2940
Ind. Eng. Chem. Res. 1995,34, 2940-2948
Improving Methylcyclohexane Dehydrogenation with ex-Situ Hydrogen Separation in a Reactor-Interstaged Membrane System Jawad K. Ali, David W. T. Rippin,? and Alfons Baiker" Department of Chemical Engineering and Industrial Chemistry, Swiss Federal Institute ETH-Zentrum, CH-8092Zurich, Switzerland
of
Technology,
A system of two plug-flow reactors with a n interstaged tubular palladium-silver (Pd-Ag) membrane to ex-situ remove H2 was used to improve the efficiency of methylcyclohexane dehydrogenation. A monometallic noble metal catalyst produced yields of toluene and hydrogen from methylcyclohexane which were much higher than those corresponding to equilibrium a t industrially relevant conditions. Experiments in this system a t 593-673 K, 1-2 MPa, and overall liquid hourly space velocities of 1-2.6 h-l afforded conversions up to 52% greater than those without membrane and 35% greater than the equilibrium values for a conventional reactor. However, the equilibrium conversions for the system itself were never exceeded. The membrane capacity was adequate to remove -90% of the HZpresent in the first reactor product. Results obtained with the two reactors-membrane system were close to those predicted using first reactor conversion data.
Introduction The potential utilizations of energy sources such as hydropower would be increased if effective energy storage systems were available. For countries with a surplus of hydropower in summer, e.g., Switzerland, a method of storing this surplus for winter may be economically desirable. H2 storage, in the form of liquid organic hydrides, e.g., methylcyclohexane (MCH), was proposed for the seasonal storage of electricity (Taube and Taube, 1981). In summer, low-cost electricity is used for electrolysis of water to yield H2 which can be catalytically combined with toluene (TOL) to MCH. In winter, catalytic dehydrogenation of MCH releases the H2 for application as electricity by means of a turbine generator or fuel cells and so forth. Newson et al. (1993) showed that seasonal storage of H2 in MCH was economically competitive with all the other carbon-free alternatives. They also indicated that stationary H2 systems can be economically comparable with new hydropower projects. The use of H2 energy for both mobile (Griinenfelder and Schucan, 1989; Winter, 1990a; Manser et al., 1991) and stationary (Winter et al., 1990b)applications has been described. This work is related to the stationary applications of the MTH system (methylcyclohexane, toluene, hydrogen) and is aimed at improving the efficiency of the MCH dehydrogenationreaction. The precedent demonstrates that MCH dehydrogenation is an important step in the Hz energy storage cycle; however, it is highly endothermic and strongly equilibrium-limited by the product H2, requiring rather high operating temperatures. Provisions for the selective removal of H2 by a suitable membrane should in principle relieve the thermodynamic restraint and implies a lower operating temperature. The membrane separation of high-purity H2 would also allow the utilization of H2 in an energy storage cycle, e.g., in fuel cells, since sweep gas may not be needed with dense membranes. The concept of combining reaction and separation in a membrane reactor to in-situ separate a product so as
* To whom correspondence should be addressed. E-mail:
[email protected]. ' Deceased June 26, 1994.
