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Technical and economical feasibility of zeolite NaA membrane-based reactors in liquid-phase etherification reactions
Chemical Engineering and Processing 48 (2009) 1072–1079
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Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep
Technical and economical feasibility of zeolite NaA membrane-based reactors in liquid-phase etherification reactions Marc Pera-Titus ∗ , Joan Llorens, Fidel Cunill Chemical Engineering Department, University of Barcelona, 08028 Barcelona, Spain
a r t i c l e
i n f o
Article history: Received 2 October 2008 Received in revised form 8 December 2008 Accepted 28 February 2009 Available online 14 March 2009 Keywords: Pervaporation Zeolite NaA Membrane reactor Di-n-pentylether Dehydration
a b s t r a c t This paper describes the pervaporation performance of film-like zeolite NaA membranes synthesized in our laboratory by template-free seeded hydrothermal towards the dehydration of n-pentanol/water and n-pentanol/water/DNPE mixtures (DNPE = di-n-pentyl ether, C10 H22 O). DNPE is a linear symmetric ether that can be used as blending additive in reformulated diesel fuels and that can be produced from C4 feedstocks via n-pentanol, obtained in its turn by selective hydroformylation of linear butenes. The membranes showed selectivities up to 3000 in the dehydration of n-pentanol/water binary mixtures, and water/DNPE and n-pentanol/DNPE selectivities approaching infinity (i.e. DNPE does not pervaporate) in the dehydration of n-pentanol/water/DNPE ternary mixtures. On the basis of the separation results presented in this work, and using experimental reaction kinetics data, we discuss the feasibility of zeolite NaA membrane reactors to carry out the liquid-phase etherification reaction of n-pentanol to di-n-pentyl ether (DNPE) catalyzed by ion-exchange sulfonated resins. These catalysts suffer from strong deactivation in the progress of the reaction due to the generation of water. © 2009 Elsevier B.V. All rights reserved.
1. Introduction Zeolite membranes have been largely advocated as alternative candidates to polymeric membranes for the separation of species relying on molecular size, adsorption affinity and/or surface diffusion differences [1]. In the special case of liquid mixture separation by pervaporation (PV), separation mainly occurs on the basis of adsorption differences [2]. This is the case for instance of solvent dehydration by zeolite NaA membranes, where water adsorbs preferentially on the zeolite material due to the low Si/Al ratios A-type zeolites [1,4–8] (approaching to 1, the minimum allowed by the Loewenstein rule [3]). The first open report on large-scale alcohol dehydration using A-type zeolite membranes dates back to 2001 [9], when the Japanese Mitsui Engineering & Shipbuilding Co. started up in Japan a plant for isopropanol dewatering by vapour permeation (VP) with a production of 20,000 L/day. By now, this company and its subsidiary the Nano-Research Institute Inc. (XNRI), as well as the European alliance between Smart (UK) and Inocermic (Germany), supply commercially LTA zeolite membranes for PV and VP dehydration separations. New VP demonstration plants have been recently installed in Brazil (3000 L/day) and India (30,000 L/day) for bioethanol dehydration.
∗ Corresponding author. Present address: Institut de Recherches sur la Catalyse et l’Environnement de Lyon (IRCELYON), UMR 5256-CNRS/University of Lyon, France. E-mail address: [email protected] (M. Pera-Titus). 0255-2701/$ – see front matter © 2009 Elsevier B.V. All rights reserved. doi:10.1016/j.cep.2009.02.006
Three are the main drawbacks when dealing with zeolite NaA membranes compared to other zeolite membranes: (1) this zeolite suffers from dealumination at strong acidic conditions, which in practice limits its application to environments with pH > 6; (2) the presence of a large amount of Al in the framework structure acts as a promoter for membrane shrinkage; and (3) hydrothermal synthesis is in practice difficult to scale up due to formation of large intercrystalline defects in the active layer. This third shortcoming can be partially overcome, for instance, by seeding the support surface prior to hydrothermal synthesis [1,4], or by refreshing the synthesis solution in direct contact with the support. Gel renewal can be accomplished either by the action of a centrifugal force field [4], or by pulse- or continuous renewal of the synthesis solution [5,6]. One of the advantages of pervaporation compared to gas separation is that, despite the presence of a reduced amount of defects in the zeolite membrane structure, these can still show selectivity due to strong water adsorption. In addition, as we have suggested in a previous study [10], high compressive tensions due to negative capillary forces might still allow discrimination between water and organic species in small-sized mesopores (i.e. <5 nm), contributing positively to the membrane separation performance. Taking into account all the above stated aspects, zeolite NaA membranes appear to be promising candidates for in situ water removal in extractor-type catalytic membrane reactor (CMR) applications. This is the case for instance of etherification reactions carried out by alcohol dehydration and catalyzed by ion-exchange acid resins, usually suffering from severe catalyst deactivation due
M. Pera-Titus et al. / Chemical Engineering and Processing 48 (2009) 1072–1079
to strong water adsorption on the resin active sites. The use of a CMR with selective water extraction from the reaction system might not only improve the catalytic activity, but also help circumventing the high-energy costs ascribed to reactive distillation. Note that, as a rule, membranes might help saving up to 85% of the energy demands ascribed to distillation [11], involving a reduction of 30–50% of the energy costs [12]. Several examples of this concept applied to the synthesis of MTBE and ETBE catalyzed by ion-exchange resins have been reported in the literature [13,14]. Acid-resistant mordenite (MOR), merlionite (MER) and phillipsite (PHI) membranes [15,16] have been applied as well to esterification reactions catalyzed either by mineral or solid catalysts for shifting the equilibrium position. As an example, in this paper we focus our attention on the synthesis of di-n-pentyl ether (DNPE) from bimolecular dehydration of n-pentanol (SN 2) catalyzed by acid ion-exchange resins. DNPE is a linear and symmetric ether of molecular formula C10 H22 O that, among other long-chain oxygenates, can be used as blending additive in reformulated diesels to increase their cetane number, improve their cold flux properties, and reduce tail-pipe emissions (especially polyaromatics and sulfur) [17,18]. This ether also shows industrial interest, since n-pentanol can be produced industrially from valorization of C4 feedstocks by selective hydroformylation of linear butenes (oxo process) [19]. On the basis of experimental n-pentanol/water/DNPE dehydration data, as well as reaction rate data of DNPE formation, we discuss how the coupling of a zeolite NaA membrane with an ion-exchange resin in an extractortype CMR configuration could enhance DNPE synthesis from n-pentanol dehydration. We also discuss the technico-economical feasibility of DNPE mass production in a zeolite NaA catalytic membrane reactor (ZCMR) compared to a reactive distillation unit. 2. Experimental The membranes used in this study consisted of a continuous active film (thickness 30–35 m) of randomly oriented zeolite NaA crystals located on the inner surface of porous ␣-alumina tubular supports (active surface 12.6 cm2 , effective length 5 cm) grown by seeded hydrothermal synthesis. The main characteristics of the membranes are summarized in Table 1. The details dealing with the membrane synthesis can be found in Refs. [4,5]. All the membranes displayed low N2 permeance (<10−7 mol m−2 s−1 Pa−1 ) and good PV performance for the separation of 92:8 wt.% ethanol/water mixtures at 323 K, feed pressure in the range 101–303 kPa, and permeate pressure <0.2 mbar. The membranes showed an extremely low amount of defects or intercrystalline porosity (εinter < 10−6 ), as inferred using the characterization method reported in Ref. [10]. A set of steady-state PV experiments with a set of C1–C5 primary alcohol/water mixtures and n-pentanol/water/DNPE mixtures were performed in a laboratory-scale PV test bench. The details dealing
Table 1 Water/ethanol selectivity (˛w/E ), total flux (NT ), zeolite film thickness (ZA ) and defective structure (d¯ C : mean pore size; εinter : intercrystalline porosity) of the zeolite NaA membranes used in this study. Membrane
ZA1 ZA2 ZA3 a
˛W/E
185 294 1084
NT [kg m−2 h−1 ]
0.58 0.36 0.47
ZA [m]a
35 30 35
Large defectsb d¯ C [nm]
εinter × 107
589 ± 8 397 ± 4 490 ± 4
5.6 ± 0.6 2.9 ± 0.2 1.8 ± 0.3
Determined by SEM. Determined from the evolution of the total pervaporated flux with the feed pressure according to the characterization method described in Ref. [8]. b
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Fig. 1. Evolution of total flux and water/alcohol and water/mixture selectivities as a function of the chain length of the primary alcohol for membrane ZA2. Conditions: xw = 6.6–8.0 mol%, T = 323 K, Po = 303 kPa, Pv = 1 mbar, and xDNPE = 19.4–20.4 mol% (ternary mixture).
