Ind. Eng. Chem. Res. 1997, 36, 4549-4556
4549
Characteristics of Fluidized-Bed Membrane Reactors: Scale-up and Practical Issues Alaa-Eldin M. Adris* SABIC R&D, P.O. Box 42503, Riyadh 11551, Saudi Arabia
John R. Grace Department of Chemical and Bio-Resources Engineering, University of British Columbia, Vancouver, Canada V6T 1Z4
Fluidized-bed membrane reactors (FBMR) are examined from a scale-up and practical point of view. Mathematical modeling is utilized to explore potential configurations for commercial FBMR steam methane reforming (SMR) as well as to quantify the effects of key design parameters such as membrane capacity, distribution of membrane surface between the dense bed and dilute phase, permeate side pressure, and sweep gas flow. Key factors affecting the performance of a commercial FBMR are analyzed and qualitatively compared with corresponding factors in packedbed membrane reactors. Issues which pose challenges to the commercial viability of this technology are identified. These include maintenance of bed mobility in the presence of gas withdrawal, providing sufficient membrane capacity, wear, and mechanical forces on vertical surfaces. Introduction The use of palladium-based membranes was initiated by a discovery by Graham (1866) that metallic palladium absorbs an unusually large amount of hydrogen. Inorganic membrane reactors are still in their infancy, with the possible exception of some dense palladiumbased membrane reactors used in the industrial production of chemicals and pharmaceuticals in Russia. However, there has been increasing research exploring the use of inorganic membranes as catalytic reactors, particularly in Japan and Russia (Hsieh, 1989). Extensive reviews of membrane reactors have been given by Tsotsis et al. (1993) and Saracco and Specchia (1994). The productivity of membrane reactors has been constrained by the limited permeability of the membranes. Commercially available permeable but nonporous membranes are either a thick film or thickwalled tubes. Since the permeability is inversely proportional to film thickness, a thick film membrane acts as a poor permseparator. Thus, developing permselective thin solid films, without compromising their structural integrity, is critical to future applications of membrane reactor technology and gas separation. The importance of the membrane film thickness can be demonstrated by considering the permeation rate of hydrogen gas through a palladium membrane (Itoh, 1987). For isothermal, isobaric, and plug flow conditions, the hydrogen permeation rate is approximated by a half power pressure law (Bohmholdt and Wicke, 1967):
QH ) kH[(PhH)1/2 - (PlH)1/2]
(1)
where QH is the hydrogen permeation rate in mol/h, PhH and PlH are the hydrogen partial pressures on the hig- and low-pressure sides, respectively, and kH is the permeation rate constant defined as
kH ) (Am/d)DFC0
(2)
where Am is the membrane surface area, DF is the * Corresponding author. S0888-5885(97)00108-5 CCC: $14.00
hydrogen diffusivity within the membrane material, C0 is the hydrogen solubility in the metal, and d is the membrane wall thickness. The permeation flux is computed to be 44.4 mol/h‚m2 for a membrane thickness of 0.2 mm at an opening temperature of 600 °C and hydrogen partial pressures of 0.25 and 0.1 MPa on the high- and low-pressure sides, respectively. For a membrane film thickness of 10 µm, the permeation flux increases to 0.89 kmol/h‚m2, representing a 20-fold increase in the product removal rate and consequently in the productivity of a reactor, provided that all other operating conditions remain unchanged. Such a reduction in the film thickness is achievable by applying a controlled film deposition technique (Chopra, 1969; Lee, 1980; Hsieh, 1989). The results presented in this work are based on a membrane film thickness of 10 µm, which is a conservative estimate of an achievable thickness in view of the current efforts of making membranes by these film deposition techniques. For instance, Yueng and Varma (1995) reported a novel deposition method which employs a combination of electroless plating and osmosis to prepare a thin film, as thin as 8 µm, of palladium on a ceramic substrate. Also Buxbaum and Kinney (1996) reported a membrane tube made of a tubular palladiumcoated tantalum and niobium membrane where the membrane layer is 5 microns thick. Despite some clear advantages which fluidized-bed operation offer to catalytic membrane reactors, reported attempts to combine fluidization and membrane separation in a reactor system have been quite limited. Earlier work by the authors and their co-workers (Adris, 1994; Adris et al., 1991a,b, 1994a,b, 1997) has focused on proving the concept and usefulness of this combination as well as on investigating the hydrodynamic aspects of this special type of fluidized bed which involves lowclearance internals. Mleczko and co-workers (Mleczko et al., 1996; Ostrowski et al., 1977) have studied the FBMR for a new reaction system, catalytic partial oxidation of methane (CPOM) to synthesis gas. This paper is divided into two parts: The first provides commercial reactor projections and quantifies the importance of major design issues, such as sweep © 1997 American Chemical Society
