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Directing selectivity of ethanol steam reforming in membrane reactors
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Directing selectivity of ethanol steam reforming in membrane reactors Maria Anna Murmura b, Michael Patrascu a, Maria Cristina Annesini b, Vincenzo Palma c, Concetta Ruocco c, Moshe Sheintuch a,* a
Technion-Israel Institute of Technology, Department of Chemical Engineering, Haifa 32000, Israel University of Rome “Sapienza”, Department of Chemical Engineering, Materials and Environment, Via Eudossiana 18, 00184 Rome, Italy c University of Salerno, Dipartimento di Ingegneria Industriale, Via Giovanni Paolo II 132, 84084 Fisciano, SA, Italy b
article info
abstract
Article history:
In a system of parallel reversible reactions, separating the corresponding product can enhance
Received 21 January 2015
a desired reaction. If the same product is produced in several reactions, its concurrent sepa-
Received in revised form
ration enhances the reaction with the higher stoichiometry. Here we demonstrate this effect
26 February 2015
by separating hydrogen from ethanol during steam reforming in a Pd membrane reactor
Accepted 5 March 2015
packed with a Pt/NieCeO2 catalyst. Interest in ethanol SR stems from the need to produce
Available online 30 March 2015
ultra-pure H2 from renewables (the energy will be supplied by solar-heated molten salt).
Keywords:
H2, CO2, CH4 and CO as products. These products can be represented by three reactions
For the conditions tested full conversion (of reaction (1) below) has been achieved with Catalytic membrane reactor
(W ¼ H2O): 1. EtOH 4 CH4 þ CO þ H2 (ethanol decomposition), 2. CH4 þ 2W 4 CO2 þ 4H2
Reforming
(methane steam reforming, MSR) and 3. CO þ W 4 CO2 þ H2 (water gas shift, WGS).
Hydrogen
Separating H2 directs the selectivity towards CO and CO2, resulting also in increased ratio
Ethanol
between CO and CH4 mole fraction.
Modeling
In general, increasing temperature (613e753 K), pressure (6e10 bar) and introducing sweep flow (0.5 NL/min N2 for a similar feed rate) led to better separation, to an increase in selectivity towards CO and CO2 and in hydrogen yield. Increasing pressure and introducing sweep flow also increased hydrogen recovery. A further increase of sweep gas flow rate (to 1 NL/min) did not result in an appreciable improvement. The results of this work show that the combination of theNi/Pt catalyst and the Pd membrane for hydrogen removal produce very high values of hydrogen yield, despite the low steam to ethanol ratio and the moderate pressure levels. In particular about 4.5 mol H2/ mol EtOH were produced at 753 K, feed and sweep flow rates of 0.5 NL/min each in the entire pressure range examined. A one-dimensional model based on the kinetics of a limited but realistic set of reactions has been developed to simulate the behavior of the reactor. With a literature kinetics model and only two adjustable parameters, the permeance and heat transfer coefficient, a very good agreement with experimental data has been obtained. The results indicate strong permeance inhibition compared to pure hydrogen measurements. Copyright © 2015, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved.
* Corresponding author. Tel.: þ972 4 8202823. E-mail address: [email protected] (M. Sheintuch). http://dx.doi.org/10.1016/j.ijhydene.2015.03.013 0360-3199/Copyright © 2015, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved.
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CO þ H2O $ CO2 þ H2
Introduction The worldwide hydrogen demand is increasing, and its use in Fuel Cells (FCs) is expected to make hydrogen one of the widely-used fuels in a future energy system, mainly thanks to its low environmental impact when used in FC systems. On the other hand, the extensive introduction of hydrogen as energy vector is limited by the actual lack (and costs) of a reliable hydrogen distribution infrastructure. This barrier can be surmounted by the development of systems for decentralized hydrogen production (i.e. close to the end-user). Fuelflexibility and the possibility to power the process with renewables (solar and biomass) are two additional preconditions for sustainable hydrogen production. The above considerations are at the basis of the CoMETHy project (Compact Multifuel-Energy to Hydrogen converter). CoMETHy's general objective is to support the intensification of hydrogen production processes, by developing a membrane reformer that can operate with various fuels and that eventually will be heated by solar-heated molten salt [1]. An electrically-heated lab membrane reformer was tested in the Technion with methane SR [2], and this work extends that study to ethanol SR using the same catalyst, membrane and system. Hydrogen production from biomass derived oxygenates has attracted interest for its potential application in fuel cells. Bio-fuels for the production of hydrogen could bring significant environmental benefits, as the carbon dioxide produced is consumed in biomass production giving a CO2 neutral energy supply. Steam reforming of bioethanol has been widely investigated over supported transition and noble metal catalysts [3,4]. Hydrogen production by SR of ethanol often encounters several problems, which are addressed in the CoMETHy project: - High-purity hydrogen is required for its application in fuel cells: the most promising FCs, based on PEM, require low levels of CO; hydrogen separation by a Pd membrane should satisfy this condition. - Steam reforming is highly endothermic and requires heat supply; in this project the reformer is heated by solar energy. - The reaction is accompanied by the formation of various by-products, which greatly affect the selective production of hydrogen; as we show in this article, hydrogen separation will shift the product distribution in the desired direction, i.e., toward the reaction that produces most hydrogen. - The reaction is limited by equilibrium; hydrogen separation will shift the equilibrium conversion. The ethanol steam reforming (ESR) reaction is: C2H5OH þ H2O $ 2CO þ 4H2
(1)
The conversion of CO to CO2 is considered through the wateregas shift (WGS) reaction:
(2)
The overall ESR reaction, which refers to complete conversion of CO to CO2, is therefore:
C2H5OH þ 3H2O $ 2CO2 þ 6H2
(3)
Aside from these desired reactions, several others may take place [5].
C2H5OH $ C2H4O þ H2
(4)
C2H4O $ CH4 þ CO
(5)
C2H4O þ H2O $ 2CO þ 3H2
(6)
C2H5OH $ C2H4 þ H2O
(7)
2C2H5OH $ C3H6O þ CO þ 3H2
(8)
C2H5OH þ 2H2 $ 2CH4 þ H2O
(9)
C2H5OH $ 0.5CO2 þ 1.5CH4
(10)
CH4 þ H2O $ CO þ 3H2
(11)
2CO $ CO2 þ C
(12)
CO2 þ 4H2 $ CH4 þ 2H2O
(13)
Achieving the desired product is usually obtained by a proper choice of catalyst and support (see below) and operating conditions. Selectivity may be further directed through the use of membrane reactors, which shift the equilibrium in the direction of the selectively permeating component (hydrogen in this case). Recently, catalysts containing more than one active species have also been investigated because of their significantly different catalytic properties with respect to either of the parent metals [6,7]. A synergic effect of Pt addition to Ni-based catalysts was shown to improve the activity of the non-noble metal towards hydrogenation reactions, resulting in decreased coke formation rates with respect to the ones observed over monometallic Ni-catalysts [8]. When Pt and Ni are supported on CeO2, by depositing the non-noble metal on the support surface earlier than the noble one, as Pt is directly available at the gasesolid interface, ethanol adsorption is favored and a better agreement with thermodynamic data in a regular fixed bed was observed [9].
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Table 1 e Summary of experimental conditions and results previously reported in literature. Ref
T [K]
P [bar]
S/E
Sweep to feed flow rate ratio
Inlet flow rate [molEtOH/kgcatmin]
[10] [11] [12] [13]
598e673 573e673 673e723 673e723
5e14 1.3 1.5e2 1.5e2
6 3e9 8.4e13 8.4e13
0 0.5e1.4 0.9e3.7 0.9e3.7
1.67 0.83 0.04e0.64 0.04e1
Conversion and product distribution in ESR in membrane reactors (MRs) depend on various operating conditions: - Increasing temperature will increase hydrogen partial pressure and its permeating flux, leading to higher conversion and selectivity. - While increasing pressure has detrimental effect on conversion in a closed system at equilibrium, it will have an advantageous effect on conversion when significant hydrogen separation is achieved, as follows from stoichiometry. - At low temperatures, increasing the S/C ratio (to avoid coking) leads to an increase in H2 and CO2 selectivity, while decreasing CO and CH4 selectivity. Hydrogen recovery decreases due to dilution effects with increasing S/C ratios. Table 1 summarizes the main operating parameters and results of ethanol steam reforming in membrane reactors previously reported in literature. Domı`nguez et al. [10] used a PdeAg membrane for ESR on cobalt talc (Co3[Si2O5]2(OH)2). Studies conducted on a PdeAg membrane reformer with a Rubased catalyst at 1.3 bar [11] showed that increasing temperature shifts the reactions toward the production of CO2 and CH4. In a study over a metal foil membrane, using Ru, Pt, Ni based catalyst [12,13] positive effect of increasing the retentate pressure was observed. In these last studies the effect of co-current vs. counter-current sweep flow was also investigated. None of these works included the development of a model to describe the experimental results obtained. In the present work, experimental data collected on lowtemperature (753 K) ethanol steam reforming in a membrane reactor, equipped with a high permeance Pd membrane, are presented and analyzed, with a particular focus on the operating conditions which maximize selectivity towards carbon dioxide and hydrogen. Full ethanol conversion is achieved under conditions studied and the problem is reduced to competition between carbondioxide and methane production. High hydrogen yields and recoveries have been obtained even when working with stoichiometric steam to ethanol ratios (S/E ¼ 3), moderate pressures and with low ratios of sweep gas to feed gas flow rates. Previous works (Table 1) used high S/E ratios (6e13) and obtained high conversion due to dilution. No deactivation was detected even when running with a stoichiometric S/C ratio. This work, in a high permeance membrane, also presents better selectivities and hydrogen separation, and higher throughputs, than those presented in previous works. Furthermore, the work developed and tests a simple model that described well the results and can be used for further optimization; it shows that results are insensitive to catalytic activity.
EtOH conversion [%] 98.6e100 16e100 100 100
H2 yield 3.2e3.7 0.6e3.3 0.2e5 0.3e3.6
The structure of the paper is as follows: thermodynamic equilibrium analysis of a limited but realistic set of reactions is presented (Section Thermodynamic analysis), followed by a description of the experimental system and results obtained (Section Experimental activity). A one-dimensional model, developed to simulate the behavior of the reactor, is also presented and analyzed in Section Membrane reactor modeling.