the experimental setup and protocols to carry out the PV tests are described elsewhere [4]. Methanol, ethanol and 1-propanol (99.5% purity) were supplied by Panreac (Spain), 1-butanol (99.7% purity) by Romil (England) and 1-pentanol (99.0% purity) by Aldrich (Germany). DNPE (98.0% purity) was synthesized in our laboratory in a batch reactor from dehydration of n-pentanol. Care was taken in the experiments involving BuOH and PeOH alcohols and DNPE to ensure complete miscibility of their liquid mixtures with water. The physical properties of the alcohols and DNPE were obtained from Refs. [20,21]. The density of n-pentanol and DNPE are, respectively, 0.815 and 0.787 g cm−3 , while their corresponding normal boiling points are 141.2 and 183.7 ◦ C. The water/alcohol selectivity was calculated as the quotient between the ratios of the molar fractions of water and alcohol at the permeate and feed sides, i.e. ˛w/A = (Yw /YA )/(Xw /XA ). In the case of n-pentanol/water/DNPE mixtures, the selectivity was calculated as the ratio of the molar fractions of water and mixture (mixture = PeOH + DNPE) at the permeate and feed sides. The composition in the feed and permeate was analyzed in a HP6980A GC equipped with a 50 cm × 0.2 mm × 0.5 mm methyl silicone capillary column and a TCD detector. The final values correspond to the mean of at least 2–6 steady-state measurements during 1–3 h taken after 2–4 h of stabilization. The reproducibility of the experiments was optimal, with standard deviations of the total flux and selectivity lower than 5% and 10%, respectively. 3. Results Fig. 1 shows the evolution of the water/alcohol selectivity and total flux with the chain length of the primary alcohol in the PV of alcohol/water binary mixtures within membrane ZA2 at 323 K and for a ∼7 mol% water composition in the feed. As can be seen, while the total flux seems to be practically independent on the nature of the alcohol in the mixture within the limits of the experimental error (except for methanol), the water/alcohol selectivity appears to be strongly affected. Fig. 1 also shows that, in the presence of DNPE (xDNPE = 19.4–20.4 mol%), the water/mixture selectivity appears to be promoted when compared to the value obtained for the PeOH/water binary system at the same temperature and water composition. However, the total flux appears to be unaffected by the presence of DNPE. Fig. 2 shows the effect of the feed composition on the total flux and water/alcohol selectivity at 323 K for EtOH/water and
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Fig. 2. Evolution of total flux (on top) and water/alcohol (or water/mixture) selectivity (on bottom) with water feed composition for membranes ZA1 and ZA2. Conditions: T = 323 K, Po = 303 kPa, Pv = 1 mbar, and xDNPE = 19.4–20.4 mol% (ternary mixture). The lines are a guide to the eye.
PeOH/water mixtures at low feed water compositions (2–30 mol%) for membranes ZA1 and ZA2. As can be seen, the total flux increases practically linearly with the water composition in the range 7.5–18 mol% with a similar slope for both mixtures, while the selectivity remains practically invariable, except for membrane ZA1, showing a slight increase. For membrane ZA2, the water/PeOH selectivities are one order of magnitude higher than those obtained for the EtOH/water system. Fig. 3 plots the effect of temperature on the water flux and water/alcohol selectivity for membrane ZA2 at low feed water compositions (2–30 mol%) for the dehydration of EtOH/water and PeOH/water binary mixtures and PeOH/water/DNPE ternary mixtures. As can be seen, water flux shows an Arrhenius trend for all the mixtures, with effective activation energies for water of 34 kJ mol−1 for EtOH/water mixtures and 20–24 kJ mol−1 for PeOH/water and PeOH/water/DNPE mixtures. For the latter two mixtures, regardless of the presence of DNPE, the total flux seems to depend only on feed composition. Furthermore, the water/EtOH and water/PeOH selectivities are almost constant with temperature in the range 303–410 K, the latter being one order of magnitude higher than the former, with values up to 3000. Moreover, water/PeOH selectivities are higher for higher water feed compositions (range 22.9–27.9 mol%), the values being comparable to those obtained for water/mixture selectivities in the presence of ∼20%mol DNPE in the liquid feed. However, in the case of ternary mixtures, the water/mixture selectivities show a slight increase with temperature.
Fig. 3. Evolution of water flux (on top) and water/alcohol selectivity or water/mixture selectivity (on bottom) with temperature for membrane ZA2. Conditions: Po = 303 kPa, Pv = 0.2–1 mbar, and xDNPE = 19.4–20.4 mol% (ternary mixtures). The lines for selectivity are a guide to the eye, while those for water fluxes correspond to Arrhenius-type fittings.
4. Discussion 4.1. Equation for water flux within zeolite NaA membranes The tends plotted in Figs. 1–3 reflect that, as expected, hydrophobic species are hardly separated by zeolite NaA membranes due to the inherent hydrophilic character of this zeolite. Moreover, the fact that water/mixture selectivities in the presence of DNPE are higher than PeOH/water selectivities obtained for the binary system at the same temperature and comparable water composition should be ascribed to the stronger hydrophobic character of the organic phase in the presence of DNPE, impelling therefore its separation. The correlations plotted in Figs. 1–3 allow the derivation of an equation for water flux from an adsorption–diffusion model formulated on the basis of the Maxwell–Stefan (MS) theory using the approach outlined in a previous study [6]. At low water feed activities, aw , and sufficiently low vacuum permeate pressure, the water flux can be expressed by the following equation (pseudo solutiondiffusion model): S Nw =
S S qM,w Mw DSw Kw o aw Pw ZA
(1)
Eq. (1) can be rewritten to Eq. (2) taking into account the explicit dependence of the MS surface diffusivity, the adsorption constant and the saturation vapour pressure of water with temperature (Eq.
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Fig. 5. Scheme of a zeolite NaA catalytic membrane reactor (extractor-type) with the catalyst packed inside the membrane tubes.
Fig. 4. Evolution of experimental and fitted water fluxes for PeOH/water and PeOH/water/DNPE mixtures in the water composition and temperature ranges 0.05–0.30 mol% and 303–403 K, respectively. The DNPE composition in the ternary mixtures was 19.4–20.4 mol%.
(2)): S = Nw
S (0) exp(B) S qM,w Mw DSw (0)Kw exp ZA
S Ew,eff
RT
(2)
aw
S In Eq. (2), the effective activation energy of water, Ew,eff , includes three main contributions (Eq. (3)): S Ew,eff
=
S Ew
◦ + Hw
+ RA
(3)
S is the activation energy for water surface diffusion, H ◦ where Ew w is the adsorption enthalpy of water, and parameter A belongs to the Antoine equation accounting for the evolution of saturation vapour pressure of water with temperature (A = 5300 K−1 in the temperature range 293–363 K). Eq. (2) has been fitted to the experimental water flux data in PeOH/water and PeOH/water/DNPE mixtures plotted in Figs. 2 and 3 to derive an expression useful for modelling purposes. In the fitting process, the effective activation energy of water and the preexponential factor in Eq. (2) have been adjusted by the least-square method using a Levenberg–Marquard optimization algorithm. The activity coefficients of water in the mixtures have been estimated by the UNIFAC method [22]. The final expression obtained is described by the following equation:
S = (1.49 ± 0.09) exp −(2700 ± 220) Nw
1 T
−
1 353
[kg m−2 h−1 ]
aw (4)
Fig. 4 shows the good agreement between the experimental and fitted water fluxes, the adjusted parameters lacking of significant correlation effects. 4.2. Synthesis of DNPE in a zeolite membrane reactor The strong dehydration performance of the zeolite NaA membranes synthesized in this work is especially remarkable for the dehydration of the PeOH/water/DNPE mixtures, since PV of DNPE within these membranes is practically negligible. In view of this result, zeolite NaA membranes appear to be promising candidates for DNPE synthesis in zeolite membrane reactors in the absence of acid conditions. On the basis of the results obtained for PeOH/water/DNPE separation in zeolite NaA membranes, we present here the simulation of a multitubular ZCMR to carry out the liquid-phase etherification reaction of n-pentanol to DNPE. For comparison purposes, we have also performed some simulations
on a conventional fixed-bed reactor to illustrate the enhancement of n-pentanol conversion that could be achieved when coupling the catalyst with selective in situ water removal. As an example, we have chosen the recently developed thermally stable S/DVB Amberlyst 70 resin, offering high activity and selectivity to DNPE (the main byproducts are pentenes generated by bimolecular elimination, E2). The kinetics of the main reaction can be well described by a model stemming from an Eley–Rideal (ER) mechanism wherein DNPE remains practically adsorbed on the catalyst surface, the number of free active sites is negligible, and the surface reaction is the rate limiting step (Eq. (5)) [23,24]. rDNPE =
k(a2PeOH − aw aDNPE /K)
(5)
1/2
aPeOH (1 + Kw aw )
with K = 8.9229 exp
778.69
T
Kw = exp 1.46 − 6615
1
k = exp 2.808 − 11595
T
1 T
(6)
−
1 438
−
1 438
(7)
[mol h−1 kgcat −1 ]
(8)
