S. Mota, S. Miachon, J.-C. Volta, J.-A. Dalmon∗ IRC-CNRS-Lyon, 2 Avenue Albert Einstein, 69626 Villeurbanne Cedex, France Abstract
A simulation of a packed-bed membrane reactor acting as an oxygen distributor for the selective oxidation of n-butane to maleic anhydride (MA) has been performed by recreating specific reactive atmospheres in a microreactor. In the membranereactor, the oxidation state of the catalyst depends on its position in the bed, leading to an important change in the MA yield.
However, this heterogeneity can be turned to an advantage using a reverse of n-butane flow. Co-promoted catalysts have alsobeen developed to enhance the global performance of the membrane reactor. 2001 Published by Elsevier Science B.V.
Keywords: Maleic anhydride; Catalyst; Reactors 1. Introduction
posed to name this type of membrane reactor an ex-tractor.
For 20 years, membrane catalysis has been receiv- In other applications, the role of the membrane is to ing an increasing interest. The number and variety of dose a reactant that may originate successive reactions.
applications continuously grew and different classifi- As the targeted product is often a product of primary cations have been tentatively proposed. Most of them addition, the regulation of this reactant concentration describe how catalyst and membrane are combined by permeation through the membrane may improve the [1]. However, as suggested in [2,3], it is also possible selectivity. When compared to a conventional reactor, to base the classification on the role of the membrane the same amounts of reactants can be introduced but in the reactor. These two modes of classification are here, one of them is distributed by the membrane along quite complementary and, in the following, specific the catalyst bed. It is therefore proposed to call this names are proposed for the different types of mem- type of membrane reactor a distributor.
The third type of membrane reactor takes advantage In most cases, at least those studied at the beginning of the unique geometry of a membrane, i.e. a perme- of membrane catalysis, the function of the membrane able wall separating two media. If the membrane is was principally to selectively remove, from the reac- also a support for a catalyst, it is therefore possible to tor, a product of an equilibrium-restricted reaction, in feed it from both sides with reactants (for instance gas order to gain yield on conventional reactors. It is pro- from one side, liquid from the other) or to force a reac-tive mixture through the active wall. In the first case,it is possible to favour the contact between the cata- lyst and the reactant that is limiting the performance Corresponding author. Tel.: +33-4-7244-5368; in conventional reactors (e.g. gas in gas–liquid–solid E-mail address: [email protected] (J.-A. Dalmon).
processes, hydrophobic reactant with hydrophilic cata- 0920-5861/01/$ – see front matter 2001 Published by Elsevier Science B.V.
PII: S 0 9 2 0 - 5 8 6 1 ( 0 1 ) 0 0 3 1 0 - 8 S. Mota et al. / Catalysis Today 2398 (2001) 1–8 lyst, etc.). In the second case, the residence time in the 2. Experimental
active pore of reactants and products is controlled byoperating parameters (pressure drop across the mem- brane) and not by diffusion. This may lead to a bet-ter control of activity or selectivity. In the two cases Two catalysts were used, both based on Vanadium the role of the membrane is to favour contact between Phosphorous mixed oxides. The common precursor reactants and catalyst. It is therefore proposed to call of the active solids was VOPO4·2H2O, as obtained this type of membrane reactor a contactor. The first by reacting V2O5 and H3PO4 [11]. Standard solids mode being named an interfacial contactor, the second were obtained by direct reduction of VOPO4·2H2O by isobutanol leading to the vanadyl phosphohemi- In this paper, we report on some aspects of mem- hydrate, VOHPO4·(1/2)H2O. For cobalt-promoted brane reactors of the distributor type, using the selec- solids, reduction was performed in presence of tive oxidation of n-butane to maleic anhydride (MA) Co(C5H7O2), previously dissolved in isobutanol. For as an example. This reaction is until now the