to increase conversions in equilibrium-limited reactions is well-known. Originating in the 1960s, Pd alloy membranes were used t o selectively remove H2 from dehydrogenation reactions and conversions higher than the traditional equilibrium limits were realized (Pfefferle, 1966; Wood, 1968). Since then, numerous industrially important reactions involving the simultaneous removal of H2 by membranes have been examined as reported in the literature. These include dehydrogenation of ethane (Gobina and Hughes, 1994; Champagnie et al., 1990, 19921, propane (Ziaka et al., 19931, isobutane (Matsuda et al., 1993; Zhu et al., 19931, n-butane (Zaspalis et al., 19911, cyclohexane (Okubo et al., 1991; Itoh, 19871, methylcyclohexane (Ali et al., 1994a), ethylbenzene (Wu et al., 1990; Abdalla and Elnashaie, 19941, and formaldehyde (Deng and Wu, 1994). Examples of other relevant reactions are methane steam reforming (Shu et al., 1994; Minet et al., 1992; Uemiya et al., 1991a), water gas shift reaction (Uemiya et al., 1991b), and propane aromatization (Uemiya et al., 1991~).The majority of investigators claim exceeding equilibrium conversions and better selectivities achieved with membrane reactors. A number of comprehensive reviews (Saracco and Specchia, 1994; Zaman and Chakma, 1994; Saracco et al., 1994; Tsotsis et al., 1993; Shu et al., 1991; Hsieh, 1991; Armor, 1989) addressing the potentials of membrane reactors, the reactions tested, the properties of inorganic membranes, and obstacles ahead of practical applications have recently been published. Of particular interest t o our work is the review by Shu et al. (1991) on catalytic Pd-based membrane reactors, in which the various reactions studied, the current techniques for preparing thin metallic films, and a number of reactor designs were discussed. Almost the entire reported studies involved membrane reactors with in-situ product separation. For dehydrogenationreactions, although membrane reactor combines reaction and in-situ HZremoval in a single unit, ex-situ separation of H2 may offer some attractive features over the in-situ mode in the case of catalysts with sufficient dehydrogenation activities (Mi and Rippin, 1994). Nevertheless, future commercial exploitation of membrane reactors necessitates the circumven-
0888-5885/95/2634-2940$09.Q0~Q 0 1995 American Chemical Society
Ind. Eng. Chem. Res., Vol. 34, No. 9, 1995 2941 tion of a number of hurdles. Some of the hurdles outlined by Saracco et al. (1994) are the following: 1. The sealing of membranes into modules is an obstacle to be overcome. 2. There are mechanical stability problems due to thermal and startup/shutdown cyclic stresses since the Pd membranes are confined within the solid catalyst bed. 3. The task of charging the catalyst into reactor tubes in a shell-and-tube configuration containing a tubular membrane inside each tube is a difficult one. 4. The susceptibility of Pd membranes to strong poisons such as sulfur and chlorine which may come from the catalyst during activation is a problem. 5. For endothermic dehydrogenation reactions, supplying heat to a high temperature membrane reactor is not a trivial task. Besides the aforementionedobstacles, membrane reactors are restricted to a fixed specific membrane and a leaking Pd membrane tube in a reactor module cannot be easily identified. For comparable duties, there are several advantages in applying the ex-situ over the in-situ H2 separation modes: 1. Existing traditional reactors can be used with interstage membrane units. 2. Suitable H2 diffusion units are commercially available (Philpott and Coupland, 1988) which can be combined with existing preheaters between reactors for endothermic reactions. 3. Separate activation or regeneration of the catalyst and the membrane is possible so as to minimize membrane deactivation. For example, germanium and rhenium in platinum-based catalysts may form volatile compounds during such operations (Ali et al., 1994c; Garland et al., 1991) which upon decomposition deposit the element on the surface of the membrane and impair its H2 permeability. 4. Future development of thin Pd alloy films would reduce the membrane unit size and makes such processes more attractive economically. 5. Operation at permeate H2 pressures above atmospheric pressure may be possible, allowing ultra pure Ha t o reach users' utilities. 6. Source gas can be passed inside the Pd alloy tubes for higher pressure resistance. 7. Large scale applications of Pd alloy diffusion units, capacities up t o 4.106 ft31day, have been successful for several years (Hwang and Krammermeyer, 1975). Whether H2 separation is in-situ or ex-situ, the necessity t o develop asymmetric composite membranes with Pd-based deposits < 1pm is stressed out. It was reported (Guy, 1992)that the membrane area necessary for a typical propylene production of 100 000 tonslyear is reasonable, Le., 50 m2 for a 1:l mole ratio of HdC3H8 feed in the case of a dense Pd membrane 1 pm thick. Indeed, recent research activities are focused on the preparation of thin Pd-based alloy films deposited on porous metallic and ceramic supports (Govind and Atnoor, 1991; Shu et al., 1993; Uemiya et al., 1991d,e; Gryaznov et al., 1993). Here we report MCH dehydrogenation in a system of two plug-flow reactors coupled in series with interstaged tubular Pd-Ag membrane to ex-situ separate the H2 produced in the first reactor. Using a catalyst and PdAg membranes, both commercially available, our aim was to improve the TOL and H2 yields over those achievable in a conventional packed-bed reactor at economically viable throughputs. Interestingly, Rezac
et al. (1994) have recently applied the same concept t o n-butane dehydrogenation and reported a significant increase in conversion over the equilibrium limit. A particular feature of this study is the elevated pressure experiments, up to 2 MPa. As previously discussed, a major hurdle in setting up an experimental membrane reactor is the leakproof sealing of the membrane. This seems to have restricted most previous investigators to atmospheric pressure operation. An exception to this, is the interesting experiment reported by Schmitz et al. (Schmitz and Gerke, 1987, 1988a,b; Schmitz et al., 19881, who examined the steam reforming of methane a t 1.05 MPa and 973 K. An industrial scale packed-bed reformer tube (1.4 m long, 4.5 cm inside diameter) with a Pd-Ag membrane sealed at its lower part was employed. The membrane (100 pm wall thickness, surface area 200 cm2) was supported by a perforated Inconel tube having 0.4 mm diameter holes spread a t 1mm intervals. They reported a gain of 25% in methane conversion with the packed-bed membrane reactor over the nonmembrane reactor.