In Eq. (5), ai is the activity of species i, the activity coefficients having been estimated by the UNIFAC method. The term (1 + Kw aw 1/2 ) in the denominator of Eq. (5) accounts for catalyst deactivation due to water adsorption. It is noteworthy that, according to the expression of the equilibrium constant (Eq. (6)), the reaction is only slightly exothermic, with a reaction enthalpy only about −6.5 kJ mol−1 [25]). The multitubular ZCMR modelled in this study consists of a catalyst packed bed in the lumen of the membrane tubes, while the vapour permeate side of the membrane is kept under vacuum (see Fig. 5). Assuming a quasi-isothermal regime, the ZCMR can be modelled by a microscopic mass balance for each species in the lumen and in the permeate sides of the tubes. For the sake of simplicity, plug-flow and perfect mixing regimes have been assumed, respectively, to describe the hydrodynamics of both zones. The pressure drop along the axial position inside the membrane tubes can be modelled by the Ergun Equation. The set of Eqs. (9)–(11) is obtained at steady state: • Microscopic mass balance (i = 1,N) −
∂(wR xi ) − Ni am + S (1 − εb )ri = 0 Ab ∂z
(9)
• Permeation (i = 1,N) −Ni am =
∂(wP yi ) Ab ∂z
(10)
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Table 2 Input data for reactor modelling. Parameter
Value
xPeOH,in xw,in = xDNPE,in T [K] Po,in [kPa] Pv [mbar] Residence time [kgcat h kmol−1 ] PeOH mass flow [tm year−1 ] Db [mm] am [m2 m−3 ] uo range (lumen) [m min−1 ] εb Catalyst: Amberlyst 70 [23] Density [kg m−3 ] Apparent density [kg m−3 ] Surface area [m2 g−1 ] Dp [m]
1.00 0.00 443 500 <5 53 55,000 20 300 0.20–0.70 35% 1542 647a 176a 549b
a
Values obtained in swollen state (ISEC method). This particle size ensures the absence of significant internal mass (IMT) contribution to the reaction rate. b
• Pressure drop (Ergun equation) ∂Po (1 − εb )2 L uo (1 − εb ) L u2o = 150 + 1.75 3 2 Dp ∂z Dp εb ε3b Boundary conditions: z = 0 → xi = xi,in , yi = yi,in ,
(11)
Po = Po,in
It is interesting to note that the set of Eqs. (9)–(11) for a ZCMR can be reduced to a fixed-bed reactor if Ni → 0 and to a tubular PV unit if ri → 0 for ∀i. The water permeation and reaction kinetics have been taken, respectively, from Eqs. (4) and (5). A value of 3000 has been chosen for the water/mixture selectivity of the membrane according to the results plotted in Figs. 2 and 3. The reactor model has been solved numerically through discretization using finite differences. The number of intervals (130) has been chosen to avoid any dependence of the simulation results on the discretization. As an example, the reactor has been simulated for a PeOH inlet flow of 55,000 tm/year. The input data used for the modelling are summarized in Table 2. Fig. 6 compares the simulation results obtained for n-pentanol conversion, reaction rate and water concentration inside the tubes along the dimensionless axial position, = z/L, in the case of a ZCMR and a fixed-bed reactor using the input data from Table 2. As can be seen, at 433 K and for a residence time of 53 kgcat h kmolPeOH −1 , the n-pentanol conversion in a ZCMR would reach a value up to 64%, while a value of 35% might only be obtained in a fixed-bed reactor. As can be deduced from Fig. 6 (on top), this difference would be ascribed to the higher reaction rates achieved in the former case due to selective water removal from the reaction mixture. Accordingly, a fairly stable and extremely low water composition (<0.8 mol%) would be obtained along the axial position in a ZCMR, while the water composition would increase progressively in a fixed-bed reactor. The output data obtained from our simulations are summarized in Table 3. According to our calculations, about 27,000 membrane tubes of 1 m length and 20 mm i.d. would be required. Table 3 also reflects the suitability of the plug-flow regime to describe the hydrodynamics inside the membrane tubes, since the conditions Db /Dp > 30 and L/Dp > 200 are fulfilled [26]. The surface velocity, uo , would be around 0.20 m min−1 , ensuring negligible external mass transfer effects on the reaction rate. The pressure drop along the reactor would be ∼7.3 kPa, translating into ∼13 kW pumping requirements for the given tube effective length. Moreover, ca. 98% of the water generated during the reaction would be removed in the permeate stream of a ZCMR significant reactant loss (<1%).
Fig. 6. Evolution of PeOH conversion and reaction rate (on top) and water molar fraction (on bottom) with the axial position inside the reactor for a fixed-bed (thin lines) and a ZCMR (thick lines).
4.3. ZCMR-based process for DNPE production: techno-economical analysis On the guidance of the simulation results obtained for DNPE production (35,000 tm/year) from n-pentanol dehydration in a ZCMR (64% reactor conversion), a process for its production can be proposed according to the process flow scheme illustrated in Fig. 7. In addition, Table 4 summarizes the values of the main parameters in the relevant streams. In principle, only one distillation column would be required at the outlet of the ZCMR to purify the DNPE, the n-pentanol and low amounts of water obtained in the distillate being recirculated to the reactor after condensation. The system would be also provided with a condenser operating at vacuum Table 3 Output data obtained from the simulations in the ZCMR. Variable
Value
Effective length [m] Number of tubes of 1 m length [–] XPeOH [%] Water removed [%] PeOH removed by membrane [%]
∼16 ∼27,000 64 >98 ∼1
Hydrodynamics Db /Dp L/Dp Pressure drop [kPa] uo [m min−1 ]
36 > 30 280 > 200 ∼7.3 kPa ∼0.20
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Fig. 7. DNPE process flow scheme based on a catalytic zeolite membrane reactor. Nomenclature: C, condenser; DC, distillation column; E, expansion system; MR, membrane reactor; HE, heat exchanger; P, liquid pump; R, reboiler; VP, vacuum pump.
(pressure ∼1 kPa) equipped with a primary vacuum pumping unit to remove the water extracted by the membrane. The condensed water would be further removed by pumping to atmospheric pressure. The distillation column has been modelled by the Smith– Brinkley method extended to non-ideal mixtures [27]. According to our simulations, the system PeOH/water/DNPE is not expected to form any azeotrope neither in the distillate nor in the residual streams in the conditions indicated in Table 4. The analysis of the economic feasibility of the technological solution here proposed has been carried out in a similar way as that proposed by Criscuoli et al. [28] for costs analysis in the case of a water gas shift membrane reactor. In terms of capital cost (CAPEX), our estimations reflect that, for 35,000 tm DNPE/year production, about 30% of the investment costs of the plant would be ascribed to the membrane reactor. In this estimate, a unit cost of 5000 US$/m2 for the zeolite membrane tubes has been considered, the zeolite Table 4 Properties of the streams in Fig. 7. Stream
T [K]
P [kPa]
F [kmol/h]
xPeOH
xH2 O
xDNPE
Phase
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15
293.1 357.9 433.1 433.1 283.3 283.3 433.1 443.1 469.3 462.7 462.7 462.7 521.0 521.0 293.1
101 535 505 1 1 101 498 468 495 465 515 535 545 495 101
42.1 68.2 68.2a 21.4 21.4 21.4 47.2 47.2 139.2 139.2 139.2 26.1 122.7 21.1 21.1
1.0000 0.9926 0.9926 0.0525 0.0525 0.0525 0.5428 0.5428 0.9808 0.9808 0.9808 0.9808 0.0011 0.0011 0.0011
0.0000 0.0050 0.0050 0.9475 0.9475 0.9475 0.0072 0.0072 0.0131 0.0131 0.0131 0.0131 0.0000 0.0000 0.0000
0.0000 0.0024 0.0024 0.0000 0.0000 0.0000 0.4524 0.4524 0.0062 0.0062 0.0062 0.0062 0.9989 0.9989 0.9989
L L L V L L L L V L L L L L L
a Value corresponding to 55,000 tm/year n-pentanol included in Table 2 (64% reactor conversion).
layer being assumed to account for 30% of the zeolite NaA membrane cost. The cost of the equipments would be about 9 M$, 7.7 M$ being ascribed to the ZCMR, while the overall capital costs of the plant (equipment + installation) would reach 29 M$. The capital cost of the distillation column, including the condenser and reboiler, as well as the capital cost of the vacuum condenser, heat exchangers, vacuum pumps and liquid pumps have been estimated and upscaled using the Williams method, the costs having been capitalized to present monetary units using Marshall & Swift correction factors [27]. The analysis of the operational costs (OPEX) of the plant following conventional standards reflects a year cost about 80 M$ for 35,000 tm DNPE/year production (see Table 5). The higher contribution to the OPEX would be ascribed to raw materials (n-pentanol and catalyst replacement), membrane replacement and refrigeration water, accounting, respectively, for 33%, 32% and 19% of the overall costs. These costs would translate into a unit cost of DNPE ∼1.6 US$/L. Note that this unit cost might be reduced by 25% when increasing the life in service of the membranes to 1 year. 4.4. ZCMR vs. reactive distillation: economical feasibility The estimation of the operation cost reduction of an etherification reaction when carried out in a ZCMR compared to a reactive distillation unit is not straightforward. However, as a first approach, for 35000 tm DNPE/year production and taking into account a conversion about 10% in the reaction zone at 433 K and at least 505 kPa total pressure (this low conversion is ascribed to the low volatility of DNPE at the reaction temperature), reactive distillation would involve an increase up to 6 times of refrigeration water and steam demands. According to our estimates, the lower energy consumption in a ZCMR compared to a reactive distillation unit would translate into a reduction up to 60% of the operation costs in the former case. In this estimate, a mean life in service about 4 months for the zeolite membrane tubes has been considered.