Thermodynamic analysis Thermodynamic analysis will provide a limit to the system operation. Kinetic tests previously carried out on the catalyst formulation employed in the present work have shown that the system may be fully described by considering the following reactions: ethanol dehydrogenation (4), acetaldehyde decomposition (5), modified ethanol decomposition (10), acetaldehyde steam reforming (6), wateregas shift (2), and CO2 methanation (13) [5]. Since no acetaldehyde formation was noticed in the course of the present work, reactions (4) and (5) can be combined and (6) may be disregarded. Of the four remaining reactions only three are linearly independent, the system may therefore be studied by considering only three reactions. Table 2 summarizes the reactions considered, along with their DH0 and DG0. The same three reactions can be represented in other linear-dependent ways, leading to same results. The equilibrium conditions of the three reactions considered yield: KEDðTÞ ¼
pCO pCH4 pH2 yCO yCH4 yH2 ¼ P2TOT pEtOH yEtOH
(14)
KMSRðTÞ ¼
pCO2 p4H2 yCO2 y4H2 ¼ P2TOT 2 p2H2 O pCH4 yH2 O yCH4
(15)
KWGSðTÞ ¼
pH2 pCO2 yH yCO2 ¼ 2 pCO pH2 O yCO yH2 O
(16)
A preliminary thermodynamic analysis maps hydrogen partial pressure in a closed system as a function of total pressure and temperature (Fig. 1, S/C ¼ 3) while considering the three reactions mentioned above. Hydrogen separation occurs if its partial pressure in the reactor side exceeds that in the permeate side. Increasing pressure leads to a less than linear increase of hydrogen partial pressure, because hydrogen concentration decreases while the total pressure increases. To carry out low-temperature ESR efficiently, taking advantage of the entire membrane surface, it is therefore essential to consider the use of sweep gas in the permeate side.
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Table 2 e Reactions accounted for in ethanol steam reforming.
1 2 3
Reaction name
Reaction
DH0 [kJ/mol]
DG0 [kJ/mol]
Ethanol decomposition (ED) Methane steam reforming (MSR) Water gas shift (WGS)
C2H5OH $ CH4 þ CO þ H2 CH4 þ 2H2O(g) $ CO2 þ 4H2 CO þ H2O(g) $ CO2 þ H2
49.4 164.6 41.2
355.6 113.3 28.8
Assuming reaction (1) goes to completion the equilibrium selectivity ratio (s.r.) between CO and CH4 can be expressed as: If H2 is s:r:≡yCO =yCH4 ¼ KMSRðTÞ =ðKWGSðTÞ P2TOT Þ$ðyH2 O =y3H2 Þ. removed during reaction it will direct the selectivity towards CO and CO2 but the ratio yCO/yCH4 will increase by increasing the conversion of MSR to a greater extent than the conversion of WGS reactions. This is evident by the (yH2)-3 dependence in the above expression. Recall that yH2O also varies, and in a stoichiometric feed (as in this work), yH2O ~ yCO þ 2yCH4. Considering that a different subset of reactions may take place in correspondence of the inlet section (e.g., EtOH þ 3H2O $ 2CO2 þ 6H2, which goes to complete conversion) suggests that the hydrogen driving force will allow for some transport across the membrane even at temperatures lower than 700 K in the absence of sweep.
Experimental activity Experimental setup Experimental activity, designed to test ethanol steam reforming in a laboratory scale membrane reactor, has been carried out in the Chemical Reactor Engineering and Environmental Catalysis Laboratory of the Technion. The system employed is the one used for methane SR [2] after some modifications. The 40 cm long membrane reactor (Fig. 2a) consists of a tube-and-shell structure. Reactions take place in
Fig. 1 e Equilibrium partial pressure of hydrogen (in bar) as a function of temperature and pressure for feed of S/E ¼ 3. Gray area indicates feasible conditions for non-sweep operation assuming equilibrium compositions.
the shell side, packed with a structured foam catalyst to expedite radial heat transfer. Hydrogen permeates through the Pd membrane (the inner reactor wall, having a 14 mm outer diameter) into the tube-side, where a sweep gas may flow (counter-currently). Reactor characteristics have been summarized in Table 3. In the present work, nitrogen was used as a sweep gas; however, it is likely that steam will be used in the pilot scale application. This will allow an easy separation of hydrogen from the permeate stream. The stainless-steel reactor is heated from the outside (40 mm outer diameter) by four concentric electric heaters, shown as transparent bodies in Fig. 2a. Each heater is controlled independently to maintain a set temperature, measured on the outer wall of the reactor, in the midpoint of the heater body. A fifth heating element is placed at the entrance of the reactor. This last heating element acts as a vaporizer and its temperature is set to 300 C throughout all the experiments. Four additional thermocouples are located inside the membrane, in the same axial position as the outer control thermocouples. Pipeline elements were sealed using Hamlet® Let-Lok Tube Fittings (Double Ferrule). The effluent from the reactor is cooled, to allow the condensation of unreacted ethanol and steam, and sent to a silica gel column, where any residual vapor is absorbed. The dried effluent is then sent to an analyzing system, comprised of IR analyzers for CH4, CO2, and CO (Edinburgh Instruments Ltd.) and a gas chromatograph (Trace GC Ultra, ThermoScientific). Either the permeate or the retentate may be sent to the analyzing system. The liquid product has been collected and analyzed at the end of each experiment, showing that it only contained unreacted water. This indicates that full ethanol conversion has been achieved and that no liquid byproducts have been formed. This result is in line with the observations of Palma et al. [5], who showed that with the catalyst tested, acetaldehyde is present only for very short contact times and ethanol is fully converted almost immediately. The reactor is packed with structured foam catalyst, specifically prepared in the framework of the CoMETHy project from cooperative work of the Fraunhofer Institute for Ceramic Technologies and Systems, IKTS (Germany) and the ProCEEDlab of the University of Salerno (Italy). The structured catalyst is made of a CeO2-washcoated SiC open cell foam (see Fig. 2b), prepared by IKTS and activated with the bimetallic (Pt(3)Ni(10)) formulation by the University of Salerno. This formulation was specifically selected on the basis of the considerations mentioned in Section Introduction. Ni and Pt catalysts promote ethanol steam reforming when supported on ceria, which is expected to diminish deactivation due to coking and promote catalyst dispersion [14]. SiC open cell foams have been chosen as the structured carried for
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Fig. 2 e a) Detailed drawing of the membrane reactor; b) Structured foam.
Table 3 e Reactor characteristics. Bed and membrane length [m]
Reactor diameter [m]
Inner reactor diameter (membrane outer diameter) [m]
Mass of catalyts þ structural support [g]
34 103
14 103
200
0.4
their good heat transfer characteristics, and to minimize pressure drops [15]. The membrane tube (Hysep©-technology), provided by the Energy Research Centre of the Netherlands (ECN), is made of a 4e5 mm palladium layer, deposited by electroless plating on the outside of a porous alumina tube provided with two thin porous alumina layers to decrease the outer pore dimension to less than 0.15 mm. The membrane is capped with ECN compression seals on both sides. Membrane permeance was measured by the producer in pure hydrogen at various temperatures and pressures. The permeance varied with temperature according to an Arrhenius-type law (Em ¼ 11 kJ/mol, Am ¼ 5.6 mol/(m2 s bar0.5), see Eq. (21)) and hydrogen flux through the membrane followed Sievert's law. The experimental conditions tested in the present work have been summarized in Table 4.
Catalyst preparation Starting from the washcoated foam, the active species deposition (Pt and Ni) was performed at University of Salerno by multiple cycles of wet impregnation with an aqueous solution of the metal precursor (C4H6O4Ni∙4H2O or PtCl4, both from Aldrich), at 60 C for 1 h; the sample was then dried at 120 C for 2 h and calcined for 1 h at 400 C for Ni deposition and at 550 C for Pt deposition. The two different temperatures were selected on the basis of thermo-gravimetric analysis results. This impregnation-drying-calcination procedure was
repeated the number of times necessary to reach the desired amount of metal (3 wt% for Pt and 10 wt% for Ni), considering the solubility of the precursor salt. After the above cycles, the sample was calcined at 600 C for 3 h, to stabilize it. The whole procedure (impregnation-drying-calcination steps and final calcination) is repeated for each active species to be deposited on the washcoated foam.
Experimental results In all experiments conducted here complete conversion of ethanol was achieved with CO, CO2, CH4, and H2 as the only products. Permeate purity has been measured, and impurities, mainly methane and carbon dioxide, were always found in concentrations < 0.5% ( ~ 0.2% CH4, ~ 0.3% CO2). These measurements could be partly affected by the presence of residues in the pipeline, which was used to measure the retentate current as well. A summary of the experimental conditions and main results obtained has been reporte in Table 5. Experimental results have been expressed in terms of hydrogen recovery and yield: Hrec ¼
H2 permeated total H2 produced
Hyield ¼
total H2 produced ethanol in feed
(17)
where the maximum values obtainable are equal to 1 and 6, respectively; and in terms of selectivity towards carboncontaining species: Si ¼
moles of i produced 100 i ¼ CO2 ; CO; CH4 2ðmoles of ethanol convertedÞ
(18)
Table 4 e Experimental conditions. T [K] 613e753
Retentate pressure [bar]
Permeate pressure [bar]
Total feed flowrate [NL/min]
Sweep flowrate [NL/min]
S/E
6e10
1
0.5e1
0e1
3
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Table 5 e Summary of experimental conditions and results in present work. T [K] 613e753
P [bar]
S/E
Sweep to feed flow rate ratio
Inlet flow rate [molEtOH/kgcatmin]
EtOH conversion [%]
H2 yield
6e10
3
0e2
0.03e0.05
100
0.3e4.8
If complete ethanol conversion is assumed, the selectivity of each species may be determined by: Si ¼
yi 100 i ¼ CO2 ; CO; CH4 yCO2 þ yCO þ yCH4
(19)
In the figures that follow, product selectivity has been compared to its value at equilibrium calculated based on feed composition and outlet temperature. Selectivity to CO was found to be small under all conditions. Better separation of hydrogen, achieved by increasing temperature, decreasing flow rate and increasing sweep gas flow rates, will shift the reaction toward CO2 and suppress selectivity to methane. Hydrogen separation may lead to faster coking by reactions like (12). In this study catalytic activity seems to have remained stable during the 200 h of operation with ESR.