Theory MCH dehydrogenation to TOL proceeds only under thermodynamically favorable conditions.
C,H,, ==C,H,
+ 3H2
AH:*
= 204 kJ/mol (1)
The following equations relate the equilibrium constant,
Keq,with temperature and composition, Xe:
Keq= 4.61(109) exp{-2:350
(L T-L 650)] (2) (3)
Equation 2 was established by Rimensberger (19871, who corrected literature values for the preexponential factor (entropy term) of the equilibrium constant using experimental data from a microcontinuous reactor. Equations 2 and 3 give the composition of a mixture of MCH, TOL, and H2 at equilibrium in a closed system where no exchange of matter occurs with the surroundings. For a continuous fixed-bed reactor under steadystate conditions, mass balance imposes that the feed and the product must be equal. Thus a reactor packed with a sufficient amount of an active catalyst can attain equilibrium in a manner analogous to a closed system. By incorporating a membrane to remove a product, a different equilibrium state is created since the system is no longer a closed one. This implies that it is impossible to exceed equilibrium conversion in a specific system and the enclosure of a membrane modifies the system, resulting in a shift to a new equilibrium. A shift to higher equilibrium conversions is also possible without using membranes, by diluting the MCH feed with an inert gas. It follows that statements often encountered in the literature about exceeding the equilibrium limits with membrane reactors are not precise unless the system relative t o which the equilibrium is crossed is clearly identified. In a membrane reactor with insitu product separation, a new equilibrium corresponding to the composition of the reaction mixture at different positions along the reactor axis is established.
2942 Ind. Eng. Chem. Res., Vol. 34, No. 9, 1995 Permeate
Feed3 "y I
Catalyst, \
% H,.removcd &om tint reactor product by an interstaged membrane
(1) 0.0 (2) 30
t
(3) 70 (4) 90 (5) 100
0
310
330
350 370 390 Temperature ("c)
410
Figure 1. Equilibrium conversions for MCH dehydrogenation in a system of two plug-flow reactors with a membrane interstaged.
Our current investigation involved the dehydrogenation of MCH in a system of two plug-flow reactors with an interstaged membrane t o ex-situ remove Ha from the first reactor product. For this system, the equilibrium conversion, Xes (mole ratio of MCH converted to TOL in the two reactors to the MCH feed), was calculated by SIMUSOLV (Steiner et al., 19901, applying eq 3 once to each reactor, as given in Figure 1. As the quantity of H2 removed through the membrane increases, so does the shift in equilibrium conversion. The maximum possible shifts in equilibrium, corresponding to 100% H2 removal, occur at 613-633 K. As a result, our experimental setup was designed to ensure -90% H2 removal by the membrane, thus avoiding the use of vacuum or sweep gas inside the membrane.