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Table 5 Estimation of operation costs of the plant depicted in Fig. 7 for a 35,000 tm DNPE/year production. Designation
Unit
Amount per year
Unit cost [k$/unit]
OPEX [M$/year]
Raw materials n-pentanol Catalystb Membrane replacementb Handworking (HW)c
tm tm m2 Person
Staffc Water (refrigeration) Water (service) Steam Electricity Depreciationa Capital chargesa Taxesa Insurancea Maintenancea Supervision Others
Person m3 m3 tm kwh – – – – – – –
55,000 11 1700 × 3 Direct: 32 Indirect: 8 3 1.3 × 106 1.3 × 104 2.6 × 105 1.2 × 105 10% CCequip 5% CCplant 3% CCplant 3% CCplant 3% CCplant 25% HW –
0.5 20 5 45 30 65 0.015 0.15 0.04 0.045 – – – – – – –
27.5 0.2 25.4 1.4 0.2 0.1 13.3 1.3 3.8 0.1 0.9 1.4 0.3 0.9 0.9 0.7 0.7
Total OPEX
79.2
CC plant = 3.2 × CCequip (equipment + installation). b A mean life of ∼4 months is considered for the membranes and catalyst. c Calculations carried out assuming a personnel constituted of 50 people with the distribution 3 technician: 4 specialist: 5 non-qualified personnel: 2 administrative: 1 staff member. a
5. Conclusions and perspectives According to the technico-economical analysis provided in this study, hydrophilic zeolite NaA membranes appear to be promising candidates to carry out liquid-phase etherification reactions kinetically restricted by the formation of water. The coupling between the catalytic activity and water extraction by the membrane might allow a relevant enhancement of the reactant conversion. Moreover, a higher purification of the ether in terms of lower water concentration at the outlet of the reactor could be achieved. Compared to a reactive distillation unit, a CMR system might allow a significant reduction of operation costs due to a drastic reduction of energy demands. According to our experience gathered in the field of zeolite membranes, one of the main limitations of this kind of membranes is their reduced life in service due to the crack formation in the zeolitic film during operation. A possibility to increase the resistance of the zeolite material, reducing accordingly the operation costs, could be the use of nanocomposite architectures, where zeolite NaA would be grown inside a porous matrix instead of forming a film. High-thermal long-term resistant MFI-type membranes synthesized using this concept have been reported by Dalmon’s group for more than a decade [29,30]. This architecture might not also help improving the resistance of the zeolite material in terms of shrinking, but to avoid possible abrasion by the catalyst during operation. The use of A-type zeolite membranes is, however, restricted to reactions subjected to heterogeneous acid catalysis to avoid dealumination. In the presence of stronger acidic conditions (e.g. using homogeneous catalysis) and at high temperatures, higher silicacontaining zeolites like mordenite, faujasites, phillipsite or DDR would be more suitable. The use of these materials is, however, restricted to the preparation of reproducible and long-term resistant membranes with higher enough selectivities and fluxes, a challenging subject by now according to the present state-of-the art of their synthesis. Acknowledgments The authors would like to express their gratitude to the Spanish Ministry of Education and Science for financial support (project
CTQ2005-08346-C02-01). M.P.-T. gratefully acknowledges Profs. M. Villarrubia and L. Jutglar, both from the Applied Physics Dept. of the UB, for their helpful comments in the preparation of this manuscript. Appendix A. Nomenclature
a am Ab A,B d¯ C Db ÐS E K, Kw ZA L M N Po P* P qM r R T uo w x y z
activity specific surface of the membrane (m2 m−3 ) section of the packed bed (m2 ) constants in the Antoine equation mean pore size of defects (m) inner diameter of membrane tubes (m) Maxwell–Stefan surface diffusivity (m2 s−1 ) activation energy (kJ mol−1 ) constants in Eq. (5) thickness of the zeolite layer (m) length of the membrane tubes (m) molecular weight (kg mol−1 ) flux across the membrane (kg m−2 h−1 ) pressure inside the tubes (kPa) saturation vapour pressure (kPa) pressure (Pa) molar saturation loading (mol kg−1 ) reaction rate (mol h−1 kgcat −1 ) constant of gases (8.314 J mol−1 K−1 ) temperature (K) surface velocity inside the tubes (m s−1 ) molar flow (kmol/h) molar fraction in liquid feed/retentate molar fraction in vapour permeate axial position in the membrane tubes (m)
Greek symbols ˛ membrane selectivity H◦ adsorption enthalpy (kJ mol−1 ) εb porosity of the packed bed intercrystalline porosity in the zeolite NaA layer εinter activity coefficient dimensionless axial position in the reactor
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K
adsorption constant (kPa−1 ) dimensionless position in the membrane tubes parameter in Eq. (5) density (kg m−3 )
Subscripts A alcohol in initial P permeate R retentate s surface S catalyst v vapour w water ZA zeolite NaA References [1] M. Kondo, M. Komori, H. Kita, K.-I. Okamoto, Tubular-type pervaporation module with zeolite NaA membrane, J. Membr. Sci. 133 (1997) 133. [2] T.C. Bowen, R.D. Noble, J.L. Falconer, Fundamentals and applications of pervaporation through zeolite membranes, J. Membr. Sci. 245 (2004) 1. [3] R.M. Barrer, Hydrothermal Chemistry of Zeolites, Academic Press, London/New York, 1982. [4] M. Pera-Titus, J. Llorens, F. Cunill, R. Mallada, J. Santamaría, Preparation of zeolite NaA membranes on the inner side of tubular supports by means of a controlled seeding technique, Catal. Today 104 (2005) 281. [5] M. Pera-Titus, R. Mallada, J. Llorens, F. Cunill, J. Santamaría, Preparation of innerside tubular zeolite NaA membranes in a semi-continuous synthesis system, J. Membr. Sci. 278 (2006) 401. [6] M. Pera-Titus, M. Bausach, J. Llorens, F. Cunill, Preparation of inner-side tubular zeolite NaA membranes in a continuous flow system, Sep. Purif. Technol. 59 (2008) 141. [7] K.-I. Okamoto, H. Kita, K. Horii, K. Tanaka, Zeolite NaA membrane: preparation, single-gas permeation, and pervaporation and vapor permeation of water/organic liquid mixtures, Ind. Eng. Chem. Res. 40 (2001) 163. [8] D. Shah, K. Kissick, A. Ghorpade, R. Hannah, D. Bhattacharyya, Pervaporation of pharmaceutical waste streams and synthetic mixtures using water selective membranes, J. Membr. Sci. 179 (2000) 185. [9] Y. Morigami, M. Kondo, J. Abe, H. Kita, K. Okamoto, The first large-scale pervaporation plant using tubular-type module with zeolite NaA membrane, Sep. Purif. Technol. 25 (2001) 251. [10] M. Pera-Titus, J. Llorens, F. Cunill, On a rapid method to characterize intercrystalline defects in zeolite NaA membranes using pervaporation data, Chem. Eng. Sci. 63 (2008) 2367. [11] H.P. Hsieh, Inorganic Membranes for Separation and Reaction, Elsevier, Amsterdam, 1996.