4550 Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997 Table 1. Projections of a Commercial FBMR for SMR (Case 2) in Comparison with a Fluidized-Bed Reactor of the Same Dimensions without Permeation Tubes (Case 1) Reaction Conditions and Design Parameters methane feed ) 175 000 kmol/h catalyst solid density ) 3350 kg/m3 carbon monoxide feed ) 0.0 kmol/h bed diameter equivalent to free area ) 2.0 m carbon dioxide feed ) 0.0 kmol/h reactor bed height ) 12.0 m steam feed ) 612 500 kmol/h freeboard height ) 4.0 m hydrogen feed ) 250 kmol/h sweep gas pressure ) 0.2 MPa reactor bed temperature ) 1023 K sweep gas molar flow ) 175 000 kmol/h reactor pressure ) 3.0 MPa case 1 permeation capacity in the dense bed, km permeation capacity in the dilute phase, km total permeation capacity, km methane conversion at the bed surface methane conversion at the reactor exit steam conversion at the bed surface steam conversion at the reactor exit total hydrogen produced, kmol/h pure hydrogen separated, kmol/h
0.498 0.414 0.230 0.212 2.75 × 105
case 2 26 700 17 800 44 500 0.797 0.901 0.424 0.494 6.18 × 105 5.77 × 105
gas flow and pressure as well as the issue of the utility of membranes in the freeboard region. The second part deals with mechanical and other design issues which are bound to be important in developing commercial FBMR systems. Some of this treatment is of necessity qualitative and/or in comparison with packed beds. Commercial-Scale Reactor Projections Assumptions and Membrane Capacity Considerations. The reactor model which was developed and successfully validated against pilot-scale-plant data in an earlier study (Adris et al., 1997) is first utilized to explore some key characteristics of FBMR systems. Projected characteristics of a typical commercial reforming unit are given in Table 1 based on the following key assumptions: 1. Membrane separation is provided by means of palladium-coated sintered-metal membrane tubes bundled together in a modular design as disclosed by Adris et al. (1994b). Each module comprises 100 tubes of 4.8 mm outside diameter and 2 m length. This provides an equivalent permeation capacity (outside membrane surface divided by permeation layer thickness) of 3 km/tube for a palladium layer thickness of 10 µm. One of these modules therefore provides 300 km of permeation capacity. On this basis, the dense bed and dilute phase can together accommodate up to 1.2 × 105 km of membrane capacity, occupying up to 0.35-0.5 m2 of the cross section of the reactor described in Table 1, depending on the vertical stacking and configuration of the modules’ arrangement. 2. The solids concentration profile in the freeboard zone is assumed to follow an exponential decay relationship as suggested by Wen and Chen (1982). A typical solids concentration profile in the freeboard zone appears in Figure 1. 3. The temperature profile in the freeboard zone is approximated by the pattern observed in the pilot-plant testing measurements (Adris et al., 1994a). A typical freeboard temperature profile is also shown in Figure 1. This temperature profile was assumed to apply in all cases considered below. 4. Isothermal conditions are assumed in the dense catalyst bed. The heat supply issue has been addressed in an earlier study (Adris et al., 1991a), where the possibility of heat pipes immersed in the catalyst bed was raised. The reactor is sized in terms of its equivalent diameter (diameter of a circle with area equivalent to the
Figure 1. Solids concentration and temperature profiles in the freeboard region for reaction conditions and design parameters as in Table 1.
free cross-sectional area). It is to be noted here that the performance of the reactor in Table 1 is typical of the expected results for low-temperature reforming where the equilibrium conversions corresponding to the reaction conditions used (T ) 750 °C, P ) 3.0 MPa, and S/C ) 3.5) are 0.568 and 0.262 for methane and steam, respectively. It is evident from the simulation results presented in Figure 2 that membrane capacities as low as 26 700 km, corresponding to about 90 modules occupying a cross-sectional area of 0.12 m2 when 6 modules are vertically stacked, are capable of doubling the methane conversion and the hydrogen production. The membrane capacity to be installed in an FBMR depends on the balance between the capital investment made in the membrane installation and the return on investment achievable by producing more high-purity hydrogen and/or by saving energy. Figure 2 shows the effect of varying permeation capacity on methane conversion both at the bed surface and at the reactor exit. Figure 3 shows the relation between the membrane capacity and both pure as well as total hydrogen produced. Comparison between the dependence of methane conversion on permeation capacity and that of pure hydrogen indicates an important factor regarding membrane installation. The improvement on methane conversion as the permeation capacity increases almost reaches a plateau when Cep exceeds about 18 000 km. Beyond this value the methane conversion increase becomes relatively small, while such a plateau is not reached for the hydrogen flow vs permeation capacity
Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997 4551
Figure 2. Methane and steam conversions at bed surface and reactor exit as functions of permeation capacity for reaction conditions and design parameters given in Table 1.