Influence of temperature In the absence of sweep flow, hydrogen separation is small and the product distribution is similar to that at equilibrium (Fig. 3a). Increasing temperature leads to small increase in selectivity towards CO and CO2 and a decrease in selectivity towards CH4, mainly due to higher methane steam reforming and reverse WGS conversions. This situation also leads to an increase in hydrogen yield (Fig. 4a). Deviations from equilibrium become significant only above 703 K, but in this range conversion to CO cannot be ignored. Hydrogen recovery decreases with increasing temperature, despite increasing membrane permeance (Fig. 4a). This occurs because hydrogen production increases, but its flow across the membrane is limited by the high partial pressure of hydrogen in the permeate. When a sweep flow rate of 1 NL/min is used (Fig. 3b), hydrogen separation is more significant, and selectivity towards CO2 and CO increases with increasing temperature reaching 80% and 4%, respectively, at 750 K, whereas
selectivity towards CH4 decreases to 15%. The positive effect of hydrogen separation is evident in Fig. 4b: in the entire temperature range hydrogen recovery is very close to 1, and the yield increases steadily with temperature. Under these conditions, the partial pressure of hydrogen in the permeate is reduced, allowing for efficient separation and significant deviation from equilibrium.
Influence of pressure Increasing pressure may have contrasting effects on the hydrogen-producing reactions, and therefore on the overall product distribution: the ethanol steam reforming process results in an increase in the number of moles, low pressures therefore shift the equilibrium conversion towards the products; however, a high retentate to permeate pressure gradient improves hydrogen separation. As separation is promoted the equilibrium of the reaction moves further towards the products. Fig. 5 shows the influence of pressure on the selectivity of carbon-containing species at 753 K, with a feed flow rate of 0.5 NL/min, and for two sweep gas flow rates. In the absence of sweep flow (Fig. 5a), increasing pressure leads to an increase in CO2 selectivity and a decrease in CH4 selectivity, indicating that permeation has a more significant effect than thermodynamics. Under these conditions, hydrogen recovery indeed increases with pressure, as may be seen by the results shown in Fig. 6a. The same may be said for the total hydrogen yield (Fig. 6a), which is favored by the removal of hydrogen itself. A similar trend in CO2 and CH4 selectivities is observed when working with sweep gas, even though the effect of pressure appears to be much less significant. Fig. 6b shows that under these conditions, hydrogen recovery is very close to 1 even when working at 6 bar, indicating that hydrogen removal has already reached its maximum value. The total hydrogen yield is also unaffected by pressure (Fig. 6b), which indicates that the system has reached its maximum performance and
Fig. 3 e Product selectivity as a function of temperature at P ¼ 8 bar, F ¼ 0.5 NL/min, without (a) and with a sweep flow of 1 NL/min (b), and its comparison with equilibrium values.
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Fig. 4 e Hydrogen recovery and yield as a function of temperature at P ¼ 8 bar, F ¼ 0.5 NL/min, without (a) and with a sweep flow of 1 NL/min (b).
Fig. 5 e Product selectivity as a function of pressure at Tw ¼ 753 K, F ¼ 0.5 NL/min, without (a) and with a sweep flow of 0.5 NL/min (b), and comparison with equilibrium values.
further hydrogen production may only be obtained by increasing the operating temperature or sweep flow.
Influence of sweep gas to feed flow rates ratio The effect of improving separation, without acting on other phenomena, may be better understood by studying the
influence of sweep gas flow rate. Increasing sweep flow rate (Fig. 7) leads to better selectivity towards CO2 and CO and decreases that of CH4, due to H2 separation, which favors WGS and methane reforming to carbon dioxide. In the presence of sweep gas, CO2 production exceeds CH4 production. The MSR reaction is more heavily influenced by hydrogen removal
Fig. 6 e Hydrogen recovery and yield as a function of pressure at Tw ¼ 753 K, F ¼ 0.5 NL/min, without (a) and with a sweep flow of 0.5 NL/min (b).
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Fig. 7 e Product selectivity as a function of sweep flow rates at P ¼ 8 bar, Tw ¼ 703 K feed flow rate 0.5 NL/min (a) and 1 NL/ min (b) and comparison with equilibrium values.
compared to the WGS reaction, as the stoichiometric ratio between hydrogen and methane in MSR is four, whereas only 1 mol of H2 is produced for every mole of CO in WGS. The direction of selectivity towards CO is better observed in Fig. 8, where the exit ratio of CO and CH4 mole fraction (which equals that of the respective selectivity ratio) is compared to the closed system equilibrium values. As sweep flow (and consequently the permeate flow) increases, the MSR is enhanced compared to WGS, which in turn increases CO content relative to CH4 by as much as an order of magnitude. The influence of sweep gas is more pronounced when a lower feed flow rate is used, because the higher residence time allows a better hydrogen separation.
Membrane reactor modeling The behavior of the membrane reactor has been described with a one-dimensional pseudo-homogeneous model, with axial dispersion of mass and heat, in which the following main assumptions have been made:
- Ideal gas behavior; - Negligible axial pressure drop; - Negligible radial gradients radial temperature gradients were measured to be small; effects of concentration polarization were estimated to be small, based on an analysis carried out for methane steam reforming in the same reactor [2], using a correction reported in Ref. [21]. The correction depends on the reactor's radius, hydrogen diffusivity and membrane permeance, the conclusions reached in Ref. [2] are therefore extendable to the present work; - Hydrogen permeation through Pd-based membranes involves the steps of dissociative adsorption, subsurface penetration, diffusion through the membrane and the inverse steps on the other side. When the rate-limiting step is hydrogen atom diffusion through the membrane, as in the case of Pd-based membranes down to 1 mm [16], permeation in pure hydrogen follows Sievert's law (eq. (20)). When considering a gaseous mixture, the same law applies with a membrane permeance that depends on coadsorption of reactants and products:
J H 2 ¼ Qm
pffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi PyH2 PSW ySW H2
(20)
where membrane permeance is evaluated as: Em Qm ¼ QAm exp RT
(21)
and Q is a parameter accounting for permeance inhibition due to competitive adsorption of reactants and products.
Model equations The material balance equations are adapted from Ref. [2], applying them to the components and reactions of interest: Fig. 8 e Selectivity ratio between CO and CH4 directed by H2 removal. Experimental measurements of retentate composition compared to the closed system equilibrium values. P ¼ 8 bar, S/E ¼ 3.
εrg
3 vyi 1 vF vyi v2 yi X yi þ ¼ F þ Deff rg 2 þ ai;j rj ð1 εÞrs ; AO vz vt vz vz j¼1
isH2 (22)
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εrg
1 0:5 1:5 pCO2 PCH4 rED ¼ kED pEtOH kED
vyH2 1 vF vyH2 v2 yH2 yH2 þ ¼ F þ Deff rg AO vz vt vz vz2 3 X pffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi pdm þ aH2 ;j rj ð1 εÞrs Qm PyH2 PSW ySW H2 AO j¼1
1 KadsCO KadsH2 O pCO pH2 O KWGS pCO2 pH2
(23)
rg
vySW 1 vFSW SW vySW H2 H2 SW ¼ y þ F AI vt vz H vz
! þ
pffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi pdm Qm PyH2 PSW ySW H2 AI (24)
Effective dispersivity has been considered to be the same for all species, and equal to 1.7 104 m2/s. The following boundary conditions apply: Fjz¼0 ¼ F0 ; FSW z¼L ¼ FSW 0 ; ¼ yCO2 ¼ yCO z¼0 ¼ yH2 ¼ 0; yCH4 z¼0 z¼0 z¼0 1 S=E ; yH2 O ; ySW yEtOH z¼0 ¼ ¼ ¼0 H2 z¼0 z¼L 1 þ S=E 1 þ S=E
vT FH2;perm vT v2 T F0 0 ¼ kax:eff 2 cpg cp;H2 eff vt vz vz A F0 þ rs ð1 εÞ
3 X
DHj rj hw
j¼1
pD ðT Twall Þ A
(25)
vT vz
accounts for heat loss from the gas flow associated with hydrogen permeation. The boundary conditions are: vT Ng Cpg T0 Tjz¼0 ¼ kax;eff vz z¼0
1 KCO2 M
pCH4 p2H2 O
(28)
The adsorption constants KCO, KCO2 , and KH2 O have been considered constant and equal to their average value in the range 300e500 C, as suggested by Palma et al. [5] in their kinetic studies. The pre-exponential factors and activation energies of the kinetic constants are reported in Table 6 [5]. The effective axial conductivity has been evaluated through the following correlation:
jH ¼
where hw is wall to catalyst heat transfer coefficient, Tw is the controlled wall temperature, and FH2 ;perm is the hydrogen permeation rate. The term:
FH2 ;perm cp;H2
rCO2 M ¼ kCO2 M pCO2 p4H2
(27)
(29)
where ks is the intrinsic conductivity of the solid, equal to 50 W/mK and ε is the bed porosity, equal to 0.85. The wall heat transfer coefficient has been evaluated through a ChiltoneColburn type correlation:
The heat balance is: rcp
rWGS ¼ kWGS 2 1 þ KadsCO pCO þ KadsH2 O pH2 O þ KadsCO2 pCO2
1 kax;eff ¼ ks ð1 εÞ 3
vyi ¼0 vz z¼L
(26)
vT ¼0 vz z¼L
Although, axial mass dispersion is usually ignored in membrane reactor studies, the incorporation of axial conductivity is essential for this non-isothermal system. A non-stationary model has been developed to simplify the numerical solution procedure of this boundary-value problem. The system of partial differential equations has been solved using the PDE Toolbox of Matlab®.
Model parameters The kinetic expressions used in the model for the ethanol decomposition (Eqn. (10)), wateregas-shift (Eqn. (2)) and methanation (Eqn. (13)) reactions are those reported by Palma et al. [5], obtained by studying the kinetic activity of the catalyst used in the present work in an ethanol steam reforming reactor operating at different contact times between temperatures of 523 and 873 K. The rate equations have been re-written here for clarity:
Nuw Pr1=2 Re
¼ aReb1
(30)
A thermal characterization of the solid foam has been previously carried out, and the value of the parameter b has been determined to be 0.29 [17]. The value of a has been obtained in the course of the present work by fitting experimental data and is equal to 24.1. A complete list of parameters used in the model, along with their meaning and value, is reported in the List of Symbols section at the end of the paper.