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[12] P. Huang, Pervaporation Separation Processes, Elsevier, Amsterdam, 1991. [13] S. Assabumrungrat, W. Kiatkittipong, P. Praserthdam, S. Goto, Simulation of pervaporation membrane reactors for liquid phase synthesis of ethyl tertbutyl ether from tert-butyl alcohol and ethanol, Catal. Today 79–80 (2003) 249. [14] M.A. Salomon, J. Coronas, M. Menendez, J. Santamaria, Synthesis of MTBE in zeolite membrane reactors, Appl. Catal. A: Gen. 200 (2000) 201. [15] O. de la Iglesia, R. Mallada, M. Menendez, J. Coronas, Synthesis of MTBE in zeolite membrane reactors, Chem. Eng. J. 131 (2007) 35. [16] T. Inoue, T. Nagase, Y. Hasegawa, Y. Kiyozumi, K. Sato, M. Nishioka, S. Hamakawa, F. Mizukami, Stoichiometric ester condensation reaction processes by pervaporative water removal via acid-tolerant zeolite membranes, Ind. Eng. Chem. Res. 46 (2007) 3743. [17] G.C. Pecci, M.G. Clerici, F. Giavazzi, F. Ancillotti, M. Marchionna, R. Patrini, Oxygenated fuel diesels. Part 1: structure and properties correlation, in: 9th International Symposium on Alcohol Fuels, 1991, p. 321. [18] M. Marchionna, R. Patrini, F. Giavazzi, G.C. Pecci, Linear ethers as high-quality components for reformulated diesel fuels, in: Symposium on the Removal of Aromatics, Sulfur and Olefins from Gasoline and Diesel, 212th National Meeting, American Chemical Society, Washington, DC, 1996, p. 585. [19] J. Van Heerden, J.J. Botha, N.J. Roets, Improvements of diesel performance with the addition of linear ethers to diesel fuels, in: Proceedings of the 12th International Symposium on Alcohol Fuels, Beijing, 1998, p. 188. [20] R.C. Weast (Ed.), Handbook of Chemistry and Physics, 52nd ed., Ohio, 1971–1972. [21] R.W. Gallant, Physical Properties of Hydrocarbons, vol. 1, Gulf Publishing Co., Houston, TX, 1968. [22] R.C. Reid, J.M. Prausnitz, B.E. Poling, The Properties of Gases and Liquids, 4th ed., McGraw-Hill, NY, 1987. [23] R. Bringué, M. Iborra, J. Tejero, J.F. Izquierdo, F. Cunill, C. Fité, V.J. Cruz, Thermally stable ion-exchange resins as catalysts for the liquid-phase dehydration of 1pentanol to di-n-pentyl ether (DNPE), J. Catal. 224 (2006) 33. [24] R. Bringué, J. Tejero, M. Iborra, J.F. Izquierdo, C. Fité, F. Cunill, Water effect on the kinetics of 1-pentanol dehydration to di-n-pentyl ether (DNPE) on amberlyst 70, Topics Catal. 45 (2007) 181. [25] R. Bringué, J. Tejero, M. Iborra, J.F. Izquierdo, C. Fité, F. Cunill, Experimental study of the chemical equilibria in the liquid-phase dehydration of 1-pentanol to di-n-pentyl ether, Ind. Eng. Chem. Res. 46 (2007) 6865. [26] H.F. Rase, Chemical Reactor Design for Process Plants, vol. I, John Wiley & Sons, Inc., NY, 1977. [27] R.H. Perry, Perry’s Chemical Engineering Handbook, 6th ed., McGraw-Hill, NY, 1984. [28] A. Criscuoli, A. Basile, E. Drioli, O. Loiacono, An economic feasibility study for water gas shift membrane reactor, J. Membr. Sci. 181 (2001) 21. [29] A. Giroir-Fendler, J. Peureux, H. Mozzanega, J.-A. Dalmon, Characterisation of a zeolite membrane for catalytic membrane reactor application, Stud. Surf. Sci. Catal. 101A (1996) 127. [30] S. Miachon, E. Landrivon, M. Aouine, Y. Sun, I. Kumakiri, Y. Li, O. Pachtová Prokopová, N. Guilhaume, A. Giroir-Fedler, H. Mozzanega, J.-A. Dalmon, Nanocomposite MFI-alumina membranes via pore-plugging synthesis: preparation and morphological characterization, J. Membr. Sci. 281 (2006) 228.
Contents lists available at ScienceDirect
Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep
Technical and economical feasibility of zeolite NaA membrane-based reactors in liquid-phase etherification reactions Marc Pera-Titus ∗ , Joan Llorens, Fidel Cunill Chemical Engineering Department, University of Barcelona, 08028 Barcelona, Spain
a r t i c l e
i n f o
Article history: Received 2 October 2008 Received in revised form 8 December 2008 Accepted 28 February 2009 Available online 14 March 2009 Keywords: Pervaporation Zeolite NaA Membrane reactor Di-n-pentylether Dehydration
a b s t r a c t This paper describes the pervaporation performance of film-like zeolite NaA membranes synthesized in our laboratory by template-free seeded hydrothermal towards the dehydration of n-pentanol/water and n-pentanol/water/DNPE mixtures (DNPE = di-n-pentyl ether, C10 H22 O). DNPE is a linear symmetric ether that can be used as blending additive in reformulated diesel fuels and that can be produced from C4 feedstocks via n-pentanol, obtained in its turn by selective hydroformylation of linear butenes. The membranes showed selectivities up to 3000 in the dehydration of n-pentanol/water binary mixtures, and water/DNPE and n-pentanol/DNPE selectivities approaching infinity (i.e. DNPE does not pervaporate) in the dehydration of n-pentanol/water/DNPE ternary mixtures. On the basis of the separation results presented in this work, and using experimental reaction kinetics data, we discuss the feasibility of zeolite NaA membrane reactors to carry out the liquid-phase etherification reaction of n-pentanol to di-n-pentyl ether (DNPE) catalyzed by ion-exchange sulfonated resins. These catalysts suffer from strong deactivation in the progress of the reaction due to the generation of water. © 2009 Elsevier B.V. All rights reserved.
1. Introduction Zeolite membranes have been largely advocated as alternative candidates to polymeric membranes for the separation of species relying on molecular size, adsorption affinity and/or surface diffusion differences [1]. In the special case of liquid mixture separation by pervaporation (PV), separation mainly occurs on the basis of adsorption differences [2]. This is the case for instance of solvent dehydration by zeolite NaA membranes, where water adsorbs preferentially on the zeolite material due to the low Si/Al ratios A-type zeolites [1,4–8] (approaching to 1, the minimum allowed by the Loewenstein rule [3]). The first open report on large-scale alcohol dehydration using A-type zeolite membranes dates back to 2001 [9], when the Japanese Mitsui Engineering & Shipbuilding Co. started up in Japan a plant for isopropanol dewatering by vapour permeation (VP) with a production of 20,000 L/day. By now, this company and its subsidiary the Nano-Research Institute Inc. (XNRI), as well as the European alliance between Smart (UK) and Inocermic (Germany), supply commercially LTA zeolite membranes for PV and VP dehydration separations. New VP demonstration plants have been recently installed in Brazil (3000 L/day) and India (30,000 L/day) for bioethanol dehydration.
∗ Corresponding author. Present address: Institut de Recherches sur la Catalyse et l’Environnement de Lyon (IRCELYON), UMR 5256-CNRS/University of Lyon, France. E-mail address: [email protected] (M. Pera-Titus). 0255-2701/$ – see front matter © 2009 Elsevier B.V. All rights reserved. doi:10.1016/j.cep.2009.02.006
Three are the main drawbacks when dealing with zeolite NaA membranes compared to other zeolite membranes: (1) this zeolite suffers from dealumination at strong acidic conditions, which in practice limits its application to environments with pH > 6; (2) the presence of a large amount of Al in the framework structure acts as a promoter for membrane shrinkage; and (3) hydrothermal synthesis is in practice difficult to scale up due to formation of large intercrystalline defects in the active layer. This third shortcoming can be partially overcome, for instance, by seeding the support surface prior to hydrothermal synthesis [1,4], or by refreshing the synthesis solution in direct contact with the support. Gel renewal can be accomplished either by the action of a centrifugal force field [4], or by pulse- or continuous renewal of the synthesis solution [5,6]. One of the advantages of pervaporation compared to gas separation is that, despite the presence of a reduced amount of defects in the zeolite membrane structure, these can still show selectivity due to strong water adsorption. In addition, as we have suggested in a previous study [10], high compressive tensions due to negative capillary forces might still allow discrimination between water and organic species in small-sized mesopores (i.e. <5 nm), contributing positively to the membrane separation performance. Taking into account all the above stated aspects, zeolite NaA membranes appear to be promising candidates for in situ water removal in extractor-type catalytic membrane reactor (CMR) applications. This is the case for instance of etherification reactions carried out by alcohol dehydration and catalyzed by ion-exchange acid resins, usually suffering from severe catalyst deactivation due
M. Pera-Titus et al. / Chemical Engineering and Processing 48 (2009) 1072–1079