only experiments in the membrane reactor, in order to avoid selective oxidation of light alkanes performed at an too large a pressure drop in the catalyst bed, powders industrial scale. Most of the plants use conventional were transformed into pellets of ca. 2 mm size, using fixed-bed reactors [4] in which the flammability of the a lab extrusion machine. Activation of the samples O2/C4 mixtures leads to a very low productivity, owing was made under flowing reactants (O2/C4 = 12) at to the limited concentration of butane in the reactants (1.5%). The resulting low concentration of MA at theoutlet of the reactor also leads to problems during the subsequent separation step. To limit the problems dueto the exothermicity of the reaction, fluidized bed and The MFI membrane was obtained by synthesis of circulating fluidized bed [5] reactors have been devel- zeolite crystals inside the pores of a macroporous host oped but catalysts had to be adapted due to attrition material (pore-plugging method) [10]. The compos- problems [6]. In all cases, the MA yield is limited ite membrane obtained presented several advantages to ca. 60%, the selectivity decreasing rapidly at high when compared to conventional supported zeolite films (less long-range stresses during thermal cy- Several studies have proposed the use of membrane cling, maximum defect not larger than the pore size reactors of distributor type to enhance selectivity to- of the host). The chosen support was a commercial wards primary products in catalytic selective oxida- SCT-USFilter T1-70 tube, made of 3 layers of macro- tions [7–9]. In the present reaction, another advan- porous ␣-alumina (from outer to inner side, respective tage could be expected from the feeding separation thickness: 1500, 40 and 20 ␮m and pore size: 10, of the two reactants by the membrane, which lim- 0.9 and 0.2 ␮m). The precursor solution of the MFI its the flammability problems. This allows the use of zeolite was obtained by mixing silica (Aerosil 380) butane-rich feed, leading to higher MA productions and a template (tetrapropylammonium hydroxide, TPAOH). After a 3-day ageing period, that solution In the present study, special attention has been given was poured in a Teflon-lined autoclave containing to the effects of the oxygen distribution on the cata- the SCT porous ceramic tube. Hydrothermal synthe- lyst characteristics. These characteristics will indeed sis was then performed at 443 K for 3 days, and the directly affect the membrane reactor performance.
membrane was calcined at 773 K under a flow of 5% The membrane reactor combines a tubular porous O2 diluted in N2. Characterisation of the membrane ceramic membrane and a fixed bed catalyst placed in showed it could be considered defect-free (i.e. the the core volume of the ceramic tube. According to transport through the membrane is controlled by the the classification proposed by Tsotsis et al. [1], it is micropores of the MFI structure) and that the sepa- a packed bed membrane reactor (PBMR). Combina- rative layer was a composite MFI-alumina material tion of the two names, i.e. PBMR-distributor, gives a [12]. The permeance of the membrane for oxygen complete description of the reactor and its function.
was ca. 3 × 10−7 mol Pa−1 s−1 m−2 at 670 K (tem- S. Mota et al. / Catalysis Today 2398 (2001) 1–8 perature of membrane reactor catalytic experiments).
5890A, Intersmat IGC) equipped with FID and TCD Other transport properties of the zeolite membrane Different feed configurations were compared. Bu- tane was always introduced in the tube side, oxygen was either co-fed with butane (as in conventional re-actors) or distributed through the membrane (Fig. 1).
They were performed in both the membrane reactor Mixed configurations (some of the oxygen reactant is and the conventional microreactor at 670 K, GHSV co-fed with butane) were also used. The transmem- brane differential pressure was fixed in order to avoid The membrane reactor module was made of a stain- back-permeation of butane to the shell side and to con- less steel shell containing the composite membrane trol the amount of oxygen fed through the membrane.
tube, which was packed with the catalyst (Fig. 1).
