Kinetic Study of Hydrogen Production by the High Temperature Water

Article pubs.acs.org/IECR

Kinetic Study of Hydrogen Production by the High Temperature Water Gas Shift Reaction of Reformate Gas in Conventional and Membrane Packed Bed Reactors over Ca-Promoted CeO2−ZrO2 Supported Ni−Cu Catalyst Ishioma Judith Oluku, Hussameldin Ibrahim,* and Raphael Idem* Clean Energy Technologies Institute (CETI), Process Systems Engineering, Faculty of Engineering and Applied Science, University of Regina, Regina, Saskatchewan S4S 0A2, Canada S Supporting Information *

ABSTRACT: The kinetics of the high temperature water gas shift (HTWGS) reaction was performed in a hydrogen selective membrane reactor (MR) and compared with that in a packed bed tubular reactor (PBTR) at a pressure range of 150−250 psi, temperature range of 400−500 °C and weight of catalyst/CO molar flow rate ratio of 1.1−2.7 (g-cat·hr)/mol using reformate gas over a Ni−Cu catalyst on Ca-promoted CeO2−ZrO2 support. The purpose was to estimate the effect of hydrogen removal on conversion and reaction products yield. The characteristics and hydrogen permeation properties of the MR were also studied. The activation energy obtained for the reaction in the PBTR was 190 kJ/mol, whereas the MR exhibited activation energy of ca. 50% lower than the PBTR. MR characterization results showed that hydrogen permeation and recovery increased with temperature and trans-membrane pressure, which could be correlated using Sievert’s law to yield an activation energy of 22.7 kJ/ mol. forces.10,11 The trans-membrane pressure is achieved by operating at a high pressure on the retentate side while the permeate side is kept at atmospheric pressure in the presence or absence of a sweeping gas. Several materials have been used for manufacturing the membrane reactor for WGSR including silica, zeolites, palladium and its alloys of Cu and Ag, but Pd based membranes have been proven to be better because of their high selectivity to hydrogen.13 The use of a membrane reactor in the WGSR thus offers the advantage of allowing the use of higher temperatures while maintaining a forward reaction path, lower steam/CO ratio, integration of reaction and separation steps, utilization of higher pressure, higher hydrogen recovery and purity.2 Although the membrane reactor reduces the amount of steam required for the WGSR, a low steam/CO ratio results in certain side reactions such as methanation and Boudouard reaction, producing unwanted products like carbon and methane, as shown in eqs 2−4.

1.0. INTRODUCTION It is well-known that hydrogen is one of the key sources of clean and/or renewable energy, which can be used as a feed to a fuel cell1 because it burns clean and produces water without any carbon emissions. The reforming reaction is typically a gas phase process that can produce hydrogen cost-effectively and efficiently lending itself toward use at the point of application, thereby meeting the requirements of a hydrogen fuel cell.2,3 The water gas shift reaction (eq 1) occurs downstream of the reformer and involves the reaction between carbon monoxide (CO) and steam (H2O) to form hydrogen and carbon dioxide. CO + H 2O ⇆ H 2 + CO2

(1)

The conventional process for the water gas shift reaction involves two shift reactors and several purification units downstream resulting in significant energy and infrastructure requirements in order to meet the desired purity of hydrogen. This leads to reduced plant efficiency and hydrogen recovery.4−6 Several efforts have been made to simplify the conventional water gas shift process such as the cyclic water gas shift reaction7 and sorption enhanced water gas shift reaction.8 However, with the recent development of membrane reactors that combine the hydrogen production and separation in one step, the traditional process can be redesigned in order to achieve more compact plant size and improved hydrogen yield and CO conversion.2 Typically, the reaction takes place in the presence of a catalyst in the retentate side and the hydrogen produced is transferred to the permeate side; thus, the continual removal of a product promotes the forward reaction, according to Le Chatelier’s principle.9 The rate of permeation of the species across the membrane is dependent on temperature and trans-membrane pressure acting as the driving © XXXX American Chemical Society

CO + 3H 2 ⇔ CH4 + H 2O (methanation)

(2)

CO2 + 4H 2 ⇔ CH4 + H 2O (methanation)

(3)

2CO ⇔ C + CO2 (Boudouard reaction)

(4)

Typically, the catalyst should be able to operate above 680 K (407 °C) to limit the poisoning effect of CO on the palladium (Pd) alloy membrane,12 be considerably more active than the Received: May 21, 2014 Revised: November 28, 2014 Accepted: December 26, 2014

A

DOI: 10.1021/ie502082q Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research

Figure 1. Schematic diagram of membrane reactor experimental setup.