to strong water adsorption on the resin active sites. The use of a CMR with selective water extraction from the reaction system might not only improve the catalytic activity, but also help circumventing the high-energy costs ascribed to reactive distillation. Note that, as a rule, membranes might help saving up to 85% of the energy demands ascribed to distillation [11], involving a reduction of 30–50% of the energy costs [12]. Several examples of this concept applied to the synthesis of MTBE and ETBE catalyzed by ion-exchange resins have been reported in the literature [13,14]. Acid-resistant mordenite (MOR), merlionite (MER) and phillipsite (PHI) membranes [15,16] have been applied as well to esterification reactions catalyzed either by mineral or solid catalysts for shifting the equilibrium position. As an example, in this paper we focus our attention on the synthesis of di-n-pentyl ether (DNPE) from bimolecular dehydration of n-pentanol (SN 2) catalyzed by acid ion-exchange resins. DNPE is a linear and symmetric ether of molecular formula C10 H22 O that, among other long-chain oxygenates, can be used as blending additive in reformulated diesels to increase their cetane number, improve their cold flux properties, and reduce tail-pipe emissions (especially polyaromatics and sulfur) [17,18]. This ether also shows industrial interest, since n-pentanol can be produced industrially from valorization of C4 feedstocks by selective hydroformylation of linear butenes (oxo process) [19]. On the basis of experimental n-pentanol/water/DNPE dehydration data, as well as reaction rate data of DNPE formation, we discuss how the coupling of a zeolite NaA membrane with an ion-exchange resin in an extractortype CMR configuration could enhance DNPE synthesis from n-pentanol dehydration. We also discuss the technico-economical feasibility of DNPE mass production in a zeolite NaA catalytic membrane reactor (ZCMR) compared to a reactive distillation unit. 2. Experimental The membranes used in this study consisted of a continuous active film (thickness 30–35 m) of randomly oriented zeolite NaA crystals located on the inner surface of porous ␣-alumina tubular supports (active surface 12.6 cm2 , effective length 5 cm) grown by seeded hydrothermal synthesis. The main characteristics of the membranes are summarized in Table 1. The details dealing with the membrane synthesis can be found in Refs. [4,5]. All the membranes displayed low N2 permeance (<10−7 mol m−2 s−1 Pa−1 ) and good PV performance for the separation of 92:8 wt.% ethanol/water mixtures at 323 K, feed pressure in the range 101–303 kPa, and permeate pressure <0.2 mbar. The membranes showed an extremely low amount of defects or intercrystalline porosity (εinter < 10−6 ), as inferred using the characterization method reported in Ref. [10]. A set of steady-state PV experiments with a set of C1–C5 primary alcohol/water mixtures and n-pentanol/water/DNPE mixtures were performed in a laboratory-scale PV test bench. The details dealing
Table 1 Water/ethanol selectivity (˛w/E ), total flux (NT ), zeolite film thickness (ZA ) and defective structure (d¯ C : mean pore size; εinter : intercrystalline porosity) of the zeolite NaA membranes used in this study. Membrane
ZA1 ZA2 ZA3 a
˛W/E
185 294 1084
NT [kg m−2 h−1 ]
0.58 0.36 0.47
ZA [m]a
35 30 35
Large defectsb d¯ C [nm]
εinter × 107
589 ± 8 397 ± 4 490 ± 4
5.6 ± 0.6 2.9 ± 0.2 1.8 ± 0.3
Determined by SEM. Determined from the evolution of the total pervaporated flux with the feed pressure according to the characterization method described in Ref. [8]. b
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Fig. 1. Evolution of total flux and water/alcohol and water/mixture selectivities as a function of the chain length of the primary alcohol for membrane ZA2. Conditions: xw = 6.6–8.0 mol%, T = 323 K, Po = 303 kPa, Pv = 1 mbar, and xDNPE = 19.4–20.4 mol% (ternary mixture).
the experimental setup and protocols to carry out the PV tests are described elsewhere [4]. Methanol, ethanol and 1-propanol (99.5% purity) were supplied by Panreac (Spain), 1-butanol (99.7% purity) by Romil (England) and 1-pentanol (99.0% purity) by Aldrich (Germany). DNPE (98.0% purity) was synthesized in our laboratory in a batch reactor from dehydration of n-pentanol. Care was taken in the experiments involving BuOH and PeOH alcohols and DNPE to ensure complete miscibility of their liquid mixtures with water. The physical properties of the alcohols and DNPE were obtained from Refs. [20,21]. The density of n-pentanol and DNPE are, respectively, 0.815 and 0.787 g cm−3 , while their corresponding normal boiling points are 141.2 and 183.7 ◦ C. The water/alcohol selectivity was calculated as the quotient between the ratios of the molar fractions of water and alcohol at the permeate and feed sides, i.e. ˛w/A = (Yw /YA )/(Xw /XA ). In the case of n-pentanol/water/DNPE mixtures, the selectivity was calculated as the ratio of the molar fractions of water and mixture (mixture = PeOH + DNPE) at the permeate and feed sides. The composition in the feed and permeate was analyzed in a HP6980A GC equipped with a 50 cm × 0.2 mm × 0.5 mm methyl silicone capillary column and a TCD detector. The final values correspond to the mean of at least 2–6 steady-state measurements during 1–3 h taken after 2–4 h of stabilization. The reproducibility of the experiments was optimal, with standard deviations of the total flux and selectivity lower than 5% and 10%, respectively. 3. Results Fig. 1 shows the evolution of the water/alcohol selectivity and total flux with the chain length of the primary alcohol in the PV of alcohol/water binary mixtures within membrane ZA2 at 323 K and for a ∼7 mol% water composition in the feed. As can be seen, while the total flux seems to be practically independent on the nature of the alcohol in the mixture within the limits of the experimental error (except for methanol), the water/alcohol selectivity appears to be strongly affected. Fig. 1 also shows that, in the presence of DNPE (xDNPE = 19.4–20.4 mol%), the water/mixture selectivity appears to be promoted when compared to the value obtained for the PeOH/water binary system at the same temperature and water composition. However, the total flux appears to be unaffected by the presence of DNPE. Fig. 2 shows the effect of the feed composition on the total flux and water/alcohol selectivity at 323 K for EtOH/water and
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Fig. 2. Evolution of total flux (on top) and water/alcohol (or water/mixture) selectivity (on bottom) with water feed composition for membranes ZA1 and ZA2. Conditions: T = 323 K, Po = 303 kPa, Pv = 1 mbar, and xDNPE = 19.4–20.4 mol% (ternary mixture). The lines are a guide to the eye.
PeOH/water mixtures at low feed water compositions (2–30 mol%) for membranes ZA1 and ZA2. As can be seen, the total flux increases practically linearly with the water composition in the range 7.5–18 mol% with a similar slope for both mixtures, while the selectivity remains practically invariable, except for membrane ZA1, showing a slight increase. For membrane ZA2, the water/PeOH selectivities are one order of magnitude higher than those obtained for the EtOH/water system. Fig. 3 plots the effect of temperature on the water flux and water/alcohol selectivity for membrane ZA2 at low feed water compositions (2–30 mol%) for the dehydration of EtOH/water and PeOH/water binary mixtures and PeOH/water/DNPE ternary mixtures. As can be seen, water flux shows an Arrhenius trend for all the mixtures, with effective activation energies for water of 34 kJ mol−1 for EtOH/water mixtures and 20–24 kJ mol−1 for PeOH/water and PeOH/water/DNPE mixtures. For the latter two mixtures, regardless of the presence of DNPE, the total flux seems to depend only on feed composition. Furthermore, the water/EtOH and water/PeOH selectivities are almost constant with temperature in the range 303–410 K, the latter being one order of magnitude higher than the former, with values up to 3000. Moreover, water/PeOH selectivities are higher for higher water feed compositions (range 22.9–27.9 mol%), the values being comparable to those obtained for water/mixture selectivities in the presence of ∼20%mol DNPE in the liquid feed. However, in the case of ternary mixtures, the water/mixture selectivities show a slight increase with temperature.
Fig. 3. Evolution of water flux (on top) and water/alcohol selectivity or water/mixture selectivity (on bottom) with temperature for membrane ZA2. Conditions: Po = 303 kPa, Pv = 0.2–1 mbar, and xDNPE = 19.4–20.4 mol% (ternary mixtures). The lines for selectivity are a guide to the eye, while those for water fluxes correspond to Arrhenius-type fittings.
4. Discussion 4.1. Equation for water flux within zeolite NaA membranes The tends plotted in Figs. 1–3 reflect that, as expected, hydrophobic species are hardly separated by zeolite NaA membranes due to the inherent hydrophilic character of this zeolite. Moreover, the fact that water/mixture selectivities in the presence of DNPE are higher than PeOH/water selectivities obtained for the binary system at the same temperature and comparable water composition should be ascribed to the stronger hydrophobic character of the organic phase in the presence of DNPE, impelling therefore its separation. The correlations plotted in Figs. 1–3 allow the derivation of an equation for water flux from an adsorption–diffusion model formulated on the basis of the Maxwell–Stefan (MS) theory using the approach outlined in a previous study [6]. At low water feed activities, aw , and sufficiently low vacuum permeate pressure, the water flux can be expressed by the following equation (pseudo solutiondiffusion model): S Nw =
S S qM,w Mw DSw Kw o aw Pw ZA
(1)
Eq. (1) can be rewritten to Eq. (2) taking into account the explicit dependence of the MS surface diffusivity, the adsorption constant and the saturation vapour pressure of water with temperature (Eq.
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Fig. 5. Scheme of a zeolite NaA catalytic membrane reactor (extractor-type) with the catalyst packed inside the membrane tubes.
