A. M.; Vosileiadis, S. P.; Zoika, Z. D.; Minet, R. G. Packed bed catalytic membrane reactors. Chem. Eng. Sci. 1992, 47, 9-11. Uemiya, S.; Sato, N.; Ando, H.; Matsuda, T.; Kikuchi, E. Promotion of steam reforming by use of palladium membrane. Sekiyu Gakkaishi 1990, 33 (6), 418-421. Xu, J.; Froment, G. F. Methane steam reforming, methanation and water gas shift: I. intrinsic kinetics. AIChE J. 1989a, 35 (1), 88-96.
Xu, J.; Froment, G. F. Methane steam reforming: II. Diffusional limitations and reactor simulation. AIChE J. 1989b, 35 (1), 97103. Zaman, J.; Chakma, A. Inorganic membrane reactors: review. J. Membr. Sci. 1994, 92, 1-28.
Received for review January 10, 1997 Revised manuscript received May 19, 1997 Accepted May 20, 1997X IE970033W Abstract published in Advance ACS Abstracts, July 1, 1997. X