Figure 3. Permeate hydrogen and total hydrogen as functions of permeation capacity for reaction conditions and design parameters given in Table 1.
curve until Cep reaches about 30 000 km. To illustrate this point further, increasing the membrane capacity from 17 800 to 71 200 km (by a factor of 4) leads to only a small increase in methane conversion (12%) but to an appreciable increase in the amount of pure hydrogen separated (28%). Economics may possibly justify this increase in the membrane capacity for a gain of 28% in the production
of pure hydrogen, especially if hydrogen is the target product and its purity is a critical product specification. However, it is unlikely that a 12% methane conversion increase or 28% extra pure hydrogen would be enough economic incentive for such a gross addition of permeation capacity when the reforming plant is targeting the production of syngas and the separated hydrogen is to be mixed with the main reactor effluent stream. Membranes in the Freeboard. When reversible reactions, e.g., dehydrogenation or steam methane reforming, are carried out in fluidized-bed reactors, product gases leaving the catalyst bed usually contain entrained catalyst particles which cause further chemical reaction to take place. This reaction may or may not be desirable. In the SMR process, since the freeboard zone is normally cooler than the catalyst bed, the reaction over the surface of the entrained catalyst bed in the freeboard zone causes some of the products to revert into reactants, therefore losing some of the conversion achieved in the main catalyst bed. For steam methane reforming, Geotsch and Say (1989) patented a sudden quenching method to cool product gases in the freeboard to temperatures below those which favor re-formation of reactants. Adris et al. (1994b) employed the thermodynamic equilibrium shift caused by hydrogen withdrawal through membrane tubes to counter the problem of reaction reversal in the freeboard zone. The FBMR design has membrane tubes extending through the freeboard as well as through the main catalyst bed. The removal of hydrogen from the gas mixture within the freeboard greatly reduces undesirable reverse reactions. The results presented in Figures 2 and 3 are for cases where the membrane modules are available only for the dense catalyst bed and not the freeboard region. The distribution of the membrane capacity between the dense and dilute phases is an important design parameter. It is addressed here by a series of simulation runs where different scenarios for the membrane distribution have been considered. Simulation results are given in Tables 2-4 for reactors with total permeation capacities of 17 800, 35 600, and 53 400 km, respectively. The results presented in these three tables indicate the impact of membrane distribution between the dense bed and dilute phase on both the overall reaction conversion and the amount of hydrogen separated. Putting a portion of the permeable surfaces in the freeboard in Table 2 does not offer any improvement in the reactor performance, indicating that, at this modest membrane capacity, all of the permeable surface should be located in the dense bed. Results in Table 3 show that a better reactor performance is achievable at a bed to freeboard capacity ratio of 3:1 (case 2) in terms of both the methane conversion and the total hydrogen production. However, a ratio of 1:1 (case 3) gives a higher pure hydrogen productivity, even though it gives a lower methane conversion than case 2. Table 4 indicates the distribution effect for a high total permeation capacity (Cep ) 53 400 km). The reactor performance is seem to improve as the portion of the membrane surface in the dilute phase is increased until the ratio of the surface in each region reaches 1:1, after which the performance decreases. The following can be concluded from these three tables: i. The distribution of the membrane capacity between the dense bed and dilute phase of an FBMR is an important design parameter which needs to be opti-
4552 Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997 Table 2. Effect of Membrane Distribution between the Dense Bed and Dilute Phase for a Total Membrane Capacity of 17 800 kma case 1 permeation capacity in the dense bed, km permeation capacity in the dilute phase, km methane conversion at the bed surface methane conversion at the reactor exit steam conversion at the bed surface steam conversion at the reactor exit total hydrogen produced, kmol/h pure hydrogen separated, kmol/h a
17 800 0.777 0.767 0.397 0.408 5.18 × 105 3.93 × 105
case 2
case 3
case 4
13 350 4450 0.740 0.732 0.368 0.391 4.69 × 105 3.89 × 105
8900 8900 0.678 0.675 0.328 0.361 4.58 × 105 3.63 × 105
4450 13 350 0.595 0.599 0.281 0.320 4.06 × 105 3.21 × 105
Reaction conditions and design parameters are as given in Table 1.
Table 3. Effect of Membrane Distribution between the Dense Bed and Dilute Phase for a Total Membrane Capacity of 35 600 kma case 1 permeation capacity in the dense bed, km permeation capacity in the dilute phase, km methane conversion at the bed surface methane conversion at the reactor exit steam conversion at the bed surface steam conversion at the reactor exit total hydrogen produced, kmol/h pure hydrogen separated, kmol/h a
35 600 0.804 0.843 0.431 0.451 5.72 × 105 4.72 × 105
case 2
case 3
case 4
26 700 8900 0.797 0.870 0.424 0.473 5.95 × 105 5.35 × 105
17 800 17 800 0.777 0.850 0.397 0.464 5.82 × 105 5.36 × 105
8900 26 700 0.678 0.749 0.328 0.405 5.10 × 105 4.67 × 105
Reaction conditions and design parameters are as given in Table 1.