Fig. 4. Evolution of experimental and fitted water fluxes for PeOH/water and PeOH/water/DNPE mixtures in the water composition and temperature ranges 0.05–0.30 mol% and 303–403 K, respectively. The DNPE composition in the ternary mixtures was 19.4–20.4 mol%.
(2)): S = Nw
S (0) exp(B) S qM,w Mw DSw (0)Kw exp ZA
S Ew,eff
RT
(2)
aw
S In Eq. (2), the effective activation energy of water, Ew,eff , includes three main contributions (Eq. (3)): S Ew,eff
=
S Ew
◦ + Hw
+ RA
(3)
S is the activation energy for water surface diffusion, H ◦ where Ew w is the adsorption enthalpy of water, and parameter A belongs to the Antoine equation accounting for the evolution of saturation vapour pressure of water with temperature (A = 5300 K−1 in the temperature range 293–363 K). Eq. (2) has been fitted to the experimental water flux data in PeOH/water and PeOH/water/DNPE mixtures plotted in Figs. 2 and 3 to derive an expression useful for modelling purposes. In the fitting process, the effective activation energy of water and the preexponential factor in Eq. (2) have been adjusted by the least-square method using a Levenberg–Marquard optimization algorithm. The activity coefficients of water in the mixtures have been estimated by the UNIFAC method [22]. The final expression obtained is described by the following equation:
S = (1.49 ± 0.09) exp −(2700 ± 220) Nw
1 T
−
1 353
[kg m−2 h−1 ]
aw (4)
Fig. 4 shows the good agreement between the experimental and fitted water fluxes, the adjusted parameters lacking of significant correlation effects. 4.2. Synthesis of DNPE in a zeolite membrane reactor The strong dehydration performance of the zeolite NaA membranes synthesized in this work is especially remarkable for the dehydration of the PeOH/water/DNPE mixtures, since PV of DNPE within these membranes is practically negligible. In view of this result, zeolite NaA membranes appear to be promising candidates for DNPE synthesis in zeolite membrane reactors in the absence of acid conditions. On the basis of the results obtained for PeOH/water/DNPE separation in zeolite NaA membranes, we present here the simulation of a multitubular ZCMR to carry out the liquid-phase etherification reaction of n-pentanol to DNPE. For comparison purposes, we have also performed some simulations
on a conventional fixed-bed reactor to illustrate the enhancement of n-pentanol conversion that could be achieved when coupling the catalyst with selective in situ water removal. As an example, we have chosen the recently developed thermally stable S/DVB Amberlyst 70 resin, offering high activity and selectivity to DNPE (the main byproducts are pentenes generated by bimolecular elimination, E2). The kinetics of the main reaction can be well described by a model stemming from an Eley–Rideal (ER) mechanism wherein DNPE remains practically adsorbed on the catalyst surface, the number of free active sites is negligible, and the surface reaction is the rate limiting step (Eq. (5)) [23,24]. rDNPE =
k(a2PeOH − aw aDNPE /K)
(5)
1/2
aPeOH (1 + Kw aw )
with K = 8.9229 exp
778.69
T
Kw = exp 1.46 − 6615
1
k = exp 2.808 − 11595
T
1 T
(6)
−
1 438
−
1 438
(7)
[mol h−1 kgcat −1 ]
(8)
In Eq. (5), ai is the activity of species i, the activity coefficients having been estimated by the UNIFAC method. The term (1 + Kw aw 1/2 ) in the denominator of Eq. (5) accounts for catalyst deactivation due to water adsorption. It is noteworthy that, according to the expression of the equilibrium constant (Eq. (6)), the reaction is only slightly exothermic, with a reaction enthalpy only about −6.5 kJ mol−1 [25]). The multitubular ZCMR modelled in this study consists of a catalyst packed bed in the lumen of the membrane tubes, while the vapour permeate side of the membrane is kept under vacuum (see Fig. 5). Assuming a quasi-isothermal regime, the ZCMR can be modelled by a microscopic mass balance for each species in the lumen and in the permeate sides of the tubes. For the sake of simplicity, plug-flow and perfect mixing regimes have been assumed, respectively, to describe the hydrodynamics of both zones. The pressure drop along the axial position inside the membrane tubes can be modelled by the Ergun Equation. The set of Eqs. (9)–(11) is obtained at steady state: • Microscopic mass balance (i = 1,N) −
∂(wR xi ) − Ni am + S (1 − εb )ri = 0 Ab ∂z
(9)
• Permeation (i = 1,N) −Ni am =
∂(wP yi ) Ab ∂z
(10)
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Table 2 Input data for reactor modelling. Parameter
Value
xPeOH,in xw,in = xDNPE,in T [K] Po,in [kPa] Pv [mbar] Residence time [kgcat h kmol−1 ] PeOH mass flow [tm year−1 ] Db [mm] am [m2 m−3 ] uo range (lumen) [m min−1 ] εb Catalyst: Amberlyst 70 [23] Density [kg m−3 ] Apparent density [kg m−3 ] Surface area [m2 g−1 ] Dp [m]
1.00 0.00 443 500 <5 53 55,000 20 300 0.20–0.70 35% 1542 647a 176a 549b
a
Values obtained in swollen state (ISEC method). This particle size ensures the absence of significant internal mass (IMT) contribution to the reaction rate. b
• Pressure drop (Ergun equation) ∂Po (1 − εb )2 L uo (1 − εb ) L u2o = 150 + 1.75 3 2 Dp ∂z Dp εb ε3b Boundary conditions: z = 0 → xi = xi,in , yi = yi,in ,
(11)
Po = Po,in
It is interesting to note that the set of Eqs. (9)–(11) for a ZCMR can be reduced to a fixed-bed reactor if Ni → 0 and to a tubular PV unit if ri → 0 for ∀i. The water permeation and reaction kinetics have been taken, respectively, from Eqs. (4) and (5). A value of 3000 has been chosen for the water/mixture selectivity of the membrane according to the results plotted in Figs. 2 and 3. The reactor model has been solved numerically through discretization using finite differences. The number of intervals (130) has been chosen to avoid any dependence of the simulation results on the discretization. As an example, the reactor has been simulated for a PeOH inlet flow of 55,000 tm/year. The input data used for the modelling are summarized in Table 2. Fig. 6 compares the simulation results obtained for n-pentanol conversion, reaction rate and water concentration inside the tubes along the dimensionless axial position, = z/L, in the case of a ZCMR and a fixed-bed reactor using the input data from Table 2. As can be seen, at 433 K and for a residence time of 53 kgcat h kmolPeOH −1 , the n-pentanol conversion in a ZCMR would reach a value up to 64%, while a value of 35% might only be obtained in a fixed-bed reactor. As can be deduced from Fig. 6 (on top), this difference would be ascribed to the higher reaction rates achieved in the former case due to selective water removal from the reaction mixture. Accordingly, a fairly stable and extremely low water composition (<0.8 mol%) would be obtained along the axial position in a ZCMR, while the water composition would increase progressively in a fixed-bed reactor. The output data obtained from our simulations are summarized in Table 3. According to our calculations, about 27,000 membrane tubes of 1 m length and 20 mm i.d. would be required. Table 3 also reflects the suitability of the plug-flow regime to describe the hydrodynamics inside the membrane tubes, since the conditions Db /Dp > 30 and L/Dp > 200 are fulfilled [26]. The surface velocity, uo , would be around 0.20 m min−1 , ensuring negligible external mass transfer effects on the reaction rate. The pressure drop along the reactor would be ∼7.3 kPa, translating into ∼13 kW pumping requirements for the given tube effective length. Moreover, ca. 98% of the water generated during the reaction would be removed in the permeate stream of a ZCMR significant reactant loss (<1%).
Fig. 6. Evolution of PeOH conversion and reaction rate (on top) and water molar fraction (on bottom) with the axial position inside the reactor for a fixed-bed (thin lines) and a ZCMR (thick lines).
4.3. ZCMR-based process for DNPE production: techno-economical analysis On the guidance of the simulation results obtained for DNPE production (35,000 tm/year) from n-pentanol dehydration in a ZCMR (64% reactor conversion), a process for its production can be proposed according to the process flow scheme illustrated in Fig. 7. In addition, Table 4 summarizes the values of the main parameters in the relevant streams. In principle, only one distillation column would be required at the outlet of the ZCMR to purify the DNPE, the n-pentanol and low amounts of water obtained in the distillate being recirculated to the reactor after condensation. The system would be also provided with a condenser operating at vacuum Table 3 Output data obtained from the simulations in the ZCMR. Variable
Value
Effective length [m] Number of tubes of 1 m length [–] XPeOH [%] Water removed [%] PeOH removed by membrane [%]
∼16 ∼27,000 64 >98 ∼1
Hydrodynamics Db /Dp L/Dp Pressure drop [kPa] uo [m min−1 ]
36 > 30 280 > 200 ∼7.3 kPa ∼0.20
M. Pera-Titus et al. / Chemical Engineering and Processing 48 (2009) 1072–1079
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Fig. 7. DNPE process flow scheme based on a catalytic zeolite membrane reactor. Nomenclature: C, condenser; DC, distillation column; E, expansion system; MR, membrane reactor; HE, heat exchanger; P, liquid pump; R, reboiler; VP, vacuum pump.