The high temperature operation also obviates the need to cool the product stream of dry reforming or gasification to below 500 °C. It can be seen that there is the need and benefit to develop a highly stable and active catalyst for the water gas shift reaction (WGSR) that can withstand high temperatures and pressures, and can be used in the presence of other components which typically make up reformate gas (from dry reforming of natural gas, biogas methane or from gasification) such as CO2, CH4 and H2. Moreover, in order to design an appropriate reactor system around the developed catalyst, information on the reaction kinetics is essential. Also, since the water gas shift reaction (WGSR) is thermodynamically limited, the incorporation of a selective membrane reactor that continuously removes hydrogen from the reaction system should enhance the forward driving force of the WGSR. Several kinetic studies have been performed on the water gas shift reaction in a membrane reactor but were limited to low temperatures and/or low pressures or performed via simulation.18,20 In this study, a newly developed HTWGSR catalyst was tested in a membrane reactor and its reaction kinetics studied at high temperatures and pressures in comparison with corresponding conditions and at atmospheric pressure in a plug flow tubular reactor. These results are presented and discussed in this paper.

current commercial high temperature formulations and be very stable, as well as not form carbonaceous residues even with low H2O/CO ratios because carbon has a great affinity for Pd, which could rapidly destroy the membrane.12 The majority of the studies on membrane reactors have focused on improvement of the membrane material, manufacturing methods and modeling studies. The majority of these studies employed commercially available catalysts (Fe/Cr and Cu/Zn/Al2O3) and artificial feeds that do not represent an industrial scenario,13−15 which are the product streams of reformate gas from gasification processes or dry reforming of biogas or natural gas. There are only a few literature reports on catalyst development for WGS in membrane reactors. In some cases, the catalyst developed is only tested in a traditional fixed bed reactor under potential membrane conditions and, in other cases, mild temperature and pressure conditions are employed in the tests to accommodate limitations in membrane performance that could occur at elevated reaction conditions.12,13,16−18 In the present study, the experimental conditions followed represent industrial operations as well as membrane reactor testing at conditions of interest to industry. Accordingly, we have developed a highly stable and active catalyst19 for the water gas shift reaction (WGSR) that can withstand high temperatures and pressures, and can be used in the presence of other components that typically make up reformate gas (from dry reforming of natural gas, biogas methane or from gasification) such as CO, CO2, CH4 and H2. B


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2.0. EXPERIMENTAL SECTION 2.1. Catalyst. The catalyst used in this study is a low cost, active and stable formulation developed in-house for the WGSR of reformate gas at high temperatures. The catalyst was prepared using surfactant assisted route for its ternary oxide support, Ca-promoted CeO2−ZrO2 (CeZrCaOx) followed by wetness impregnation of the active metals Ni and Cu. Details of the preparation of the catalyst and the catalyst characteristics are reported elsewhere.19 2.2. Experimental Apparatus. The goal of these experiments was to show the effect of using a membrane reactor on the water gas shift reaction. The membrane reactor was first characterized to determine the rate of permeation of hydrogen and its dependence on temperature and pressure. Thereafter, the catalyst performance was evaluated in the membrane reactor. The membrane reactor (REB Research & Consulting, USA) consists of an Inconel reactor shell 0.5 in. I.D. enclosing a single tubular Pd−Ag membrane 0.125 in. O.D. and 6.5 in. effective length. There are two outlet tubes from the reactor; the raffinate/retentate stream and the pure H2 stream (permeate). Details such as ratio of Pd to Ag, presence of other metals as well as thickness and support mechanism were not provided. The catalyst is placed into the annular space between the reactor shell and membrane tube. A sliding three-point thermocouple is also inserted into the reactor for temperature measurement. The reactor is placed vertically in an electrical furnace (REB Research & Consulting, USA). The experimental setup used is shown in Figure 1. 2.3. Membrane Characterization Studies. Membrane characterization studies were conducted to show the relationship between hydrogen flux and pressure as well as temperature and to determine the permeance of hydrogen through the membrane tube. The experiments were performed using 30% H2/70%N2, 40%H2/60%N2 and 50%H2/50%N2 gas mixtures within a temperature range of 300−500 °C and pressure range of 150−250 psi (10.2−17 atm). Prior to the test, the reactor was filled with α-Al2O3 particles of 0.8 mm to ensure even heat and flow distribution and to mimic real-world catalytic runs such as high temperature and pressure, and the presence of other components which typically make up reformate gas (from dry reforming of natural gas, biogas methane or from gasification) such as CO, CO2, CH4 and H2. The temperature was raised to the desired value in the presence of ultrahigh pure (UHP) N2 and then the gas switched to the desired H2/N2 mix. The pressure on the retentate side was then increased gradually from atmospheric pressure to the desired pressure by means of the upstream gas regulator on the cylinder and maintained by a back pressure regulator (TESCOM Corporation) downstream of the reactor, keeping the inlet gas flow (275 mL/min) hence residence time constant for all pressures. The permeate side of the MR was operated at atmospheric pressure thereby creating a trans-membrane pressure differential acting as a driving force to allow for the permeation of hydrogen. The retentate and pure H2 stream were sent to online gas chromatography equipped with a thermal conductivity detector (GC-TCD) for analysis and the outlet flow rates of each stream was recorded so as to determine the hydrogen recovery (RH2) defined according to eq 5. R H2 =