Table 4. Effect of Membrane Distribution between the Dense Bed and Dilute Phase for a Total Membrane Capacity of 53 400 kma case 1 permeation capacity in the dense bed, km permeation capacity in the dilute phase, km methane conversion at the bed surface methane conversion at the reactor exit steam conversion at the bed surface steam conversion at the reactor exit total hydrogen produced, kmol/h pure hydrogen separated, kmol/h a
53 400 0.815 0.859 0.438 0.461 5.84 × 105 4.89 × 105
case 2
case 3
case 4
case 5
44 500 8900 0.810 0.893 0.435 0.487 6.12 × 105 5.56 × 105
35 600 17 800 0.804 0.914 0.431 0.502 6.27 × 105 5.88 × 105
26 700 26 700 0.797 0.920 0.424 0.506 6.32 × 105 5.98 × 105
17 800 35 600 0.777 0.890 0.397 0.489 6.11 × 105 5.78 × 105
Reaction conditions and design parameters are as given in Table 1.
mized. Suboptimal values can have a serious impact on the performance of FBMR systems. ii. Membranes in the freeboard zone contribute significantly to the reactor performance at high overall membrane capacities, while their contribution is small at low membrane capacities. iii. The optimum ratio of membrane capacity in the two regions depends on the freeboard zone conditions as well as on sweeping parameters, flow, and pressure. Permeate Side Conditions. Another parameter with an impact on the economics of this process technology is the environment on the permeate side, particularly the permeate side pressure and the sweep gas flow. In this work, the sweep gas is assumed to be steam, given its ease of separation from the permeate hydrogen. The sweep gas flow was varied from nil to a value equal to the molar flow rate of the key reactant, methane. The sweep gas pressure was also varied over the range of 0.02-1.0 MPa. The simulation results for the cases studied are presented in Figure 4. The reported results suggest that the pressure of the sweep side is far more important than the sweep gas flow rate. In addition, it is observed that the sweep gas flow has almost no impact on the reactor performance at low permeate side pressures. Scale-up and Practical Issues Operating Environment for Catalytic Membranes. The permeable tubes or other permeable surfaces of a membrane reactor are affected by their
Figure 4. Effect of permeate side pressure and sweep gas flow rate on methane conversion at the bed surface and reactor exit. Reaction conditions and design parameters are as in Table 1.
environment. In particular, they are influenced by permeate concentration gradient, temperature gradient, and sweep gas velocity as well as configuration, surfaceto-volume ratio, mechanical forces applied, and sweep gas conditions. In the present work, the principal factors affecting the performance of membranes in a catalytic membrane reactor are discussed with emphasis on two configuration options: packed and fluidized beds. Each of the factors influencing membrane performance is discussed from two perspectives: (i) impact on membrane permeation capacity utilization and (ii) impact on membrane longevity.
Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997 4553
In principle, FBMR systems can be designed to operate in any of the hydrodynamic regimes of fluidization (Bi and Grace, 1995; Bi et al., 1995). However, heat-transfer and temperature uniformity considerations appear to preclude operation in the low-velocity bubble-free regime, while the need for large-scale operation and relatively small catalyst particles makes operation in the slug flow regime unlikely. Hence, the regimes to be considered are the bubbling regime, turbulent regime, and fast fluidization. The choice is likely to be dictated by competing factors such as compactness, temperature uniformity, high bed-to-factor heat transfer, and limitation of the forces on and erosion of the membrane surfaces. The latter factor is likely to be favored by high-velocity regimes (turbulent or fast fluidization), while the bubbling-bed regime may be best in terms of the other criteria. So far, all FBMR experience has been in bubbling beds, so we discuss the FBMR in terms of the bubbling-bed regime in this paper. It should be borne in mind, however, that the turbulent regime in particular may also prove to be advantageous. Tubes in fluidized bed are almost always either horizontal or vertical, since other orientations lead to bypassing and relatively low heat-transfer rates (Harrison and Grace, 1971). In the bubbling-bed regime, horizontal tubes are subject to strong time-varying buffeting forces due to the passage of bubbles and pressure waves arriving from the bed surface (Kennedy et al., 1981; Turner and Irving, 1982; Grace and Hosny, 1985). The forces tend to be especially large when bubbles coalesce just as the front bubble reaches the tube surface (Levy and Bayat, 1989; Zhu et al., 1989). Some damping occurs when tubes are in the interior of a bundle of horizontal tubes (Hosny and Grace, 1984), but forces remain much higher than those for comparable vertical tubes. Heat transfer to or from vertical tubes is at least as favorable as that for horizontal tubes. Moreover, the vertical orientation avoids the tubes having to bear the weight of particles sitting on them during periods of shutdown. Given all of these factors, the vertical orientation is clearly the favored one, even though mechanical design may then be more complex. This is also the orientation used in our previous work (Adris et al., 1994a), and we assume vertical tubes in the rest of this paper. Temperature Profile. The operating temperature profile to which a membrane surface is exposed is critically important for the performance of a membrane. First, due to the delicate nature of membrane construction based on dense metallic materials, strong temperature gradients can create thermal stresses, which, in turn, can compromise the mechanical integrity of the membrane element. In this respect, fluidized-bed and fixed-bed operating modes represent two extremes of temperature gradients. Fluidized beds provide virtually uniform temperature over a wide range of superficial gas velocities. The contrary is true for packed beds, where isothermal conditions are seldomly achievable for either exothermic or endothermic reactions. The issue of fluidized-bed thermal uniformity for a highly endothermic (SMR) reaction system was addressed by Adris et al. (1994a), who observed that at a gas velocity of about 5 times the minimum fluidization velocity the catalyst bed temperature becomes fairly uniform. A membrane surface immersed in a fluidized