Microreactor experiments were performed in order The endings of the membrane tube were enamelled to simulate the local situation of the catalyst along the and equipped with compression fittings of graphite fixed-bed. As a matter of fact, the present membrane seals (Fargraf Cefilac), in order to ensure tightness reactor configuration, as shown in Fig. 1, implies that between the inner (retentate or tube side) and the the O2/C4H10 ratio continuously decreases along the outer (permeate or shell side) compartments. The re- catalyst bed. Butane, introduced at the inlet of the actor temperature was measured by a thermocouple catalyst bed, was progressively consumed, when oxy- moving in a stainless steel tube (1/16 in.) installed in gen was evenly distributed along the bed (the pressure the central axis of the inner compartment. An electric drop due to the catalyst bed being neglected). In or- tape wound around the stainless tube heated the mem- der to characterize the catalyst under fuel rich condi- brane module. The density of whorls was adjusted tions, typical of the situation at the membrane reactor along the reactor length, in order to obtain a temper- inlet, some tests have been performed in a conven- ature profile within ±5 K. The reactor could be fed tional (non-membrane) microreactor with O2/C4 = with n-butane, oxygen and helium (used as dilutent) 0.6. This value has been calculated from oxygen per- with different ratios, from both retentate and permeate meation measurements with the MFI membrane, the inlets. All streams were mass flow (Brooks 5850E) ratio of the total amounts of O2 (distributed) and C4 P, the pressure difference across the introduced in the membrane reactor being 12, close membrane was regulated by an automatic valve (Käm- to the value of conventional fixed-bed industrial pro- mer 800377) connected to a differential manometer cesses (oxygen-rich mixture). For comparison, other (Keller DP232), whereas the pressure at the retentate experiments were carried out using the microreactor outlet is measured by a manometer (Keller PAA23).
under O2/C4 = 12 (standard ratio) and 20 (simulation Permeate and retentate outlet flow rates were mea- of the membrane reactor outlet). Results were quite sured with film flowmeters. Reaction products were similar for these two fuel lean conditions. Catalysts analyzed on-line using a gas chromatograph (HP were characterized before and after testing.
Fig. 1. Schematic of the membrane reactor (segregated feed).
S. Mota et al. / Catalysis Today 2398 (2001) 1–8 Fig. 2. 31P (SEM) NMR spectra of the catalysts: (a) fresh standard catalyst; (b) standard catalyst after testing in the microreactor underfuel rich conditions; (c) fresh Co-promoted catalyst; (d) Co-promoted catalyst after testing in the microreactor under fuel rich conditions.
3. Results
conditions (simulation of the inlet of the membranereactor).
Characterizations of the solids before and after test- ing in the microreactor (20 h on stream) have been performed using different techniques [13]. After ac- Fig. 3 compares the performance (MA yield) of the tivation (fresh catalyst), the solids showed a specific standard catalyst under conventional (O2/C4 = 12) area close to 20 m2 g−1 and XRD analysis disclosed and fuel rich (O2/C4 = 0.6) conditions. In the last in all cases the presence of vanadyl pyrophosphate case, a rapid deactivation occurred, and only butenes, (VO)2P2O7. Fig. 2 shows the 31P (SEM) NMR spec- via oxidative dehydrogenation of butane, were formed.
tra of the standard and Co-modified catalysts before Fig. 4 illustrates the reversibility of the catalyst prop- and after testing in the microreactor under fuel rich erties when cycling the operating conditions from fuel S. Mota et al. / Catalysis Today 2398 (2001) 1–8 Fig. 5. Comparison of the standard (᭡) and Co-promoted (᭿) cat-alysts under fuel rich conditions (T = 670 K, GHSV = 3000 h−1).
1. Co-feed: oxygen and n-butane are mixed and fed the tube side. Permeate inlet and outlet are closed Fig. 3. Performance (MA yield) of the standard catalyst in the (this mimics the conventional reactor).
microreactor, T = 670 K, GHSV = 3000 h−1: (᭿) O2/C4 = 12, 2. Totally segregated feed (n-butane to the tube side, conventional fuel lean conditions; (᭜) O2/C4 = 0.6, fuel rich 3. Mixed feed (ca. 20% of total oxygen entering the catalyst bed is co-fed with n-butane).
lean to fuel rich atmospheres. Fig. 5 shows the re- 4. Similar to configuration 3, but with reversing of the spective behaviors of the standard and modified cata- feed containing n-butane. The performance mea- lysts under fuel rich conditions. Even though the stan- sured 15 min after the reversal is clearly above that dard solid rapidly deactivated, the Co-promoted cata- obtained before 30 min after reversal, the MA yield Note that all the membrane experiments presented 4. Discussion
here have been performed with the standard catalyst.
Fig. 6 represents schematics of the different feeding modes of the membrane reactor. The correspondingcatalytic performances are given in Table 1. All exper- If the standard catalyst performance was stable iments were performed with a O2/C4 ratio (introduced when operated under conventional fuel lean conditions in the catalyst bed) between 8 and 9, close to the con- (O2/C4 = 12), the MA yield rapidly drooped to zero ventional fuel lean ratio. These different modes are: when placed under fuel rich atmosphere (Fig. 3). At Fig. 4. Reversibility of the performance of the standard catalyst in the microreactor.