Model results Fitting the experimental results presented here (conditions summarized in Table 2) with the model described above, using thermodynamic and kinetic parameters from the literature, and only two adjustable parameters, has shown good agreement (Fig. 9). However, the estimated permeance at 723 K is about 12 Nm3h1m2bar0.5. This value should be compared with the permeance of the membrane used in the present work that was measured by the producer (ECN) to be about 75 Nm3h1 m2 bar0.5 at 723 K in pure hydrogen (Em ¼ 11 kJ/ mol and Am ¼ 5.6 mol/(m2sbar0.5)). Hydrogen flow in pure hydrogen and mixtures was found to follow Sievert's law, indicating that the limiting step in its permeation is the diffusion of atoms through the bulk of the membrane. This is a
Table 6 e Pre-exponential factors and activation energies of the kinetic constants. k0
Reaction Modified ethanol decomposition Water gas shift CO2 methanation
4
7.94 10 mol/(gcat s bar) 2.77 108 mol/(gcat s) 7.65 101 mol/(gcat s bar4)
Ea [kJ/mol] 104 74 8
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while the concentration of other components is higher. This phenomenon is known as concentrationpolarization. In narrow reactors, such as the one used in the present study, the effect of concentrationpolarization is not negligible, but cannot account for the large decline (see below); (ii) Co-adsorption of reactants and products on the Pdsurface may lead to inhibition: While there are many studies of the inhibiting effects of CO [18] the effects of methane adsorption is contradictory, while computational (DFT) studies show negligible effects due to CO2, CH4 or H2O adsorption [19]. The relatively small concentrations of CO make this explanation questionable as well; (iii) Surface reactions (such as reverse-WGS) may cause adsorption of CO on the Pd surface.
Fig. 9 e Comparison between experimental (points) and calculated (lines) molar fractions on dry basis as a function of temperature at P ¼ 9 bar, 0.5 NL/min feed flow rate, sweep gas flow rate 0 NL/min (a), 0.5 NL/min (b), and 1 NL/ min (c). significant inhibition (Q ¼ 0.15.) and we try to understand its sources. No attempt was made to fit the data using an inhibition Q function that depends on composition, as that would involve adjusting too many parameters. Furthermore, gas composition in proximity of the entrance section is not known precisely. The decrease in apparent permeance may be due to several reasons: (i) In wide reactors the incorporation of high-permeance membranes leads to the hydrogen concentration at the membrane wall to be much lower than in the bulk,
An inhibition effect similar to the one noticed here was previously observed in a study of methane steam reforming in the same membrane reactor [2] and an attempt was made to discriminate between various sources of inhibition. It was argued that the strong inhibition cannot be explained only by concentration polarization or by competitive adsorption by gas-phase species. The former effect was ruled out by approximate criteria and corrections to the 1-D model accounting for transversal effects, which were developed for hydrogen separation without [20] or with [21] reaction. The latter effect was ruled out using partial information on adsorption properties from the literature: CO inhibition is strong but its concentration is small. On the other hand, the inhibition can be accounted for by competitive adsorption by species produced on the surface. This issue is hardly addressed in the literature, but some indications do exist: Li et al. [22] found that both CO and H2O reduce hydrogen permeation through a Pd/stainless steel membrane at 653 K. Permeation tests with binary H2/CO and H2/H2O mixtures of the same concentrations of CO and H2O showed that steam had a stronger inhibition effect than CO on hydrogen permeation. Hulme et al. [23] found significant decline of hydrogen permeability through a PdeAg/VeNi alloy membrane in presence of CO2 at 473e673 K (60% decline with 30% CO2 and stronger at lower temperatures). These results can only be explained by higher CO surface concentrations than those expected from gas phase composition. Fig. 10 shows the calculated temperature profile along the reactor axis when the wall temperature has been set at 753 K and the reactor operates at 8 bar with a 1 NL/min feed flow rate and 1 NL/min sweep gas flow rate. The inlet temperature was equal to 573 K: the calculated temperature increases rapidly in proximity of the inlet, due to the combined effect of the exothermicity of the ethanol decomposition reaction and the heat transfer from the reactor wall. Temperature tends to decrease towards the final part of the reactor. This behavior may be attributed to the fact that when the sweep gas is sent counter-currently with respect to the reactor flow, the hydrogen partial pressure in the permeate side is close to zero in correspondence of the exit section of the reactor. Hydrogen permeance through the membrane shifts the carbon dioxide methanation reaction backwards, absorbing heat and causing a reduction in temperature.
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providing the structured foam, and the Energy Research Centre of the Netherlands (ECN), for providing the membrane. The experimental equipment was partially supported by Technion Grand Energy Program (TGEP).
List of Symbols a A0 AI Am
Fig. 10 e Calculated (continuous) and measured (points) temperature along reactor axis for Tw ¼ 753 K, P ¼ 8 bar, 1 NL/min feed flow rate, sweep gas flow rate 1 NL/min.
Conclusions The present work showed that in a system of competing reversible reactions, the desired one can be accelerated by separating one of its products, hydrogen in this case. This novel idea has the potential of suggesting new designs for improving selectivity. It was shown experimentally that pure hydrogen can be produced by ethanol steam reforming in an integrated catalytic membrane reactor, with complete conversion, good selectivity, and with no appreciable catalyst deactivation. As explained in the Introduction a membrane reformer heated by solar energy through molten salt represents an interesting process for renewable hydrogen production, which can be converted to electricity via FC. The possibility of working with low amounts of water and relatively low pressures reduces the energetic cost of conducting ethanol steam reforming. Using steam instead of nitrogen as sweep gas will also make the overall process easier, by allowing separation of the permeate side current by simple condensation of water. Process optimization can proceed with the mathematical model that was developed; the model uses only two adjustable parameters: a permeance inhibition factor, which was found to be significant, and the ChiltoneColburn coefficient of the wall heat transfer coefficient. Understanding the drop in permeance will allow to reduce the membrane (and reactor) size and cost.
Acknowledgments This work has been conducted within the framework of the CoMETHy (Compact Multifuel-Energy to Hydrogen converter) project, funded by the European Commission under the Fuel Cells and Hydrogen Joint Undertaking (GA No. 279075). The authors wish to acknowledge the Fraunhofer Institute for Ceramic Technologies and Systems, IKTS (Germany), for
b dm dw Deff Ea Em F FH2 ;perm jH kax,eff kg ks k0 KadsCO KadsCO2 KadsH2 O KCO2 M KED KWGS Ng Nuw P Pr Qm rj Re S/E T u y
coefficient of ChiltoneColburn correlation (24.1, from data fitting), annular section (1.1 103), m2 membrane cross section (6.24 105), m2 membrane permeance pre-exponential factor (6.6), mol/m2/s/bar0.5 exponent of ChiltoneColburn correlation (0.29), outer membrane diameter (0.014), m outer reactor diameter (0.04), m effective dispersivity (1.7 104), m2/s rate constant activation energy, kJ/mol membrane permeance activation energy (11), kJ/mol gas flow rate (evaluated from model), mol/s hydrogen permeation rate (evaluated from model), mol/s ChiltoneColburn coefficient (evaluated from model), effective axial conductivity (2.5), W/m/K gas conductivity (6.9 102), W/m/K intrinsic solid conductivity (50), W/m/K pre-exponential factor of the rate constants, mol/ gcat s bar CO constant (0.37), atm1 CO2 adsorption constant (1.701), atm1 H2O adsorption constant (80073.2), atm1 CO2 methanation reaction equilibrium constant ethanol decomposition reaction equilibrium constant water gas shift reaction equilibrium constant specific gas flow rate (evaluated from model), mol/m2/s wall Nusselt number (evaluated from model), pressure, bar Prandtl number (Pr ¼ cp,gmg/kg), membrane permeance (evaluated from model), mol/ m2/s/bar0.5 j-th reaction rate (evaluated from model), mol/gcat/s Reynolds number for a packed bed (Re ¼ rgudw/mg), steam:ethanol molar ratio at inlet, temperature, K gas superficial velocity (evaluated from model), m/s molar fraction (evaluated from model),
Greek letters ε bed porosity (0.85), gas phase molar density (evaluated from model), mol/ rg m3 rs catalyst density (1.9), g/mL a stoichiometric coefficient, Q inhibition parameter (0.15, from data fitting), Subscripts i i-th component j j-th component w wall m membrane
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s g
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solid gas
Superscripts SW sweep side
references
[1] Giaconia A, Turchetti L, Monteleone G, Morico B, Iaquaniello G, Shabtai K, et al. Development of a solarpowered, fuel-flexible compact steam reformer: the CoMETHy project. Chem Eng Trans 2013;35:433e8. [2] Patrascu M, Sheintuch M. On-site pure hydrogen production by methane steam reforming in high flux membrane reactor: experimental validation, model predictions and membrane inhibition. Chem Eng J 2015;262:862e74. [3] Frusteri F, Freni S, Spadaro L, Chiodo V, Bonura G, Donato S, et al. H2 production for MC fuel cell by steam reforming of ethanol over MgO supported Pd, Rh, Ni and Co catalysts. Catal Commun 2004;5:611e5. [4] Mattos LV, Noronha FB. Hydrogen production for fuel cell applications by ethanol partial oxidation on Pt/CeO2 catalysts: the effect of the reaction conditions and reaction mechanism. J Catal 2005;233:453e63. [5] Palma V, Castaldo F, Ciambelli P, Iaquaniello G. CeO2supported Pt/Ni catalyst for the renewable and clean H2 production via ethanol steam reforming. App Cat B 2014;145:73e84. [6] Furtado AC, Gonc C, Alonso CG, Cantao MP, FernandesMachado NRC. Bimetallic catalysts performance during ethanol steam reforming: influence of support materials. Int J Hydrogen Energy 2009;34:7189e96. [7] Vizcaino AJ, Carrero A, Calles JA. Ethanol steam reforming on Mg- and Ca-modified CueNi/SBA-15 catalysts. Cat Today 2009;146:63e70. [8] Orucu E, Gocalimer F, Aksoylu AE, Onsan ZI. Ethanol steam reforming for hydrogen production over bimetallic PteNi/ Al2O3. Catal Lett 2008;120:198e203. [9] Palma V, Castaldo F, Ciambelli P, Iaquanilello G, Capitani G. On the activity of bimetallic catalysts for ethanol steam reforming. Int J Hydrogen Energy 2013;38:6633e45. [10] Domı`nguez M, Taboada E, Molins E, Llorca J. Ethanol steam reforming at very low temperature over cobalt talc in a membrane reactor. Cat Today 2012;193:101e6.