bed operated above this velocity experiences virtually isothermal operation. Note that the permeation process is highly temperature dependent, with the permeation rate constant following Arrhenius behavior. Therefore, a higher operating temperature should lead to a higher permeation rate. On the other hand, too high a temperature shortens the membrane life. For practical cases there should be an optimum operating temperature that gives a desirable permeation rate without unduly compromising the life of membrane. The optimum temperature is determined by the economics of the membrane installation and is different for each application, depending on the type of reaction, its kinetics, and the permeation capacity requirement. One should note in this context that strong temperature gradients, as in packed beds, work against the achievement of an optimal membrane operating temperature. The temperature profile imposed by the operating mode of the catalyst bed tends to compromise the level of utilization of the membrane capacity because the “optimum” temperature range can only be realized over a limited portion of the membrane surface, while the rest of the membrane is exposed to less favorable temperatures, with colder temperatures giving lower permeation rates and higher temperatures curtailing membrane life. Concentration Profile. Like temperature profiles, concentration profiles tend to be steep in packed-bed reactors. However, while permeability of a desired component is favored by a high partial pressure or concentration of that component on the reaction side, as per eq 1, there is no negative impact associated with the exposure of membrane surfaces to very high concentrations of the permeating component, as is the case with membrane deterioration at elevated temperatures. For fluidized-bed reactors the concentration issue is more complicated because of the differences in concentration between the bubbles or voids and the dense phase. Most reaction tends to take place in the dense phase, so that reactant concentrations tend to be lower there (and product concentrations higher) than in the bubble phase. Since the surface of a membrane immersed in the fluidized catalyst bed is exposed to both phases, the driving force for the permeation process across the membrane varies from the distributor level up to the surface of the catalyst bed. The concentration profile depends on the mass exchange rates between the phases as well as on the bubble volume fraction. Both are functions of superficial gas velocity and of bubble coalescence and splitting. The net permeation rate in a fluidized-bed membrane reactor system is affected by both the bubble and dense phases. In an effort to quantify such a property in a fluidizedbed membrane reactor (FBMR), experimental measurements from a pilot-scale study (Adris et al., 1994a) have been used to estimate a permeability effectiveness factor for steam methane reforming (SMR) over a wide range of operating conditions (at different operating temperatures and hydrogen partial pressures on the reaction side). Actual permeation rates were compared with theoretically calculated rates based on permeability rate constants measured by Katsuta et al. (1979). The deficiency in membrane performance due to bubble bypassing, and possibly due to other effects such as resistance offered by a dust layer deposited on the membrane surface, discounts the membrane capacity by as much as 60%.
4554 Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997
Construction and Mechanical Aspects. In addition to thermal stresses and the total pressure difference between the inside and outside of the membrane discussed above, a number of other mechanical factors can affect the performance and longevity of the membrane surface. In order to address the mechanical forces acting on membrane surfaces in the two reactor configurations considered in this paper, packed and fludized beds, it is useful to first treat the construction features of membranes in each configuration. Construction/Mechanical Aspects in a Packed Bed. A packed-bed reactor is generally characterized by poorto-moderate rates of heat transfer and significant temperature gradients, as mentioned above. Reaction conversions and species concentrations in fixed-bed reactors are spatially dependent. These two features of packed-bed reactors lead to the common shape of a multitubular catalyst bed placed inside an enclosure for heat-exchange purposes. This enclosure may be a fired heater, steam still, molten salt baths, or thermal fluid jacket. Dimensions of the catalyst tube are, also, imposed by the above two constraints. Better heatexchange rates require small diameter tubes for desirable surface-to-volume ratios, while reaction completion dictates long tubes. Typical reactor tubes (e.g., steam methane reforming, a highly endothermic reaction) can have diameters as small