(H2)perm (H2)perm + (H2)ret

where (H2)perm is the amount of hydrogen that permeated through the membrane and H2ret is the amount of hydrogen left in the retentate section. For membrane reactor characterization, hydrogen recovery (RH2) is the fraction of hydrogen that can be recovered on the permeate side as compared to the original amount fed to the reactor ((H2)perm + (H2)ret). In the case involving WGSR, hydrogen recovery is defined as the fraction of hydrogen that can be recovered on the permeate side of the membrane reactor as compared to the total amount of hydrogen available in the reactor (i.e., (H2)feed + (H2)WGSR), as shown in eq 6. R H2 =

(H 2)perm (H 2)feed + (H 2)WGSR

× 100%

(6)

where (H2)feed is the molar flow of hydrogen introduced from the feed, and (H2)WGSR is the net hydrogen produced from WGSR. 2.4. WGSR in Membrane Reactor. A new catalyst developed in an earlier study by our group with formulation of 3Ni-5Cu/CeZrCa was selected for the WGSR in the membrane reactor in order to determine its suitability for MR conditions. Reformate gas from biogas dry reforming composed of 46% CO, 38% H2, 9% CH4, 7% CO2 (Praxair) was used. The flow rates of the gases were measured using a mass flow controller (Aalborg Instruments) downstream of the setup. The pressure of the gases was controlled using the upstream and downstream pressure regulators. The downstream backpressure regulator was used to build the pressure as well as regulate and maintain flow to the desired rate. An HPLC (Alltech) pump supplied water at the desired flow rate, while a three-point thermocouple was used to measure the reaction temperature inside the catalyst bed. The thermocouple (REB Research & Consulting) was designed to measure temperatures at the bottom, middle and top of the reactor. This is important so as to place the catalyst such that the thermocouple is in the middle of the catalyst bed where the reactor temperature is controlled. In a typical run, about 15 g of α-Al2O3 is poured into the reactor to create an elevated bed for the actual catalyst mix to sit on, such that the catalyst bed is positioned at a point close to the middle of the reactor were optimum heat transfer occurs. 0.25 g of 0.8 mm sized (Dp) catalyst particles was then mixed with another 15 g of α-Al2O3 (Sasol GmbH, Germany) used as a diluent, thus forming a catalyst bed height (L) of about 9 cm in the reactor. The catalyst weight, size, bed height, temperature and steam/CO ratio were maintained at similar values to the experiments performed in the packed bed reactor for the purpose of comparison. Plug flow conditions (zero radial velocity and temperature gradient profiles) were met by ensuring the values of L/Dp and D/Dp were greater than the minimum requirement of 50 and 10, respectively, in the reactor.21 Prior to the actual reaction, the catalyst was activated at atmospheric pressure and 500 °C in situ by flowing 5%H2/ balance N2 through the catalyst bed for 2 h. Water at a flow rate of 0.26 mL/min and gas mixture flowing at 275 mL/min was fed into the reactor by means of a HPLC pump (Alltech) and maintained at the desired reaction temperature of 500 °C for the duration of the reaction, typically 6 h. The two product streams, raffinate and pure H2, exiting the reactor were passed through a condenser/gas−liquid separator and cooler, respectively, to condense the water in the raffinate stream, thus leaving the gaseous phase and reduce the temperature of the pure H2 stream, which was sent to the GC. The criteria

× 100% (5) C


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Industrial & Engineering Chemistry Research used to evaluate the performance of the reaction are shown below. CO conversion (XCO) =

(moles of COin − moles of COout ) × 100 moles of COin

H 2 yield (YH2) =

moles of H 2produced theoretically expected H2

H 2 selectivity (SH2) = TOF =

(7)

× 100

H 2 yield × 100 conversion (Xco)

(8)

(9)

number of molecules of reactant converted per second −1 [s ] number of atoms of active site

(10)

Theoretically expected H2 was calculated to be equal to the moles of CO converted in the WGSR. 2.5. Kinetic Experiments. The kinetic studies were carried out using reformate gas at reaction temperatures of 400, 450 and 500 °C. The catalyst weight time (W/FA0) was varied between 1.1, 1.6 and 2.7 (g-cat·hr)/mol by keeping the weight of the catalyst constant at 0.5 g while changing the feed flow rate between 161, 275 and 389 mL/min. These conditions were selected such that equilibrium conversion was avoided22 while keeping the catalyst bed height and feed composition constant. The kinetic reactions were carried out in the PBTR at atmospheric pressure and elevated pressure of 200 psi (13.6 atm) as well as in the membrane reactor.