S. Mota et al. / Catalysis Today 2398 (2001) 1–8 Fig. 6. Schematics of the different feed configurations of the membrane reactor.
this state, only butanes were formed via C4H10 After modification by cobalt addition, the perfor- mance under fuel rich conditions (O2/C4 = 0.6) was The 31P (SEM) NMR spectrum of the standard cat- much better than that of the non-promoted system alyst before testing shows (Fig. 2) that the fresh cat- (Fig. 5). Even if the performance was lower than that alyst contained both V4+ (near 2500 ppm) and V5+ obtained under fuel lean atmosphere, some MA was species (near 0 ppm). The very same NMR spectrum produced at steady state. The NMR spectrum after (not shown here) was observed after testing under a testing (Fig. 2) showed the presence of a small peak O2/C4 ratio of 12, in good keeping with the stability close to 0 ppm, indicating that, even under reducing of the catalytic performance. This was no more the (fuel-rich) conditions, V5+ species were still present case after catalytic testing under fuel rich conditions in the Co-modified solid, at least enough to lead to se- (O2/C4 = 0.6), as only V4+ were then visible (Fig. 2).
lective oxidation. Note also that this peak was much On the basis of similar experiments, it has been pro- larger in the fresh promoted catalyst than in the fresh posed that the active site for selective oxidation in- standard one. This also suggests that cobalt favors the volves both V4+ and V5+ species [13]. In the pres- presence of V5+ species. It has also been suggested ence of mainly V4+ species, only ODH takes place.
that Co is involved in the V4+/V5+ redox equilibrium When there is a large excess of V5+, total oxidation existing during selective oxidation of butane to MA The surface state was quite reversible however, XPS analyses also supported those V4+/V5+ evolu- since activity rapidly changed and recovered previous tions under fuel rich conditions for both solids [13,14].
Let us first consider the reactor with co-fed reactants (Fig. 6). The performance (Table 1) was close to that Catalytic results of the membrane reactor as a function of the feed observed with the microreactor (Fig. 3, O close values of W/F). This suggests that the pellets packed in the membrane tube work similarly to the powder in the microreactor. The co-feed membrane reactor can therefore be considered as representative For totally segregated feeds, the MA yield was about 25% lower than in the previous case. This is not sur- a O2/C4 is the ratio of reactants entering the catalyst bed. XC4 prising, as the first part of the catalyst bed was proba- is the butane conversion (%), SMA and YMA the maleic anhydrideselectivity and yield expressed in %.
bly deactivated, due to the reductive atmosphere pre- S. Mota et al. / Catalysis Today 2398 (2001) 1–8 vailing in this zone. Such a deactivation was also ob- lyst bed and a reversal of the C4 flow may be an inter- served in the microreactor under similar fuel rich con- esting subject for further investigations. It is clear that the above-mentioned observations deserve more ex- This led us to introduce some oxygen diluted in the perimental studies, especially about the transient MA butane feed, as was proposed by Mallada et al. [10].
production. The data reported here give the situation In the case of such a mixed feed, the performance 15 min after the reversal, and a peak in the MA produc- of the membrane reactor was similar to that of the tion must occur during each period, which timing will conventional one. A simple calculation of the O2/C4 be a key parameter for an optimization of the process.
ratio in the first part of the reactor (1/10th of the total Further studies are currently under way to better de- length of the packed bed) gives a value in the range scribe this type of membrane reactor, including pos- 2–3, which should be too low to ensure a high activity sible heat transfer effects. Moreover, cobalt-promoted towards MA. Considering however that a significant catalysts will be used to enhance the global perfor- part of the packed bed was placed in this inadequate reactive atmosphere, the global performance was stillhigh, and there is certainly room for an optimizationof the feed conditions.