as 0.1 m and lengths up to 12 m (Hyman, 1986; Rostrup-Nielsen, 1983). Even smaller diameters can also be found for highly exothermic reactions (e.g., 0.0245 m for vinyl acetate reactors; Satterfield, 1980). The number of catalyst tubes in one reactor unit can vary from tens to thousands. There is reasonable consensus in the membrane technology community that tubular membranes are most practical and most suitable for large-scale applications (Tsotsis et al., 1993). This is due to the high surface-to-volume ratio and the ease of interconnections and manifolding using standard piping. A packed-bed membrane reactor then contains a double pipe or multiple membrane tubes inside a tubular catalyst bed placed, in turn, in another enclosure for heat exchange as described above. There is almost no room for interconnections of the membrane tubes whether in the axial or the radial direction. Therefore, the length of the membrane tubes should also be extended to maximize the permeation surface area. The key mechanical forces acting on membrane tubes in a packed-bed reactor are the differential thermal expansion and the associated potential tube buckling, as well as the weight of the catalyst column. Construction/Mechanical Aspects in a Fluidized Bed. Fluidized beds are known to present quite different, harsher, mechanical environments than packed beds. Forces and vigorous particle motion caused by gas bubbles within the bed are responsible for the harsh mechanical conditions in fluidized-bed systems. Fluidized beds often include tube internals, usually for heatexchange purposes. Buffeting action by bubbles and tubes wear caused by solid particles could affect the longevity of the tube internals. A number of studies have therefore been carried out which have addressed mechanical forces acting on tube internals of different orientations. While there has been considerable work on transient forces (e.g., Kennedy et al., 1981; Turner and Irving, 1982; Grace and Hosny, 1985; Pell, 1990) and wear (e.g., Zhu et al., 1990, 1991) with respect to horizontal tubes in bubbling fluidized beds, very little has been done with
respect to vertical tubes. The peak forces are clearly much smaller for vertical tubes than for horizontal ones, but vertical tubes do, nevertheless, experience some transient forces, as they are observed to undergo vibrations when immersed on bubbling beds (Grace and Harrison, 1968). The transverse forces probably originate in the lateral motion of bubbles which occurs when bubbles travel with substantial horizontal components as they interact and coalesce with each other. Pressure waves traveling faster through dense regions of the beds than through more dilute regions may also cause transient horizontal forces on tubes. Research is clearly needed to measure the horizontal stresses and the related wear and erosion. As a first estimate the stresses may be assumed to be of order Fp(1 - mf)(Ub cos θ)2/2, which for typical catalyst particles is of order 100 N/m2. The typical frequency is likely to be of order 1 Hz, the same as for horizontal tubes in bubbling beds. The FBMR reactor also poses other research questions that have received scant attention at best in the extensive literature on fluidized beds. First, the withdrawal of gas through the permeable surfaces could lead to local defluidization in the vicinity of the surfaces. When hydrogen is being withdrawn through permeable tubes in reforming or dehydrogenation reactors, the partial pressure of the diffusing species, and hence the withdrawal, will be highest in the dense phase (where the reaction is most rapid), and this is likely to promote local defluidization. However, countering this effect is the fact that these processes also involve increases in the total number of moles as the reaction proceeds, with most of the reaction (and hence production) occurring in the dense phase. Hydrodynamic studies (Adris, 1994) indicate that at least some of the resulting extra moles are transferred to the bubble phase. Vertical tubes inserted into fluidized beds can have a profound influence on bed hydrodynamics. It is essential to maintain a separation distance between interior surfaces (e.g., between adjacent tubes or between the outer wall and nearest tubes) of about 30 mean particle diameters or more to avoid bridging and sticking of the particles leading to local defluidization (Grace and Harrison, 1968). Violating this condition is likely to lead to local temperature gradients and a serious loss of reactor effectiveness. This condition sets a limit on the number of tubes or other surfaces which can be inserted into a given cross section of bed. Below we suggest that tubes could be bundled together in an exterior sheath to overcome this limitation, while also helping to protect the thin-walled permeable tubes from lateral forces, wear, and plugging by particles. Clearly, working with small catalyst particles also helps, but for hydrodynamic reasons, these should not generally be smaller than about 50 µm in mean size. Note that PBMR reactors will also be adversely affected if tubes are too close together, although in this case the effects are unlikely to be as dramatic. In the packed-bed case, it is well-known that particles do not pack as closely near solid fixed surfaces as in the interior of the bed, causing some short-circuiting of gas to occur along the fixed surface. The distance between adjacent surfaces should be approximately 16 times the mean packing diameter to make this effect negligible. Since the packing diameter for the PBMR is bound to be many times the particle size in the FBMR, satisfying the spacing constraint is likely to be a serious problem in the PBMR also.
Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997 4555
Once the inter-tube spacing is sufficient to assure that defluidization does not occur, there are other influences of the tubes on bed behavior (Grace and Harrison, 1968; Harrison and Grace, 1971; Yates et al., 1984). If the bubbles are of smaller diameter than the immersed surfaces (which may well be the case for tube bundles), then the voids tend to act like slugs between the fixed surfaces. However, for tubes much smaller in diameter than the bubbles, which is bound to be the case for single tubes, the bubbles tend to envelope the tubes, causing the bubbles to elongate, rise more quickly than they would otherwise, and be less likely to be diverted horizontally in pursuit of other bubbles. This results in a bed which is more uniform radially. Hence, the tubes can be beneficial to the overall bed hydrodynamics, provided that the spacing constraints are satisfied. Adris et al. (1994b) proposed a modular design for membrane tubes to be used in a fluidized-bed membrane reactor for steam methane reforming. In the proposed design a number of membrane tubes (thin-walled metal tubes or thin metal film deposited on a porous substrate) are bundled together and shielded by a porous sheath. This outer shield lets all gases through, but no solids. This design should minimize the harmful effect of buffeting and wear caused by gas bubbles and solid circulation, yet maintain the gas environment needed for permeation to occur. It also helps overcome the constraints of inter-tube spacing discussed above. Conclusions Characteristics of catalytic membrane reactors are considered, with emphasis on the reactor and membrane surface configuration. Simulation was used to give typical designs for commercial-size FBMR serving as a steam methane reformer and to show the ability of the FBMR to accommodate the required permeation capacity. The impact of distributing the membrane capacity between the dense bed and dilute phase has also been studied by means of reactor simulation, and results indicated this parameter to be an important design feature which needs to be optimized for each particular application. The effects of permeate side pressure and sweep gas flow rate have also been investigated, and results suggested that the sweep gas flow rate has a limited impact on the FBMR performance compared to the impact of the pressure on the permeate side. Thermal uniformity, offered by fluidized systems, is advantageous for membrane capacity utilization as well as for minimizing thermal stresses on the membrane material. The fluidized bed also has a greater ability to accommodate membrane surfaces with connection and orientation flexibility compared to packed beds. Key mechanical stresses acting on membrane structures have been identified to be buckling and catalyst weight in packed beds, while in fluidized beds buffeting action and erosion of surfaces by moving solids are the principal concerns. It is clear that there are several areas that require additional research and clarification if FBMR reactors are to become commercially viable. Key areas requiring investigation include the magnitude of transverse forces on vertical tubes and the effects on fluidization behavior of gas withdrawal through permeable surfaces and of gas generation due to reaction. Others having to do with the long-term durability of the membrane material, the maintainance of the permeable surface free from
blockage, and the optimum sweep gas configuration are also important. Acknowledgment The authors are grateful to Prof. C. J. Lim for input to this study. Nomenclature Am ) membrane surface area, m2 C0 ) standard hydrogen solubility in palladium, kmol/ m3‚MPa1/2 Cep ) equivalent permeation capacity (membrane surface area divided by permeation layer thickness), km d ) thickness of the membrane wall, m DF ) Fickian diffusion coefficient for hydrogen dissolved in palladium, m2/s kH ) permeation rate constant, kmol/s‚MPa1/2 PhH, PlH ) partial pressure of hydrogen on the high and low sides of the membrane respectively, MPa QH ) permeating hydrogen flow, kmol/s ub ) bubble rise velocity, m/s Fp ) particle density, kg/m3 mf ) bed voidage at minimum fluidization θ ) angle of bubble rise to the vertical
Literature Cited Adris, A. M. A fluidized bed reactor for steam methane reforming: experimental verification and model validation. Ph.D. Thesis, University of British Columbia, Vancouver, Canada, 1994. 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. 1991a, 69, Oct, 1061. Adris, A. M.; Grace, J. R.; Lim, C. J.; Weherhahn, E.; Safavi, S.; Zhu, J. Effect of Low-Clearance Vertical Internals on Mechanical Attrition and Fluidization Quality in a Bubbling Fluidized Bed. 41st Canadian Chemical Engineering Conference, Vancouver, Canada, Oct 1991b. Adris, A. M.; Lim, C. J.; Grace, J. R. The Fluidized Bed Membrane Reactor (FBMR) System: a Pilot Scale Experimental Study. Chem. Eng. Sci. 1994a, 49, 5833. Adris, A. M.; Grace, J. R.; Lim, C. J.; Elnashaie, S. S. E. H. Fluidized Bed Reaction System for Steam/Hydrocarbon Gas Reforming to Produce Hydrogen. U.S. Patent 5,326,550, 1994b. Adris, A. M.; Lim, C. J.; Grace, J. R. The Fluidized Bed Membrane Reactor for Steam Methane Reforming: Model Verification and Parametric Study. Chem. Eng. Sci. 1997, 52, 1609. Bi, H. T.; Grace, J. R. Flow Regime Diagrams for Gas-Solid Fluidization and Upward Transport. Int. J. Multiphase Flow 1995, 21, 1229-1236. Bi, H. T.; Grace, J. R.; Zhu, J. Regime Transitions Affecting GasSolids Suspensions and Fluidized Beds. Trans. Inst. Chem. Eng. 1995, 73, 154-161. Bohmholdt, G.; Wicke, E. Diffusion of H2 and D2 in Pd and Pdalloys. Z. Phys. Chem. 1967, 56, 133-148. Buxbaum, R. E.; Kinney, A. B. Hydrogen Transport through Tubular Membranes of Palladium-Coated Tantalum and Niobium. Ind. Eng. Chem. Res. 1996, 35, 530. Chopra, K. L. Thin Film Phenomena; McGraw-Hill Book Co.: New York, 1969. Goetsch, D. A.; Say, G. R. Synthesis Gas Preparation and Catalyst Therefor. U.S. Patent 4,877,550, 1989. Grace, J. R.; Harrison, D. The Effect of Internal Baffles in Fluidized Beds: a Guide to Design. Inst. Chem. Eng. Symp. Ser. 1968, 27, 93-100. Grace, J. R.; Hosny, N. Forces on Horizontal Tubes in Gas Fluidized Beds. Chem. Eng. Res. Des. 1985, 63, 191-198. Graham, T. Philos. Trans. R. Soc. (London) 1866, 156, 399. Harrison, D.; Grace, J. R. Fluidized beds with internal baffles. In Fluidization; Davidson, J. F., Harrison, D., Eds.; Academic Press: London, 1971; Chapter 13, pp 599-626. Hosny, N.; Grace, J. R. Transient Forces on Tubes within an Array in a Fluidized Bed. AIChE J. 1984, 30, 974-976.