[11] Basile A, Gallucci F, Iulianelli A, Tosti S. CO-free hydrogen production by ethanol steam reforming in a PdeAg membrane reactor. Fuel Cells 2008:62e8. [12] Tosti S, Basile A, Borgognoni F, Capaldo V, Cordiner S, Di Cave S, et al. Low temperature ethanol steam reforming in a PdeAg membrane reactorepart 1 Ru-based catalyst. J Membr Sci 2008:250e7. [13] Tosti S, Basile A, Borgognoni F, Capaldo V, Cordiner S, Di Cave S, et al. Low-temperature ethanol steam reforming in a PdeAg membrane reactorepart 2. Pt-based and Ni-based catalysts and general comparison. J Membr Sci 2008;308:258e63. [14] Fatsikostas AN, Verykios XE. Reaction network of steam reforming of ethanol over Ni-based catalysts. J Catal 2004;225:439e52. [15] Turchetti L, Monteleone G, Annesini MC. Comparative analysis of heat transport in SiC solid foam and Al2O3 granular packings for fixed-bed reactors. Chem Eng Trans 2011;24:1405e10. [16] Ward TL, Dao T. Model of hydrogen permeation behavior in palladium membranes. J Membr Sci 1999;153:211e31. [17] Colonna C. Caratterizzazione e modellizzazione dei di moto meccanismi di trasporto di calore e quantita attraverso supporti catalitici di schiuma solida [Unpublished master's thesis]. Department of Chemical Engineering, Materials and Environment. University of Rome “Sapienza”; 2012. [18] Israni SH, Harold MP. Methanol steam reforming in singlefiber packed bed PdeAg membrane reactor experiments and modeling. J Membr Sci 2011;369:375e87. [19] Abir H, Sheintuch M. Modeling H2 transport through a Pd or Pd/Ag membrane, and its inhibition by co-adsorbates, from first principles. J Membr Sci 2014;466:58e69. [20] Nekhamkina O, Sheintuch M. Effective approximations for concentration-polarization in Pd-membrane and separators. Chem Eng J 2015;260:835e45. [21] Sheintuch M. Pure hydrogen production in a membrane reformer demonstration, macro-scale modeling and atomicscale modeling. Chem Eng J 2014. http://dx.doi.org/10.1016/ j.cej.2014.11.100. [22] Li A, Liang W, Hughes R. The effect of carbon monoxide and steam on the hydrogen permeability of a Pd/stainless steel membrane. J Membr Sci 2000;165:135e41. [23] Hulme J, Komaki M, Nishimura C, Gwak J. The effects of gas mixtures on hydrogen permeation through PdeAg/VeNi alloy composite membrane. Curr Appl Phys 2011;11:972e5.
Available online at www.sciencedirect.com
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Directing selectivity of ethanol steam reforming in membrane reactors Maria Anna Murmura b, Michael Patrascu a, Maria Cristina Annesini b, Vincenzo Palma c, Concetta Ruocco c, Moshe Sheintuch a,* a
Technion-Israel Institute of Technology, Department of Chemical Engineering, Haifa 32000, Israel University of Rome “Sapienza”, Department of Chemical Engineering, Materials and Environment, Via Eudossiana 18, 00184 Rome, Italy c University of Salerno, Dipartimento di Ingegneria Industriale, Via Giovanni Paolo II 132, 84084 Fisciano, SA, Italy b
article info
abstract
Article history:
In a system of parallel reversible reactions, separating the corresponding product can enhance
Received 21 January 2015
a desired reaction. If the same product is produced in several reactions, its concurrent sepa-
Received in revised form
ration enhances the reaction with the higher stoichiometry. Here we demonstrate this effect
26 February 2015
by separating hydrogen from ethanol during steam reforming in a Pd membrane reactor
Accepted 5 March 2015
packed with a Pt/NieCeO2 catalyst. Interest in ethanol SR stems from the need to produce
Available online 30 March 2015
ultra-pure H2 from renewables (the energy will be supplied by solar-heated molten salt).
Keywords:
H2, CO2, CH4 and CO as products. These products can be represented by three reactions
For the conditions tested full conversion (of reaction (1) below) has been achieved with Catalytic membrane reactor
(W ¼ H2O): 1. EtOH 4 CH4 þ CO þ H2 (ethanol decomposition), 2. CH4 þ 2W 4 CO2 þ 4H2
Reforming
(methane steam reforming, MSR) and 3. CO þ W 4 CO2 þ H2 (water gas shift, WGS).
Hydrogen
Separating H2 directs the selectivity towards CO and CO2, resulting also in increased ratio
Ethanol
between CO and CH4 mole fraction.
Modeling
In general, increasing temperature (613e753 K), pressure (6e10 bar) and introducing sweep flow (0.5 NL/min N2 for a similar feed rate) led to better separation, to an increase in selectivity towards CO and CO2 and in hydrogen yield. Increasing pressure and introducing sweep flow also increased hydrogen recovery. A further increase of sweep gas flow rate (to 1 NL/min) did not result in an appreciable improvement. The results of this work show that the combination of theNi/Pt catalyst and the Pd membrane for hydrogen removal produce very high values of hydrogen yield, despite the low steam to ethanol ratio and the moderate pressure levels. In particular about 4.5 mol H2/ mol EtOH were produced at 753 K, feed and sweep flow rates of 0.5 NL/min each in the entire pressure range examined. A one-dimensional model based on the kinetics of a limited but realistic set of reactions has been developed to simulate the behavior of the reactor. With a literature kinetics model and only two adjustable parameters, the permeance and heat transfer coefficient, a very good agreement with experimental data has been obtained. The results indicate strong permeance inhibition compared to pure hydrogen measurements. Copyright © 2015, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved.
* Corresponding author. Tel.: þ972 4 8202823. E-mail address: [email protected] (M. Sheintuch). http://dx.doi.org/10.1016/j.ijhydene.2015.03.013 0360-3199/Copyright © 2015, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved.
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CO þ H2O $ CO2 þ H2
Introduction The worldwide hydrogen demand is increasing, and its use in Fuel Cells (FCs) is expected to make hydrogen one of the widely-used fuels in a future energy system, mainly thanks to its low environmental impact when used in FC systems. On the other hand, the extensive introduction of hydrogen as energy vector is limited by the actual lack (and costs) of a reliable hydrogen distribution infrastructure. This barrier can be surmounted by the development of systems for decentralized hydrogen production (i.e. close to the end-user). Fuelflexibility and the possibility to power the process with renewables (solar and biomass) are two additional preconditions for sustainable hydrogen production. The above considerations are at the basis of the CoMETHy project (Compact Multifuel-Energy to Hydrogen converter). CoMETHy's general objective is to support the intensification of hydrogen production processes, by developing a membrane reformer that can operate with various fuels and that eventually will be heated by solar-heated molten salt [1]. An electrically-heated lab membrane reformer was tested in the Technion with methane SR [2], and this work extends that study to ethanol SR using the same catalyst, membrane and system. Hydrogen production from biomass derived oxygenates has attracted interest for its potential application in fuel cells. Bio-fuels for the production of hydrogen could bring significant environmental benefits, as the carbon dioxide produced is consumed in biomass production giving a CO2 neutral energy supply. Steam reforming of bioethanol has been widely investigated over supported transition and noble metal catalysts [3,4]. Hydrogen production by SR of ethanol often encounters several problems, which are addressed in the CoMETHy project: - High-purity hydrogen is required for its application in fuel cells: the most promising FCs, based on PEM, require low levels of CO; hydrogen separation by a Pd membrane should satisfy this condition. - Steam reforming is highly endothermic and requires heat supply; in this project the reformer is heated by solar energy. - The reaction is accompanied by the formation of various by-products, which greatly affect the selective production of hydrogen; as we show in this article, hydrogen separation will shift the product distribution in the desired direction, i.e., toward the reaction that produces most hydrogen. - The reaction is limited by equilibrium; hydrogen separation will shift the equilibrium conversion. The ethanol steam reforming (ESR) reaction is: C2H5OH þ H2O $ 2CO þ 4H2
(1)
The conversion of CO to CO2 is considered through the wateregas shift (WGS) reaction:
(2)
The overall ESR reaction, which refers to complete conversion of CO to CO2, is therefore:
C2H5OH þ 3H2O $ 2CO2 þ 6H2
(3)
Aside from these desired reactions, several others may take place [5].
C2H5OH $ C2H4O þ H2
(4)
C2H4O $ CH4 þ CO
(5)
C2H4O þ H2O $ 2CO þ 3H2
(6)
C2H5OH $ C2H4 þ H2O
(7)
2C2H5OH $ C3H6O þ CO þ 3H2
(8)
C2H5OH þ 2H2 $ 2CH4 þ H2O
(9)
C2H5OH $ 0.5CO2 þ 1.5CH4
(10)
CH4 þ H2O $ CO þ 3H2
(11)
2CO $ CO2 þ C
(12)
CO2 þ 4H2 $ CH4 þ 2H2O
(13)
Achieving the desired product is usually obtained by a proper choice of catalyst and support (see below) and operating conditions. Selectivity may be further directed through the use of membrane reactors, which shift the equilibrium in the direction of the selectively permeating component (hydrogen in this case). Recently, catalysts containing more than one active species have also been investigated because of their significantly different catalytic properties with respect to either of the parent metals [6,7]. A synergic effect of Pt addition to Ni-based catalysts was shown to improve the activity of the non-noble metal towards hydrogenation reactions, resulting in decreased coke formation rates with respect to the ones observed over monometallic Ni-catalysts [8]. When Pt and Ni are supported on CeO2, by depositing the non-noble metal on the support surface earlier than the noble one, as Pt is directly available at the gasesolid interface, ethanol adsorption is favored and a better agreement with thermodynamic data in a regular fixed bed was observed [9].