Experimental Apparatus The goal of this work was to determine the effect of interstage H2 removal by Pd-Ag membrane on the performance of a MCH dehydrogenation system which consisted of two plug-flow reactors in series. The reaction was operated under steady-state conditions, and the effects of varying the reaction temperature, pressure, liquid hourly space velocity (LHSV), and H2 pressure inside the membrane were evaluated. Pure MCH, Fluka purum, >99% GC, 0.2%TOL, rest CS-, was used as received from the manufacturer. A detailed description of the reaction system, reactors and, membrane are provided. Catalyst and Membrane. The catalyst was a commercial monometallic noble metal supported on spherical alumina particles -1.6 mm diameter, containing chlorine and sulfur. This catalyst was selected because it provided superior dehydrogenation activity and selectivity, compared to other catalysts screened, and maintained almost constant activity (Ali and Rippin, 1995;Manser, 1992). It was supplied under confidential agreement by UOP Ltd. (UK). Thirteen milliliters (6.688g) of the undiluted catalyst was packed in each
Quart. Pd-Ag Membra
Figure 2. The reactor-membrane system.
reactor and supported between two sections of quartz granules (size -1 mm). The top quartz bed served as a preheater for the reactants. The membrane was tubular Pd-23 wt % Ag, 3.1 mm diameter with a 0.1 mm wall thickness, open a t one end. Equipment. Details of the experimental apparatus are available elsewhere (Mi et al., 1994a,b), and only key features pertinent t o this study are emphasized here. It consisted of two stainless steel (SS) plug-flow reactors in series (25 cm long, 1.9 cm in diameter each) and two interstaged tubular membranes as detailed in Figure 2. Each membrane was sealed along the center line of a 9.5 mm diameter SS tube. The two membranes were connected in series to ensure the removal of -90% H2 from the first reactor product. The exposed length of the membranes to the reactants was 48 cm (24 c m / membrane). All four units were placed in a fluidized sand bath with temperature control ( f l"C), a pressure control ( f 0 . 2 bar), and mass flow controllers providing accurate reactor mass balances (f2%). The product of the first reactor acted as feed to the membranes. The effluent from the membranes was fed t o the second reactor. The permeate Hz from the two membranes was combined in a single stream. A typical run involved monitoring the feed rate to reactor 1, H2 permeation rate, reactor 2 product composition, and axial temperature profiles of both reactors by sliding thermocouples inside 3.1 mm thermowells sealed in the centers of the catalyst beds. Two thermocouples in the fluidized sand bath monitored the bath temperature. A Lewa diaphragm pump coupled to a flow pulsation dampener was used to feed MCH into the system giving uniform flow, and a balance monitored the feed rate continuously. Liquid product samples were taken about every hour. Ha permeation rates were measured over a time period of -20 min by a silicone oil (grade 47 v 100) wet gas meter. The H2 pressure inside the membranes, P,, was atmospheric unless indicated otherwise. Sweep gas was not used to avoid contaminating the H2. N2 (99.999%), Hz (99.999%),which were further cleaned by deoxo and moisture trap units, and air (99.995%)were used. The liquid product samples (condensed in an ice-water bath operated a t atmospheric pressure) were analyzed by a gas chromatograph with a flame ionization detector using a 24 ft, 3.1 mm diameter SS column packed with SE-30 as the stationary phase. Leakproof sealing of the membrane is very crucial under the somewhat severe operating conditions used in this work. In our experiments, the membranes were sealed along the center line
Ind. Eng. Chem. Res., Vol. 34, No. 9, 1995 2943 I
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X. = Equilibrium Conversion 0
F e e d 4 Figure 3. Details of sealing the top end of the membrane into the shell.
of the SS tubes by using Swagelok SS compression ferrules with 10 mm long pieces of SS tubes inside their top ends as shown in Figure 3. Catalyst and Membrane Activation Procedure. Weakly adsorbed sulfur (S)was removed from the first reactor catalyst by H2 reduction and air oxidation at 693 K t o minimize poisoning the membranes by S (Ali et al., 1994b). The desorption of some S was evident by the smell of hydrogen sulfide. The second reactor catalyst was used as received from the manufacturer with no prior treatment. H2 (100 mumin) was passed through the system as it was heated to 673 K. After air oxidation of the membrane surface at 673 K for 30 min, the H2 permeabilities of the two membranes were a few percent lower than those reported for the clean membrane (Ali et al., 1994b). Methylcyclohexane Runs. H2 flow through the reactor-membrane system was stopped and MCH was pumped until a constant pressure was reached. Then an experimental run was started recording the various relevant data while MCH was being pumped at a constant rate. At selected conditions between the runs, the rate of H2 permeation through the membranes was checked by stopping the MCH flow and replacing it by H2. A few runs were carried out with the H2 pressure inside the membrane varied between 0.1 and 0.3 MPa. Runs were also conducted with the first reactor only and the two reactors with the permeate H2 line closed; the data were then used for analyzing the system. MCH feeding was stopped overnight because of safety considerations, and the system was left under H2 flow at reaction temperature.