4556 Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997 Hsieh, H. P. Inorganic Membrane ReactionssA Review. AIChE Symp. Ser. 1989, 85, 268. Hyman, M. H. Simulate Methane Reformer Reactions. Hydrocarbon Process. 1986, 47, 131-137. Itoh, N. A Membrane Reactor Using Palladium. AIChE J. 1987, 33, 1576. Katsuta, H.; Farraro, R. J.; McLellan, R. B. The Diffusivity of Hydrogen in Palladium. Acta Metall. 1979, 27, 1111-1114. Kennedy, T. C.; Donovan, J. E.; Trigas, A. Forces on Immersed Tubes in Gas Fluidized Beds. AIChE J. 1981, 27, 351-357. Lee, S. M. Film Deposition Techniques. Encycl. Chem. Technol. 1980, 10, 247. Levy, E. K.; Bayat, R. The Bubble Coalescence Mechanism of Tube Erosion in Fluidized Beds. In Fluidization VI; Grace, J. R., Shemilt, L. W., Bergougnou, M. A., Eds.; Engineering Foundation: New York, 1989; pp 605-611. Mleczko, L.; Ostrowski, T.; Wurzel, T. A Fluidized-Bed Membrane Reactor for the Catalytic Partial Oxidation of Methane to Synthesis Gas. Chem. Eng. Sci. 1996, 51, 3187. Ostrowski, T.; Giroir-Fendler, A.; Mirodatos, C.; Mleczko, L. Comparative Study of the Catalytic Partial Oxidation of Methane to Synthesis Gas in Fixed-Bed and Fluidized-Bed Membrane Reactors. Catal. Today, 1997, in press. Pell, M. Gas Fluidization; Elsevier Science Publishers: Amsterdam, The Netherlands, 1990. Rostrup-Nielsen, J. R. Catalytic Steam Reforming. In Catalysis Science and Technology; Anderson, J. R., Boudard, M., Eds.; 1983; Vol. 4. Springer: Berlin. Saracco, G.; Specchia, V. Catalytic Inorganic Membrane Reactors: Present Experience and Future Opportunities. Catal. Rev.sSci. Eng. 1994, 36 (2), 305-384. Satterfield, C. N. Heterogeneous Catalysis in Practice; McGrawHill Book Co.: New York, 1980. Turner, M. J.; Irving, D. Forces on a Tube Immersed in a Fluidized Bed. Proc. 7th Int. Fluidized Bed Combust. Conf. 1982, 2, 831839.
Tsotsis, T. T.; Champagniem, A. M.; Minet, R. G.; Liu, P. K. T. Catalytic Membrane Reactors. In Computer-Aided Design of Catalysts; Beeker, E. R., Pereira, C. J., Eds.; Chemical Industries Volume 51; Marcel Dekker: New York, 1993. Wen, C. Y.; Chen, L. H. Fluidized Bed Freeboard Phenomenas Entrainment and Elutriation. AIChE J. 1982, 28, 117. Yates, J. G.; Cheesman, D. J.; Mashingaidze, T. A.; Howe, C.; Jefferis, G. The Effect of Vertical Rods on Bubbles in Gas Fluidized Beds. In Fluidization; Kunii, D., Toei, R., Eds.; Engineering Foundation: New York, 1984; pp 103-110. Yeung, K. L.; Varma, A. Novel Preparation Techniques for Thin Metal-Ceramic Composite Membranes. AIChE J. 1995, 41 (9), 2131-2139. Zhu, J.; Grace, J. R.; Lim, C. J. Erosion-Causing Particle Impacts on Tubes in Fluidized Beds. In Fluidization VI; Grace, J. R., Shemilt, L. W., Bergougnou, M. A., Eds.; Engineering Foundation: New York, 1989; pp 613-620. Zhu, J.; Grace, J. R.; Lim, C. J. Tube Wear in Gas Fluidized Beds. I. Experimental Findings. Chem. Eng. Sci. 1990, 45, 10031015. Zhu, J.; Lim, C. J.; Grace, J. R.; Lund, J. A. Tube Water in Gas Fluidized Beds. II. Low Velocity Impact Erosion and SemiEmpirical Model for Bubbling and Slugging Fluidized Beds. Chem. Eng. Sci. 1991, 46, 1152-1156.
Received for review February 3, 1997 Revised manuscript received August 22, 1997 Accepted August 25, 1997X IE970108W
X Abstract published in Advance ACS Abstracts, October 1, 1997.