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Table 1 e Summary of experimental conditions and results previously reported in literature. Ref
T [K]
P [bar]
S/E
Sweep to feed flow rate ratio
Inlet flow rate [molEtOH/kgcatmin]
[10] [11] [12] [13]
598e673 573e673 673e723 673e723
5e14 1.3 1.5e2 1.5e2
6 3e9 8.4e13 8.4e13
0 0.5e1.4 0.9e3.7 0.9e3.7
1.67 0.83 0.04e0.64 0.04e1
Conversion and product distribution in ESR in membrane reactors (MRs) depend on various operating conditions: - Increasing temperature will increase hydrogen partial pressure and its permeating flux, leading to higher conversion and selectivity. - While increasing pressure has detrimental effect on conversion in a closed system at equilibrium, it will have an advantageous effect on conversion when significant hydrogen separation is achieved, as follows from stoichiometry. - At low temperatures, increasing the S/C ratio (to avoid coking) leads to an increase in H2 and CO2 selectivity, while decreasing CO and CH4 selectivity. Hydrogen recovery decreases due to dilution effects with increasing S/C ratios. Table 1 summarizes the main operating parameters and results of ethanol steam reforming in membrane reactors previously reported in literature. Domı`nguez et al. [10] used a PdeAg membrane for ESR on cobalt talc (Co3[Si2O5]2(OH)2). Studies conducted on a PdeAg membrane reformer with a Rubased catalyst at 1.3 bar [11] showed that increasing temperature shifts the reactions toward the production of CO2 and CH4. In a study over a metal foil membrane, using Ru, Pt, Ni based catalyst [12,13] positive effect of increasing the retentate pressure was observed. In these last studies the effect of co-current vs. counter-current sweep flow was also investigated. None of these works included the development of a model to describe the experimental results obtained. In the present work, experimental data collected on lowtemperature (753 K) ethanol steam reforming in a membrane reactor, equipped with a high permeance Pd membrane, are presented and analyzed, with a particular focus on the operating conditions which maximize selectivity towards carbon dioxide and hydrogen. Full ethanol conversion is achieved under conditions studied and the problem is reduced to competition between carbondioxide and methane production. High hydrogen yields and recoveries have been obtained even when working with stoichiometric steam to ethanol ratios (S/E ¼ 3), moderate pressures and with low ratios of sweep gas to feed gas flow rates. Previous works (Table 1) used high S/E ratios (6e13) and obtained high conversion due to dilution. No deactivation was detected even when running with a stoichiometric S/C ratio. This work, in a high permeance membrane, also presents better selectivities and hydrogen separation, and higher throughputs, than those presented in previous works. Furthermore, the work developed and tests a simple model that described well the results and can be used for further optimization; it shows that results are insensitive to catalytic activity.
EtOH conversion [%] 98.6e100 16e100 100 100
H2 yield 3.2e3.7 0.6e3.3 0.2e5 0.3e3.6
The structure of the paper is as follows: thermodynamic equilibrium analysis of a limited but realistic set of reactions is presented (Section Thermodynamic analysis), followed by a description of the experimental system and results obtained (Section Experimental activity). A one-dimensional model, developed to simulate the behavior of the reactor, is also presented and analyzed in Section Membrane reactor modeling.
Thermodynamic analysis Thermodynamic analysis will provide a limit to the system operation. Kinetic tests previously carried out on the catalyst formulation employed in the present work have shown that the system may be fully described by considering the following reactions: ethanol dehydrogenation (4), acetaldehyde decomposition (5), modified ethanol decomposition (10), acetaldehyde steam reforming (6), wateregas shift (2), and CO2 methanation (13) [5]. Since no acetaldehyde formation was noticed in the course of the present work, reactions (4) and (5) can be combined and (6) may be disregarded. Of the four remaining reactions only three are linearly independent, the system may therefore be studied by considering only three reactions. Table 2 summarizes the reactions considered, along with their DH0 and DG0. The same three reactions can be represented in other linear-dependent ways, leading to same results. The equilibrium conditions of the three reactions considered yield: KEDðTÞ ¼
pCO pCH4 pH2 yCO yCH4 yH2 ¼ P2TOT pEtOH yEtOH
(14)
KMSRðTÞ ¼
pCO2 p4H2 yCO2 y4H2 ¼ P2TOT 2 p2H2 O pCH4 yH2 O yCH4
(15)
KWGSðTÞ ¼
pH2 pCO2 yH yCO2 ¼ 2 pCO pH2 O yCO yH2 O
(16)
A preliminary thermodynamic analysis maps hydrogen partial pressure in a closed system as a function of total pressure and temperature (Fig. 1, S/C ¼ 3) while considering the three reactions mentioned above. Hydrogen separation occurs if its partial pressure in the reactor side exceeds that in the permeate side. Increasing pressure leads to a less than linear increase of hydrogen partial pressure, because hydrogen concentration decreases while the total pressure increases. To carry out low-temperature ESR efficiently, taking advantage of the entire membrane surface, it is therefore essential to consider the use of sweep gas in the permeate side.
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Table 2 e Reactions accounted for in ethanol steam reforming.
1 2 3
Reaction name
Reaction
DH0 [kJ/mol]
DG0 [kJ/mol]
Ethanol decomposition (ED) Methane steam reforming (MSR) Water gas shift (WGS)
C2H5OH $ CH4 þ CO þ H2 CH4 þ 2H2O(g) $ CO2 þ 4H2 CO þ H2O(g) $ CO2 þ H2
49.4 164.6 41.2
355.6 113.3 28.8
Assuming reaction (1) goes to completion the equilibrium selectivity ratio (s.r.) between CO and CH4 can be expressed as: If H2 is s:r:≡yCO =yCH4 ¼ KMSRðTÞ =ðKWGSðTÞ P2TOT Þ$ðyH2 O =y3H2 Þ. removed during reaction it will direct the selectivity towards CO and CO2 but the ratio yCO/yCH4 will increase by increasing the conversion of MSR to a greater extent than the conversion of WGS reactions. This is evident by the (yH2)-3 dependence in the above expression. Recall that yH2O also varies, and in a stoichiometric feed (as in this work), yH2O ~ yCO þ 2yCH4. Considering that a different subset of reactions may take place in correspondence of the inlet section (e.g., EtOH þ 3H2O $ 2CO2 þ 6H2, which goes to complete conversion) suggests that the hydrogen driving force will allow for some transport across the membrane even at temperatures lower than 700 K in the absence of sweep.
Experimental activity Experimental setup Experimental activity, designed to test ethanol steam reforming in a laboratory scale membrane reactor, has been carried out in the Chemical Reactor Engineering and Environmental Catalysis Laboratory of the Technion. The system employed is the one used for methane SR [2] after some modifications. The 40 cm long membrane reactor (Fig. 2a) consists of a tube-and-shell structure. Reactions take place in
Fig. 1 e Equilibrium partial pressure of hydrogen (in bar) as a function of temperature and pressure for feed of S/E ¼ 3. Gray area indicates feasible conditions for non-sweep operation assuming equilibrium compositions.
the shell side, packed with a structured foam catalyst to expedite radial heat transfer. Hydrogen permeates through the Pd membrane (the inner reactor wall, having a 14 mm outer diameter) into the tube-side, where a sweep gas may flow (counter-currently). Reactor characteristics have been summarized in Table 3. In the present work, nitrogen was used as a sweep gas; however, it is likely that steam will be used in the pilot scale application. This will allow an easy separation of hydrogen from the permeate stream. The stainless-steel reactor is heated from the outside (40 mm outer diameter) by four concentric electric heaters, shown as transparent bodies in Fig. 2a. Each heater is controlled independently to maintain a set temperature, measured on the outer wall of the reactor, in the midpoint of the heater body. A fifth heating element is placed at the entrance of the reactor. This last heating element acts as a vaporizer and its temperature is set to 300 C throughout all the experiments. Four additional thermocouples are located inside the membrane, in the same axial position as the outer control thermocouples. Pipeline elements were sealed using Hamlet® Let-Lok Tube Fittings (Double Ferrule). The effluent from the reactor is cooled, to allow the condensation of unreacted ethanol and steam, and sent to a silica gel column, where any residual vapor is absorbed. The dried effluent is then sent to an analyzing system, comprised of IR analyzers for CH4, CO2, and CO (Edinburgh Instruments Ltd.) and a gas chromatograph (Trace GC Ultra, ThermoScientific). Either the permeate or the retentate may be sent to the analyzing system. The liquid product has been collected and analyzed at the end of each experiment, showing that it only contained unreacted water. This indicates that full ethanol conversion has been achieved and that no liquid byproducts have been formed. This result is in line with the observations of Palma et al. [5], who showed that with the catalyst tested, acetaldehyde is present only for very short contact times and ethanol is fully converted almost immediately. The reactor is packed with structured foam catalyst, specifically prepared in the framework of the CoMETHy project from cooperative work of the Fraunhofer Institute for Ceramic Technologies and Systems, IKTS (Germany) and the ProCEEDlab of the University of Salerno (Italy). The structured catalyst is made of a CeO2-washcoated SiC open cell foam (see Fig. 2b), prepared by IKTS and activated with the bimetallic (Pt(3)Ni(10)) formulation by the University of Salerno. This formulation was specifically selected on the basis of the considerations mentioned in Section Introduction. Ni and Pt catalysts promote ethanol steam reforming when supported on ceria, which is expected to diminish deactivation due to coking and promote catalyst dispersion [14]. SiC open cell foams have been chosen as the structured carried for
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Fig. 2 e a) Detailed drawing of the membrane reactor; b) Structured foam.
Table 3 e Reactor characteristics. Bed and membrane length [m]
Reactor diameter [m]
Inner reactor diameter (membrane outer diameter) [m]
Mass of catalyts þ structural support [g]
34 103
14 103
200
0.4
their good heat transfer characteristics, and to minimize pressure drops [15]. The membrane tube (Hysep©-technology), provided by the Energy Research Centre of the Netherlands (ECN), is made of a 4e5 mm palladium layer, deposited by electroless plating on the outside of a porous alumina tube provided with two thin porous alumina layers to decrease the outer pore dimension to less than 0.15 mm. The membrane is capped with ECN compression seals on both sides. Membrane permeance was measured by the producer in pure hydrogen at various temperatures and pressures. The permeance varied with temperature according to an Arrhenius-type law (Em ¼ 11 kJ/mol, Am ¼ 5.6 mol/(m2 s bar0.5), see Eq. (21)) and hydrogen flux through the membrane followed Sievert's law. The experimental conditions tested in the present work have been summarized in Table 4.
Catalyst preparation Starting from the washcoated foam, the active species deposition (Pt and Ni) was performed at University of Salerno by multiple cycles of wet impregnation with an aqueous solution of the metal precursor (C4H6O4Ni∙4H2O or PtCl4, both from Aldrich), at 60 C for 1 h; the sample was then dried at 120 C for 2 h and calcined for 1 h at 400 C for Ni deposition and at 550 C for Pt deposition. The two different temperatures were selected on the basis of thermo-gravimetric analysis results. This impregnation-drying-calcination procedure was
repeated the number of times necessary to reach the desired amount of metal (3 wt% for Pt and 10 wt% for Ni), considering the solubility of the precursor salt. After the above cycles, the sample was calcined at 600 C for 3 h, to stabilize it. The whole procedure (impregnation-drying-calcination steps and final calcination) is repeated for each active species to be deposited on the washcoated foam.