Results and Discussion The system was operated with the reactants flowing downward through the catalyst beds and the permeate H2 upward inside the membranes as illustrated in Figure 2. MCH dehydrogenation is a highly endothermic reaction, and consequently the reactor operation was nonisothermal.
0.0
0.5
1.0
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Overall LHSV
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/ Two Reactors (h-l)
Figure 4. MCH conversion to TOL vs overall LHSV in the reactor-membrane system. 100
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2.0
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3.0
Overall LHSV / Two Reactors (h-') Figure 5. MCH conversion to TOL vs overall LHSV in the reactor-membrane system.
Performance of the Reactor-Membrane System. Figures 4-6 show the experimental results in terms of MCH conversion to TOL in the reactor-membrane system at 593-673 K, 1-2 MPa, and overall LHSVs of 1-2.6 h-l based on the two catalyst beds (2-5.2 h-ll catalyst bed). The equilibrium conversion values (Xe indicated in figures) for a reactor without membrane are those based on the work of Rimensberger (1987). These figures show that MCH conversion to TOL in the reactor-membrane system exceeded the equilibrium limits for a conventional reactor at the temperatures of
2944 Ind. Eng. Chem. Res., Vol. 34, No. 9, 1995
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310
330
350
370
390
I
410
Temperature ("C)
Figure 6. MCH conversion to TOL vs bath temperature in the reactor-membrane system at 1.5 MPa.
the fluidized sand bath. In fact the experimental conversions are at lower temperatures than those given in Figures 4-6 due to the nonisothermal reactor operation resulting in effluent exit temperatures generally lower than the fluidized bath temperature. The extent by which the equilibrium for a plug-flow reactor was exceeded increased a t the lower temperatures and higher pressures, where strong equilibrium limitations are present. Such conditions favor the backward reaction, the rate of which diminishes after H2 removal through the membrane. For instance, at an overall LHSV of 1.5 h-' and a pressure of 1.5 MPa, the conversions of MCH to TOL for the reactormembrane operation were 2%and 31%greater than the equilibrium values for a reactor without membrane (X,) at 673 and 593 K, respectively. For the sake of clarity, the percent increase in conversions, q s and qse, are calculated as
The equilibrium conversions for the reactor-membrane system (Xes)estimated from Figure 1 were never exceeded, as evident from Figure 6. As mentioned in the Introduction, the MTH system is a closed dehydrogenationhydrogenationcycle. Therefore, a TOL yield not less than 95%is desirable to reduce MCH makeup and byproduct formation. In traditional reactors, temperatures substantially in excess of 673 K are required to achieve this goal. Inspection of Figures 4 and 5 indicates that the conditions yielding 95%TOL lie in the region bounded by the temperature and pressure ranges of 653-673 K and 1-1.5 MPa at an overall LHSV of a 2 h-l, the latter value being used in commercial catalytic reformers.
0
100
200
300
Permeate Hydrogen Pressure
400
pa)
Figure 7. MCH conversion to TOL vs permeate Hz pressure in the reactor-membrane system at 1.5 MPa and an overall LHSV = 2 h-l.