5. Conclusion
Even if the performance of the mixed fed reactor was better than that of the totally segregated configu- The simulation in a microreactor of the catalyst bed ration, the oxidation state of the catalyst bed was cer- of the PBMR-distributor showed an important hetero- tainly heterogeneous, the first part being too reduced, geneity of the catalyst oxidation state, depending on when the end of the bed was probably overoxidized.
its axial position in the reactor. The active catalyst re- This heterogeneity can however be turned to an advan- quiring an optimal V4+/V5+ ratio, the inlet of the bed tage by reversing the flow in the inner volume, leading presented a poor performance, as it was excessively to a transient MA yield clearly higher than that of the reduced to V4+ by the butane flow. This can be partly conventional reactor (Table 1). This is due to the fact solved by diluting some of the O2 reactant in the bu- that, just after reversal, butane first flows through the tane feed. For these conditions, the global performance previously overoxidized catalyst, leading in this zone (MA yield) of the membrane reactor was similar to to a high MA yield. This however is only a transient that obtained using the conventional cofeed configu- phenomenon, as this catalyst zone was progressively ration. As has been shown in [10], it is then possible reduced by butane, driving back the situation to the to take advantage of the segregated feed, which lim- previous one. However, during this interval, the end its the flammability problems, and operate with higher of the bed was oxidized by the high O2/C4 ratio pre- butane concentrations than those used in conventional vailing then in this zone. It is therefore possible, us- ing a new reversal of the C4 flow, to observe a new Another issue of the present study is the increase of exaltation of the MA yield. This catalyst flexibility is the MA yield observed when reversing the butane flow also illustrated in Fig. 4, when cycling the microreac- in the membrane reactor. During a transient period the tor between fuel rich and lean conditions.
performance was higher than that observed using a To schematise, the catalyst bed of the reversible conventional feed. This was due to the heterogeneity membrane reactor can be divided into two zones, the of the packed bed, which, at steady state, presented an first being reduced by butane (MA production), while increase of oxidation state from inlet to outlet. When the second is reoxidized. In this way this membrane re- reversed, the butane first flew in the oxidized zone actor looks like the CFB reactor developed by DuPont leading to the observed increase of the MA yield, while [3]. In both cases, reduction and reoxidation of the cat- oxygen distributed through the membrane reoxidized alyst are separated, but they occur here without mov- the end of the bed. Further studies are currently under way to better describe this type of membrane reactor Let us underline that these are preliminary observa- that couples distribution of one reactant and sequential tions, but this type of reactor, combining a membrane reverse flow of the other. It may combine two of the distributing oxygen in a continuous way to the cata- main interests of reverse flow operation, i.e. better heat S. Mota et al. / Catalysis Today 2398 (2001) 1–8 management and exploitation of the catalyst dynamic [3] A.G. Dixon, Catalysis 14 (1999) 40.
[4] F. Cavani, F. Trifiro, Chemtech (1994).
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sphere they create. In this study, a cobalt-promoted solid adapted to reducing atmospheres has been de- [7] M.P. Harold, C. Lee, A.J. Burggraaf, K. Keizer, V.T. Zaspalis, R.S.A. de Lange, MRS Bul, April 1994, p. 34.
[8] A. Pantazidis, J.-A. Dalmon, C. Mirodatos, Catal. Today 25 [9] C. Tellez, M. Menendez, J. Santamaria, AIChE J. 43 (1997) Acknowledgements
[10] R. Mallada, M. Menendez, J. Santamaria, Catal. Today 56 This work has been carried out with financial sup- [11] M.T. Sananés-Schulz, F. Ben Abdelouahab, G.J. Hutchings, port from the European Commission, Project BRPR J.C. Volta, J. Catal. 163 (1996) 346.
[12] A. Giroir-Fendler, J. Peureux, H. Mozzanega, J.-A. Dalmon, Stud. Surf. Sci. Catal. 101A (1996) 127.
[13] S. Mota, M. Abon, J.C. Volta, J.-A. Dalmon, J. Catal. 193 References
[14] S. Mota, J.C. Volta, G. Vorbeck, J.A. Dalmon, J. Catal. 193 [1] T.T. Tsotsis, R.G. Minet, A.M. Champagnie, P.K.T. Liu, in: Computer-Aided Design of Catalysts, Marcel Dekker, New [15] Y.S. Matros, G.A. Bunimovich, Catal. Rev. Sci. Eng. 38 (1) Handbook of Heterogeneous Catalysis, VCH, 1997, Chapter9.3, p. 1387.


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