Experimental results In all experiments conducted here complete conversion of ethanol was achieved with CO, CO2, CH4, and H2 as the only products. Permeate purity has been measured, and impurities, mainly methane and carbon dioxide, were always found in concentrations < 0.5% ( ~ 0.2% CH4, ~ 0.3% CO2). These measurements could be partly affected by the presence of residues in the pipeline, which was used to measure the retentate current as well. A summary of the experimental conditions and main results obtained has been reporte in Table 5. Experimental results have been expressed in terms of hydrogen recovery and yield: Hrec ¼
H2 permeated total H2 produced
Hyield ¼
total H2 produced ethanol in feed
(17)
where the maximum values obtainable are equal to 1 and 6, respectively; and in terms of selectivity towards carboncontaining species: Si ¼
moles of i produced 100 i ¼ CO2 ; CO; CH4 2ðmoles of ethanol convertedÞ
(18)
Table 4 e Experimental conditions. T [K] 613e753
Retentate pressure [bar]
Permeate pressure [bar]
Total feed flowrate [NL/min]
Sweep flowrate [NL/min]
S/E
6e10
1
0.5e1
0e1
3
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Table 5 e Summary of experimental conditions and results in present work. T [K] 613e753
P [bar]
S/E
Sweep to feed flow rate ratio
Inlet flow rate [molEtOH/kgcatmin]
EtOH conversion [%]
H2 yield
6e10
3
0e2
0.03e0.05
100
0.3e4.8
If complete ethanol conversion is assumed, the selectivity of each species may be determined by: Si ¼
yi 100 i ¼ CO2 ; CO; CH4 yCO2 þ yCO þ yCH4
(19)
In the figures that follow, product selectivity has been compared to its value at equilibrium calculated based on feed composition and outlet temperature. Selectivity to CO was found to be small under all conditions. Better separation of hydrogen, achieved by increasing temperature, decreasing flow rate and increasing sweep gas flow rates, will shift the reaction toward CO2 and suppress selectivity to methane. Hydrogen separation may lead to faster coking by reactions like (12). In this study catalytic activity seems to have remained stable during the 200 h of operation with ESR.
Influence of temperature In the absence of sweep flow, hydrogen separation is small and the product distribution is similar to that at equilibrium (Fig. 3a). Increasing temperature leads to small increase in selectivity towards CO and CO2 and a decrease in selectivity towards CH4, mainly due to higher methane steam reforming and reverse WGS conversions. This situation also leads to an increase in hydrogen yield (Fig. 4a). Deviations from equilibrium become significant only above 703 K, but in this range conversion to CO cannot be ignored. Hydrogen recovery decreases with increasing temperature, despite increasing membrane permeance (Fig. 4a). This occurs because hydrogen production increases, but its flow across the membrane is limited by the high partial pressure of hydrogen in the permeate. When a sweep flow rate of 1 NL/min is used (Fig. 3b), hydrogen separation is more significant, and selectivity towards CO2 and CO increases with increasing temperature reaching 80% and 4%, respectively, at 750 K, whereas
selectivity towards CH4 decreases to 15%. The positive effect of hydrogen separation is evident in Fig. 4b: in the entire temperature range hydrogen recovery is very close to 1, and the yield increases steadily with temperature. Under these conditions, the partial pressure of hydrogen in the permeate is reduced, allowing for efficient separation and significant deviation from equilibrium.
Influence of pressure Increasing pressure may have contrasting effects on the hydrogen-producing reactions, and therefore on the overall product distribution: the ethanol steam reforming process results in an increase in the number of moles, low pressures therefore shift the equilibrium conversion towards the products; however, a high retentate to permeate pressure gradient improves hydrogen separation. As separation is promoted the equilibrium of the reaction moves further towards the products. Fig. 5 shows the influence of pressure on the selectivity of carbon-containing species at 753 K, with a feed flow rate of 0.5 NL/min, and for two sweep gas flow rates. In the absence of sweep flow (Fig. 5a), increasing pressure leads to an increase in CO2 selectivity and a decrease in CH4 selectivity, indicating that permeation has a more significant effect than thermodynamics. Under these conditions, hydrogen recovery indeed increases with pressure, as may be seen by the results shown in Fig. 6a. The same may be said for the total hydrogen yield (Fig. 6a), which is favored by the removal of hydrogen itself. A similar trend in CO2 and CH4 selectivities is observed when working with sweep gas, even though the effect of pressure appears to be much less significant. Fig. 6b shows that under these conditions, hydrogen recovery is very close to 1 even when working at 6 bar, indicating that hydrogen removal has already reached its maximum value. The total hydrogen yield is also unaffected by pressure (Fig. 6b), which indicates that the system has reached its maximum performance and
Fig. 3 e Product selectivity as a function of temperature at P ¼ 8 bar, F ¼ 0.5 NL/min, without (a) and with a sweep flow of 1 NL/min (b), and its comparison with equilibrium values.
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Fig. 4 e Hydrogen recovery and yield as a function of temperature at P ¼ 8 bar, F ¼ 0.5 NL/min, without (a) and with a sweep flow of 1 NL/min (b).
Fig. 5 e Product selectivity as a function of pressure at Tw ¼ 753 K, F ¼ 0.5 NL/min, without (a) and with a sweep flow of 0.5 NL/min (b), and comparison with equilibrium values.
further hydrogen production may only be obtained by increasing the operating temperature or sweep flow.
Influence of sweep gas to feed flow rates ratio The effect of improving separation, without acting on other phenomena, may be better understood by studying the
influence of sweep gas flow rate. Increasing sweep flow rate (Fig. 7) leads to better selectivity towards CO2 and CO and decreases that of CH4, due to H2 separation, which favors WGS and methane reforming to carbon dioxide. In the presence of sweep gas, CO2 production exceeds CH4 production. The MSR reaction is more heavily influenced by hydrogen removal
Fig. 6 e Hydrogen recovery and yield as a function of pressure at Tw ¼ 753 K, F ¼ 0.5 NL/min, without (a) and with a sweep flow of 0.5 NL/min (b).
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Fig. 7 e Product selectivity as a function of sweep flow rates at P ¼ 8 bar, Tw ¼ 703 K feed flow rate 0.5 NL/min (a) and 1 NL/ min (b) and comparison with equilibrium values.
compared to the WGS reaction, as the stoichiometric ratio between hydrogen and methane in MSR is four, whereas only 1 mol of H2 is produced for every mole of CO in WGS. The direction of selectivity towards CO is better observed in Fig. 8, where the exit ratio of CO and CH4 mole fraction (which equals that of the respective selectivity ratio) is compared to the closed system equilibrium values. As sweep flow (and consequently the permeate flow) increases, the MSR is enhanced compared to WGS, which in turn increases CO content relative to CH4 by as much as an order of magnitude. The influence of sweep gas is more pronounced when a lower feed flow rate is used, because the higher residence time allows a better hydrogen separation.
Membrane reactor modeling The behavior of the membrane reactor has been described with a one-dimensional pseudo-homogeneous model, with axial dispersion of mass and heat, in which the following main assumptions have been made:
- Ideal gas behavior; - Negligible axial pressure drop; - Negligible radial gradients radial temperature gradients were measured to be small; effects of concentration polarization were estimated to be small, based on an analysis carried out for methane steam reforming in the same reactor [2], using a correction reported in Ref. [21]. The correction depends on the reactor's radius, hydrogen diffusivity and membrane permeance, the conclusions reached in Ref. [2] are therefore extendable to the present work; - Hydrogen permeation through Pd-based membranes involves the steps of dissociative adsorption, subsurface penetration, diffusion through the membrane and the inverse steps on the other side. When the rate-limiting step is hydrogen atom diffusion through the membrane, as in the case of Pd-based membranes down to 1 mm [16], permeation in pure hydrogen follows Sievert's law (eq. (20)). When considering a gaseous mixture, the same law applies with a membrane permeance that depends on coadsorption of reactants and products:
J H 2 ¼ Qm
pffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi PyH2 PSW ySW H2
(20)
where membrane permeance is evaluated as: Em Qm ¼ QAm exp RT
(21)
and Q is a parameter accounting for permeance inhibition due to competitive adsorption of reactants and products.
Model equations The material balance equations are adapted from Ref. [2], applying them to the components and reactions of interest: Fig. 8 e Selectivity ratio between CO and CH4 directed by H2 removal. Experimental measurements of retentate composition compared to the closed system equilibrium values. P ¼ 8 bar, S/E ¼ 3.
εrg
3 vyi 1 vF vyi v2 yi X yi þ ¼ F þ Deff rg 2 þ ai;j rj ð1 εÞrs ; AO vz vt vz vz j¼1
isH2 (22)
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εrg
1 0:5 1:5 pCO2 PCH4 rED ¼ kED pEtOH kED
vyH2 1 vF vyH2 v2 yH2 yH2 þ ¼ F þ Deff rg AO vz vt vz vz2 3 X pffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi pdm þ aH2 ;j rj ð1 εÞrs Qm PyH2 PSW ySW H2 AO j¼1
1 KadsCO KadsH2 O pCO pH2 O KWGS pCO2 pH2
(23)
rg
vySW 1 vFSW SW vySW H2 H2 SW ¼ y þ F AI vt vz H vz
! þ
pffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi pdm Qm PyH2 PSW ySW H2 AI (24)
Effective dispersivity has been considered to be the same for all species, and equal to 1.7 104 m2/s. The following boundary conditions apply: Fjz¼0 ¼ F0 ; FSW z¼L ¼ FSW 0 ; ¼ yCO2 ¼ yCO z¼0 ¼ yH2 ¼ 0; yCH4 z¼0 z¼0 z¼0 1 S=E ; yH2 O ; ySW yEtOH z¼0 ¼ ¼ ¼0 H2 z¼0 z¼L 1 þ S=E 1 þ S=E
vT FH2;perm vT v2 T F0 0 ¼ kax:eff 2 cpg cp;H2 eff vt vz vz A F0 þ rs ð1 εÞ
3 X
DHj rj hw
j¼1
pD ðT Twall Þ A
(25)
vT vz
accounts for heat loss from the gas flow associated with hydrogen permeation. The boundary conditions are: vT Ng Cpg T0 Tjz¼0 ¼ kax;eff vz z¼0
1 KCO2 M
pCH4 p2H2 O
(28)
The adsorption constants KCO, KCO2 , and KH2 O have been considered constant and equal to their average value in the range 300e500 C, as suggested by Palma et al. [5] in their kinetic studies. The pre-exponential factors and activation energies of the kinetic constants are reported in Table 6 [5]. The effective axial conductivity has been evaluated through the following correlation:
jH ¼
where hw is wall to catalyst heat transfer coefficient, Tw is the controlled wall temperature, and FH2 ;perm is the hydrogen permeation rate. The term:
FH2 ;perm cp;H2
rCO2 M ¼ kCO2 M pCO2 p4H2
(27)
(29)
where ks is the intrinsic conductivity of the solid, equal to 50 W/mK and ε is the bed porosity, equal to 0.85. The wall heat transfer coefficient has been evaluated through a ChiltoneColburn type correlation:
The heat balance is: rcp
rWGS ¼ kWGS 2 1 þ KadsCO pCO þ KadsH2 O pH2 O þ KadsCO2 pCO2
1 kax;eff ¼ ks ð1 εÞ 3
vyi ¼0 vz z¼L
(26)
vT ¼0 vz z¼L
Although, axial mass dispersion is usually ignored in membrane reactor studies, the incorporation of axial conductivity is essential for this non-isothermal system. A non-stationary model has been developed to simplify the numerical solution procedure of this boundary-value problem. The system of partial differential equations has been solved using the PDE Toolbox of Matlab®.