If permeate H2 pressures above the ambient could be tolerated, then a more efficient utilization of HZenergy is possible. In this way, high-purity Ha from the membrane can flow to locations where it is needed. Figure 7 shows the results of varying the permeate H2 pressure between 0.1 and 0.3 MPa. MCH conversion to TOL was only slightly reduced because the H2 separation capacity of the membranes was high. Conversionsin First Reactor. Figure 8 shows the experimental results obtained with the first reactor only. These results will be used later for analyzing the reactor-membrane system. Two common features describe the performances of the first reactor and the reactor-membrane system: (1) MCH conversion decreased as the LHSV increased as a consequence of the reactants' shorter residence time approach, and (2) the experimental data, when extrapolated to a LHSV = 0, tend close to the equilibrium values as illustrated in Figures 5 and 8. The small differences may be attributed to the catalytic selectivities being a few percent lower than 100%. This confirms the reliability of the experimental results and demonstrates the excellent behavior of the catalyst for MCH dehydrogenation. The performance of the catalyst was limited by the heat input since conversions very close to the equilibrium values were obtained by diluting it 10 times with inert particles to reduce the endothermicity (Ali et al., 1994a). The membrane capacity was quite adequate to separate most of the Hz present in the first reactor product (evaluated from Figure 8). For instance, -90% and 95% of H2 were removed at pressures of 1 and 1.5 MPa, respectively. This reduced the equilibrium restraint associated with the reactant feed t o the second reactor. Conversions with and without Membrane. In practice, the reactor operation is nonisothermal, so a direct comparison of MCH conversions to TOL in the reactor-membrane system with those obtained by clos-
Ind. Eng. Chem. Res., Vol. 34, No. 9, 1995 2945 I
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(h-l)
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593 K were achieved. Rezac et al. (1994), using a similar concept on n-butane dehydrogenation, increased conversion from 22% in their two reactors to 33% with a membrane interstaged between them; i.e., 50% greater conversion was achieved corresponding to =loo% removal of H2. Their polyimide-ceramic membrane was operated a t 453 K because it underwent plasticization at lower temperatures, resulting in the loss of some hydrocarbons. The temperature in their reactors was 755 K, hence cooling the feed to and heating the effluent from the membrane cell was necessary. In our work, the membrane was operated at the same temperature as the reactors and only Hz permeated the membrane, thus providing a direct source of high-purity Hz for potential future applications in energy storage systems, e.g., fuel cells. Next the selectivity of the catalyst is discussed. The observed selectivities to TOL in the reactor-membrane system, the two reactors without membrane, and the first reactor were close t o each others. However, the selectivity decreased slightly with increasing pressure and decreasing LHSV, typical of hydrocarbon dehydrogenation reactions. The selectivity varied between 93% and 99% corresponding to pressure bounds of 2 and 1 MPa, respectively. The reactor-membrane system produced a little more benzene when H2 was withdrawn (membrane open) than without Hz removal (membrane closed). For example, 0.2 and 0.07 mol % benzene were present in the liquid product with 96.5% and 95.5% selectivities to TOL corresponding to an open and a closed membrane runs, respectively, at 643 K, 1.5 MPa, and an overall LHSV of 2 h-l. The membrane activity for MCH dehydrogenation in a flow system was marginally small (Ali et al., 199413). Its Hz permeability was nearly constant within a few percent as checked periodically during the runs. Thus, the results reflect the performance of the catalyst only. Prediction of Conversions in the ReactorMembrane System. It is possible to estimate roughly the conversion of MCH to TOL in the reactor-membrane system from that in the first reactor because (i) the rate of reaction becomes first order with respect t o MCH (Manser, 1992) at the entrance t o the second reactor since equilibrium restriction was eliminated by the membrane, (ii) the catalyst maintained almost constant activity throughout the experimental runs, and (iii) both reactors contained equal amounts of catalyst. This would imply approximately equal conversions of MCH to TOL in reactors 1 and 2 of the system. For reactor 1, the degree of MCH conversion to TOL, x , is defined as x=- F - P
0 310
330
350
370
390
410
Temperature ("C) Figure 9. MCH conversion to TOL vs bath temperature in the reactor-membrane system, the two reactors, and the first reactor at 1.5 MPa and an overall LHSV = 1.2 h-l.
ing the membrane outlet would be more realistic. This system with the membranes closed mimics that of two plug-flow reactors in series. Figure 9 shows such a comparison at 593-673 K, 1.5 MPa, and a LHSV of 1.2 h-l. The improvement in conversion in the reactormembrane system over that without membrane increases at lower temperatures; e.g., 8-50% greater conversions corresponding to the temperatures 673 and
F
where F and P represent the moles of MCH in the feed to and the product from the reactor, respectively. Considering reactor 1 of the system, moles of MCH converted to TOL is XF and that in the product is (1x)F,the latter equals the MCH feed t o reactor 2. Moles of MCH converted in reactor 2 is then x ( l - x)F and that converted in reactors 1 and 2, F,,, is XF + x(1 x)F. The overall MCH conversion to TOL in reactors 1 and 2, xov,is simply F,,IF, i.e., x,, = 2x
- x2
(7)
Figure 10 compares the experimentally measured conversions in the reactor-membrane system with
2946 Ind. Eng. Chem. Res., Vol. 34, No. 9, 1995
sion to TOL in the reactor-membrane system can be estimated from that available for a traditional plug-flow reactor provided that the equilibrium restriction associated with the product from the latter reactor is eliminated by H2 removal in a membrane unit.