(pressure ∼1 kPa) equipped with a primary vacuum pumping unit to remove the water extracted by the membrane. The condensed water would be further removed by pumping to atmospheric pressure. The distillation column has been modelled by the Smith– Brinkley method extended to non-ideal mixtures [27]. According to our simulations, the system PeOH/water/DNPE is not expected to form any azeotrope neither in the distillate nor in the residual streams in the conditions indicated in Table 4. The analysis of the economic feasibility of the technological solution here proposed has been carried out in a similar way as that proposed by Criscuoli et al. [28] for costs analysis in the case of a water gas shift membrane reactor. In terms of capital cost (CAPEX), our estimations reflect that, for 35,000 tm DNPE/year production, about 30% of the investment costs of the plant would be ascribed to the membrane reactor. In this estimate, a unit cost of 5000 US$/m2 for the zeolite membrane tubes has been considered, the zeolite Table 4 Properties of the streams in Fig. 7. Stream
T [K]
P [kPa]
F [kmol/h]
xPeOH
xH2 O
xDNPE
Phase
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15
293.1 357.9 433.1 433.1 283.3 283.3 433.1 443.1 469.3 462.7 462.7 462.7 521.0 521.0 293.1
101 535 505 1 1 101 498 468 495 465 515 535 545 495 101
42.1 68.2 68.2a 21.4 21.4 21.4 47.2 47.2 139.2 139.2 139.2 26.1 122.7 21.1 21.1
1.0000 0.9926 0.9926 0.0525 0.0525 0.0525 0.5428 0.5428 0.9808 0.9808 0.9808 0.9808 0.0011 0.0011 0.0011
0.0000 0.0050 0.0050 0.9475 0.9475 0.9475 0.0072 0.0072 0.0131 0.0131 0.0131 0.0131 0.0000 0.0000 0.0000
0.0000 0.0024 0.0024 0.0000 0.0000 0.0000 0.4524 0.4524 0.0062 0.0062 0.0062 0.0062 0.9989 0.9989 0.9989
L L L V L L L L V L L L L L L
a Value corresponding to 55,000 tm/year n-pentanol included in Table 2 (64% reactor conversion).
layer being assumed to account for 30% of the zeolite NaA membrane cost. The cost of the equipments would be about 9 M$, 7.7 M$ being ascribed to the ZCMR, while the overall capital costs of the plant (equipment + installation) would reach 29 M$. The capital cost of the distillation column, including the condenser and reboiler, as well as the capital cost of the vacuum condenser, heat exchangers, vacuum pumps and liquid pumps have been estimated and upscaled using the Williams method, the costs having been capitalized to present monetary units using Marshall & Swift correction factors [27]. The analysis of the operational costs (OPEX) of the plant following conventional standards reflects a year cost about 80 M$ for 35,000 tm DNPE/year production (see Table 5). The higher contribution to the OPEX would be ascribed to raw materials (n-pentanol and catalyst replacement), membrane replacement and refrigeration water, accounting, respectively, for 33%, 32% and 19% of the overall costs. These costs would translate into a unit cost of DNPE ∼1.6 US$/L. Note that this unit cost might be reduced by 25% when increasing the life in service of the membranes to 1 year. 4.4. ZCMR vs. reactive distillation: economical feasibility The estimation of the operation cost reduction of an etherification reaction when carried out in a ZCMR compared to a reactive distillation unit is not straightforward. However, as a first approach, for 35000 tm DNPE/year production and taking into account a conversion about 10% in the reaction zone at 433 K and at least 505 kPa total pressure (this low conversion is ascribed to the low volatility of DNPE at the reaction temperature), reactive distillation would involve an increase up to 6 times of refrigeration water and steam demands. According to our estimates, the lower energy consumption in a ZCMR compared to a reactive distillation unit would translate into a reduction up to 60% of the operation costs in the former case. In this estimate, a mean life in service about 4 months for the zeolite membrane tubes has been considered.
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Table 5 Estimation of operation costs of the plant depicted in Fig. 7 for a 35,000 tm DNPE/year production. Designation
Unit
Amount per year
Unit cost [k$/unit]
OPEX [M$/year]
Raw materials n-pentanol Catalystb Membrane replacementb Handworking (HW)c
tm tm m2 Person
Staffc Water (refrigeration) Water (service) Steam Electricity Depreciationa Capital chargesa Taxesa Insurancea Maintenancea Supervision Others
Person m3 m3 tm kwh – – – – – – –
55,000 11 1700 × 3 Direct: 32 Indirect: 8 3 1.3 × 106 1.3 × 104 2.6 × 105 1.2 × 105 10% CCequip 5% CCplant 3% CCplant 3% CCplant 3% CCplant 25% HW –
0.5 20 5 45 30 65 0.015 0.15 0.04 0.045 – – – – – – –
27.5 0.2 25.4 1.4 0.2 0.1 13.3 1.3 3.8 0.1 0.9 1.4 0.3 0.9 0.9 0.7 0.7
Total OPEX
79.2
CC plant = 3.2 × CCequip (equipment + installation). b A mean life of ∼4 months is considered for the membranes and catalyst. c Calculations carried out assuming a personnel constituted of 50 people with the distribution 3 technician: 4 specialist: 5 non-qualified personnel: 2 administrative: 1 staff member. a
5. Conclusions and perspectives According to the technico-economical analysis provided in this study, hydrophilic zeolite NaA membranes appear to be promising candidates to carry out liquid-phase etherification reactions kinetically restricted by the formation of water. The coupling between the catalytic activity and water extraction by the membrane might allow a relevant enhancement of the reactant conversion. Moreover, a higher purification of the ether in terms of lower water concentration at the outlet of the reactor could be achieved. Compared to a reactive distillation unit, a CMR system might allow a significant reduction of operation costs due to a drastic reduction of energy demands. According to our experience gathered in the field of zeolite membranes, one of the main limitations of this kind of membranes is their reduced life in service due to the crack formation in the zeolitic film during operation. A possibility to increase the resistance of the zeolite material, reducing accordingly the operation costs, could be the use of nanocomposite architectures, where zeolite NaA would be grown inside a porous matrix instead of forming a film. High-thermal long-term resistant MFI-type membranes synthesized using this concept have been reported by Dalmon’s group for more than a decade [29,30]. This architecture might not also help improving the resistance of the zeolite material in terms of shrinking, but to avoid possible abrasion by the catalyst during operation. The use of A-type zeolite membranes is, however, restricted to reactions subjected to heterogeneous acid catalysis to avoid dealumination. In the presence of stronger acidic conditions (e.g. using homogeneous catalysis) and at high temperatures, higher silicacontaining zeolites like mordenite, faujasites, phillipsite or DDR would be more suitable. The use of these materials is, however, restricted to the preparation of reproducible and long-term resistant membranes with higher enough selectivities and fluxes, a challenging subject by now according to the present state-of-the art of their synthesis. Acknowledgments The authors would like to express their gratitude to the Spanish Ministry of Education and Science for financial support (project
CTQ2005-08346-C02-01). M.P.-T. gratefully acknowledges Profs. M. Villarrubia and L. Jutglar, both from the Applied Physics Dept. of the UB, for their helpful comments in the preparation of this manuscript. Appendix A. Nomenclature
a am Ab A,B d¯ C Db ÐS E K, Kw ZA L M N Po P* P qM r R T uo w x y z
activity specific surface of the membrane (m2 m−3 ) section of the packed bed (m2 ) constants in the Antoine equation mean pore size of defects (m) inner diameter of membrane tubes (m) Maxwell–Stefan surface diffusivity (m2 s−1 ) activation energy (kJ mol−1 ) constants in Eq. (5) thickness of the zeolite layer (m) length of the membrane tubes (m) molecular weight (kg mol−1 ) flux across the membrane (kg m−2 h−1 ) pressure inside the tubes (kPa) saturation vapour pressure (kPa) pressure (Pa) molar saturation loading (mol kg−1 ) reaction rate (mol h−1 kgcat −1 ) constant of gases (8.314 J mol−1 K−1 ) temperature (K) surface velocity inside the tubes (m s−1 ) molar flow (kmol/h) molar fraction in liquid feed/retentate molar fraction in vapour permeate axial position in the membrane tubes (m)
Greek symbols ˛ membrane selectivity H◦ adsorption enthalpy (kJ mol−1 ) εb porosity of the packed bed intercrystalline porosity in the zeolite NaA layer εinter activity coefficient dimensionless axial position in the reactor
M. Pera-Titus et al. / Chemical Engineering and Processing 48 (2009) 1072–1079
K
adsorption constant (kPa−1 ) dimensionless position in the membrane tubes parameter in Eq. (5) density (kg m−3 )
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