Model parameters The kinetic expressions used in the model for the ethanol decomposition (Eqn. (10)), wateregas-shift (Eqn. (2)) and methanation (Eqn. (13)) reactions are those reported by Palma et al. [5], obtained by studying the kinetic activity of the catalyst used in the present work in an ethanol steam reforming reactor operating at different contact times between temperatures of 523 and 873 K. The rate equations have been re-written here for clarity:
Nuw Pr1=2 Re
¼ aReb1
(30)
A thermal characterization of the solid foam has been previously carried out, and the value of the parameter b has been determined to be 0.29 [17]. The value of a has been obtained in the course of the present work by fitting experimental data and is equal to 24.1. A complete list of parameters used in the model, along with their meaning and value, is reported in the List of Symbols section at the end of the paper.
Model results Fitting the experimental results presented here (conditions summarized in Table 2) with the model described above, using thermodynamic and kinetic parameters from the literature, and only two adjustable parameters, has shown good agreement (Fig. 9). However, the estimated permeance at 723 K is about 12 Nm3h1m2bar0.5. This value should be compared with the permeance of the membrane used in the present work that was measured by the producer (ECN) to be about 75 Nm3h1 m2 bar0.5 at 723 K in pure hydrogen (Em ¼ 11 kJ/ mol and Am ¼ 5.6 mol/(m2sbar0.5)). Hydrogen flow in pure hydrogen and mixtures was found to follow Sievert's law, indicating that the limiting step in its permeation is the diffusion of atoms through the bulk of the membrane. This is a
Table 6 e Pre-exponential factors and activation energies of the kinetic constants. k0
Reaction Modified ethanol decomposition Water gas shift CO2 methanation
4
7.94 10 mol/(gcat s bar) 2.77 108 mol/(gcat s) 7.65 101 mol/(gcat s bar4)
Ea [kJ/mol] 104 74 8
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while the concentration of other components is higher. This phenomenon is known as concentrationpolarization. In narrow reactors, such as the one used in the present study, the effect of concentrationpolarization is not negligible, but cannot account for the large decline (see below); (ii) Co-adsorption of reactants and products on the Pdsurface may lead to inhibition: While there are many studies of the inhibiting effects of CO [18] the effects of methane adsorption is contradictory, while computational (DFT) studies show negligible effects due to CO2, CH4 or H2O adsorption [19]. The relatively small concentrations of CO make this explanation questionable as well; (iii) Surface reactions (such as reverse-WGS) may cause adsorption of CO on the Pd surface.
Fig. 9 e Comparison between experimental (points) and calculated (lines) molar fractions on dry basis as a function of temperature at P ¼ 9 bar, 0.5 NL/min feed flow rate, sweep gas flow rate 0 NL/min (a), 0.5 NL/min (b), and 1 NL/ min (c). significant inhibition (Q ¼ 0.15.) and we try to understand its sources. No attempt was made to fit the data using an inhibition Q function that depends on composition, as that would involve adjusting too many parameters. Furthermore, gas composition in proximity of the entrance section is not known precisely. The decrease in apparent permeance may be due to several reasons: (i) In wide reactors the incorporation of high-permeance membranes leads to the hydrogen concentration at the membrane wall to be much lower than in the bulk,
An inhibition effect similar to the one noticed here was previously observed in a study of methane steam reforming in the same membrane reactor [2] and an attempt was made to discriminate between various sources of inhibition. It was argued that the strong inhibition cannot be explained only by concentration polarization or by competitive adsorption by gas-phase species. The former effect was ruled out by approximate criteria and corrections to the 1-D model accounting for transversal effects, which were developed for hydrogen separation without [20] or with [21] reaction. The latter effect was ruled out using partial information on adsorption properties from the literature: CO inhibition is strong but its concentration is small. On the other hand, the inhibition can be accounted for by competitive adsorption by species produced on the surface. This issue is hardly addressed in the literature, but some indications do exist: Li et al. [22] found that both CO and H2O reduce hydrogen permeation through a Pd/stainless steel membrane at 653 K. Permeation tests with binary H2/CO and H2/H2O mixtures of the same concentrations of CO and H2O showed that steam had a stronger inhibition effect than CO on hydrogen permeation. Hulme et al. [23] found significant decline of hydrogen permeability through a PdeAg/VeNi alloy membrane in presence of CO2 at 473e673 K (60% decline with 30% CO2 and stronger at lower temperatures). These results can only be explained by higher CO surface concentrations than those expected from gas phase composition. Fig. 10 shows the calculated temperature profile along the reactor axis when the wall temperature has been set at 753 K and the reactor operates at 8 bar with a 1 NL/min feed flow rate and 1 NL/min sweep gas flow rate. The inlet temperature was equal to 573 K: the calculated temperature increases rapidly in proximity of the inlet, due to the combined effect of the exothermicity of the ethanol decomposition reaction and the heat transfer from the reactor wall. Temperature tends to decrease towards the final part of the reactor. This behavior may be attributed to the fact that when the sweep gas is sent counter-currently with respect to the reactor flow, the hydrogen partial pressure in the permeate side is close to zero in correspondence of the exit section of the reactor. Hydrogen permeance through the membrane shifts the carbon dioxide methanation reaction backwards, absorbing heat and causing a reduction in temperature.
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providing the structured foam, and the Energy Research Centre of the Netherlands (ECN), for providing the membrane. The experimental equipment was partially supported by Technion Grand Energy Program (TGEP).
List of Symbols a A0 AI Am
Fig. 10 e Calculated (continuous) and measured (points) temperature along reactor axis for Tw ¼ 753 K, P ¼ 8 bar, 1 NL/min feed flow rate, sweep gas flow rate 1 NL/min.
Conclusions The present work showed that in a system of competing reversible reactions, the desired one can be accelerated by separating one of its products, hydrogen in this case. This novel idea has the potential of suggesting new designs for improving selectivity. It was shown experimentally that pure hydrogen can be produced by ethanol steam reforming in an integrated catalytic membrane reactor, with complete conversion, good selectivity, and with no appreciable catalyst deactivation. As explained in the Introduction a membrane reformer heated by solar energy through molten salt represents an interesting process for renewable hydrogen production, which can be converted to electricity via FC. The possibility of working with low amounts of water and relatively low pressures reduces the energetic cost of conducting ethanol steam reforming. Using steam instead of nitrogen as sweep gas will also make the overall process easier, by allowing separation of the permeate side current by simple condensation of water. Process optimization can proceed with the mathematical model that was developed; the model uses only two adjustable parameters: a permeance inhibition factor, which was found to be significant, and the ChiltoneColburn coefficient of the wall heat transfer coefficient. Understanding the drop in permeance will allow to reduce the membrane (and reactor) size and cost.
Acknowledgments This work has been conducted within the framework of the CoMETHy (Compact Multifuel-Energy to Hydrogen converter) project, funded by the European Commission under the Fuel Cells and Hydrogen Joint Undertaking (GA No. 279075). The authors wish to acknowledge the Fraunhofer Institute for Ceramic Technologies and Systems, IKTS (Germany), for
b dm dw Deff Ea Em F FH2 ;perm jH kax,eff kg ks k0 KadsCO KadsCO2 KadsH2 O KCO2 M KED KWGS Ng Nuw P Pr Qm rj Re S/E T u y
coefficient of ChiltoneColburn correlation (24.1, from data fitting), annular section (1.1 103), m2 membrane cross section (6.24 105), m2 membrane permeance pre-exponential factor (6.6), mol/m2/s/bar0.5 exponent of ChiltoneColburn correlation (0.29), outer membrane diameter (0.014), m outer reactor diameter (0.04), m effective dispersivity (1.7 104), m2/s rate constant activation energy, kJ/mol membrane permeance activation energy (11), kJ/mol gas flow rate (evaluated from model), mol/s hydrogen permeation rate (evaluated from model), mol/s ChiltoneColburn coefficient (evaluated from model), effective axial conductivity (2.5), W/m/K gas conductivity (6.9 102), W/m/K intrinsic solid conductivity (50), W/m/K pre-exponential factor of the rate constants, mol/ gcat s bar CO constant (0.37), atm1 CO2 adsorption constant (1.701), atm1 H2O adsorption constant (80073.2), atm1 CO2 methanation reaction equilibrium constant ethanol decomposition reaction equilibrium constant water gas shift reaction equilibrium constant specific gas flow rate (evaluated from model), mol/m2/s wall Nusselt number (evaluated from model), pressure, bar Prandtl number (Pr ¼ cp,gmg/kg), membrane permeance (evaluated from model), mol/ m2/s/bar0.5 j-th reaction rate (evaluated from model), mol/gcat/s Reynolds number for a packed bed (Re ¼ rgudw/mg), steam:ethanol molar ratio at inlet, temperature, K gas superficial velocity (evaluated from model), m/s molar fraction (evaluated from model),
Greek letters ε bed porosity (0.85), gas phase molar density (evaluated from model), mol/ rg m3 rs catalyst density (1.9), g/mL a stoichiometric coefficient, Q inhibition parameter (0.15, from data fitting), Subscripts i i-th component j j-th component w wall m membrane
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i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 4 0 ( 2 0 1 5 ) 5 8 3 7 e5 8 4 8
solid gas
Superscripts SW sweep side
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