-@-Observed
-_ -- Estimated from First Reactor Conversions
t
Conclusions
390
We have shown that methylcyclohexane conversion to toluene in a two reactors-membrane system was increased over the equilibrium limit for a conventional reactor through the addition of a membrane to selectively separate Hz ex-situ. The extent by which the equilibrium was exceeded increased as the severity of the equilibrium limitations became stronger. This afforded 8-50% greater conversions in this system than those obtained without the membrane and 2-31% greater than the equilibrium corresponding to the temperatures of 673 and 593 K, respectively. The equilibrium restriction associated with the product from the first reactor of the system was relieved by removing -90% of the Hz produced through the membrane. The catalyst selectivity was in most cases greater than 95%, making it possible to achieve the commercially desirable yield of 95% toluene. Methylcyclohexane conversions in the reactor-membrane system are roughly predictable from those available for a plug-flow reactor provided that the equilibrium restraint associated with the reactor effluent is eliminated. There should be no difficulties in implementing the results of this investigation into practice since Hz diffusion units that can achieve the desired separation are commercially available and provided that such a proposal is economically warrantable.
380 -
Acknowledgment
0.0
0.5
1.0
Overall LHSV
/
1.5 2.0 2.5 3.0 TWO Reactors (h-')
Figure 10. Comparison of observed MCH conversions to TOL in the reactor-membrane system with those estimated from reactor 1 conversions at 1.5 MPa.
37OoC (bath)
-a
Financial support from the Swiss Federal Office of Energy (Bern) is acknowledged. Johnson Matthey plc (Royston, UK) supplied the palladium-silver membranes, and UOP Ltd. (U.K.) supplied the catalyst. Dr. E. J. Newson of the Swiss Paul Scherrer Institute proposed the project.
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Figure 11. Axial temperature profiles for MCH conversion to TOL in reactors 1 and 2 of the reactor-membrane system at 1.5 MPa and an overall LHSV = 1.65 h-I.
those estimated from conversions, x , in the first reactor (Figure 8) using eq 7 at 613 and 643 K, taking the selectivity of the catalyst into consideration. The predicted values are close to those observed in the reactormembrane system but slightly lower at the 643 K temperature level, possibly due to higher axial temperatures in reactor 2 compared to those in reactor 1 as shown in Figure 11. This suggests that MCH conver-
F = moles of MCH in the feed F,,, = total moles of MCH converted to TOL in reactors 1 and 2 Keq= equilibrium constant (kPa3) LHSV = MCH feed, mL h-lheactor volume occupied by the catalyst, mL (h-l) P = moles of MCH in the product P, = permeate Hz pressure (kPa) Pt = total pressure in the reactor (kPa) R = gas constant (8.3144J mol-' K-') T = temperature (K) z = degree of MCH conversion to TOL in a reactor X , = equilibrium conversion of MCH to TOL in a conventional reactor X,, = equilibrium conversion of MCH to TOL in the reactor-membrane system xov= overall conversion of MCH to TOL in reactors 1 and = MCH conversion to TOL in the reactor-membrane system X,,, = MCH conversion to TOL in the two reactors without the membrane vs = increase in MCH conversion to TOL in the reactormembrane system over that without the membrane
X,
Ind. Eng. Chem. Res., Vol. 34, No. 9, 1995 2947
vse= increase in MCH conversion to TOL in the reactormembrane system over the equilibrium limit for a conventional reactor
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Received for review January 23, 1995 Revised manuscript received May 18, 1995 Accepted J u n e 6, 1995@ IE9500679
Abstract published in Advance A C S Abstracts, August 1, 1995. @