Figure 2. Membrane characterization results in terms of H2 recovery at (a) 30%H2/bal.N2 for temperature range 300−500 °C and (b) 300 °C for H2/bal.N2 range of 30%−50%.

3.0. RESULTS AND DISCUSSION 3.1. Membrane Characterization Results. The membrane reactor used in this study was characterized using different H2 mixtures first to verify the perm-selectivity of the membrane to pure H2 under varying operating conditions of temperature and pressure and then to determine the relationship between the rate of permeation of H2 and the difference in partial pressure and temperature. The results of the characterization tests are presented in Figures 2 and 3 in terms of H2 recovery and flux. In all cases, pure H2 (from GC analysis) was obtained on the permeate side, indicating the 100% H2 selectivity of the membrane. In Figure 2a, the variation of the amount of H2 recovered is correlated with the temperature and total pressure employed using 30% H2/balance N2 whereas Figure 2b relates the influence of gas composition and total pressure to the H2 recovery. It is important to note that the pressure on the permeate side is at atmospheric conditions in all runs with no sweep gas employed, thus creating a transmembrane pressure differential allowing the permeation of hydrogen. The results show a linear increase in the H2 recovery with increasing temperature and pressure. In Figure 2a, the increasing gas composition from 30% to 50% H2 translates to an increase in partial pressure thereby increasing the recovery even at a constant temperature of 300 °C. This observation is in consonance with Sievert’s law in eq 11, which relates the flux of H2 through a Pd membrane (NH2) to the difference in partial pressure with an exponent of 0.5, while the temperature dependence of the H2 recovery is described by Arrhenius law relating the H2 permeability to the temperature as shown in eq 12. Combining eqs 11 and 12 yields eq 13. In these equations, Pe is the permeance.

Figure 3. Membrane characterization results showing Sievert’s law for (a) hydrogen flux at 300 °C for 30%H2/bal.N2 range of 30%−50% and (b) hydrogen flux at 30% H2 mix for temperature range of 300−500 °C.

NH2 = Pe(PH2ret 0.5 − PH2perm 0.5) D

(11) DOI: 10.1021/ie502082q Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Figure 4. WGSR performance in membrane reactor using reformate gas at 275 mL/min gas with steam/CO molar ratio of 2.7 and 3Ni5Cu/CeZrCa catalyst at 500 °C in PBTR (a) CO conversion, (b) H2 yield, (c) H2 selectivity and (d) H2 recovery.

Pe = P′exp−E / RT

(12)

NH2 = P′exp−E / RT (PH2ret 0.5 − PH2perm 0.5)

(13)

pressure further to 250 psi (17 atm) yields no significant effect on the CO conversion. This is because of the available membrane surface area in the reactor thus limiting the rate of hydrogen permeation, hence the forward reaction. A higher CO conversion may be obtained with more available membrane surface area. It is also possible that the occurrence of other processes such as mass transfer limitations, competitive adsorption of active sites on the membrane surface due to CO, CO2 and CH4, and CO poisoning may also limit the rate of hydrogen permeation, in addition to the existence in the membrane of a finite surface area. At atmospheric pressure, the membrane reactor operates as a conventional packed bed tubular reactor due to the absence of a pressure differential driving force. On the other hand, the H2 yield is independent of pressure increase while H2 selectivity tends to decrease with increasing pressure from 150 to 200 psi but stabilizes between 200 and 250 psi. This implies that although equal amounts of CO2 and H2 are produced based on the WGSR, side reactions, especially methanation reactions, do occur whereby about 4 mol of H2 requires 1 mol of CO2 for conversion to CH4. The net result is that there is more CO2 produced than H2. In all cases, there is negligible conversion of CH4 present in the reformate gas. Although, the composition of CH4 shows 15%, this does not imply formation of CH4 but results from the H2 permeating through the membrane thus increasing the fraction of the CH4 in the raffinate stream per unit volume. It is also clear from Figure 4d that the H2 recovery increases linearly from 6% to 16% with increasing pressure from 150 to 250 psi (10.2−17 atm) despite the independence of CO conversion and H2 yield on pressure above 200 psi (13.6 atm). This is in accordance with the Sievert’s law such that the increasing partial pressure improves the driving force across the membrane; hence the continuous increase in H2 recovery. However, it is pertinent to note that the H2 recovery obtained during the WGSR is lower than that obtained from the permeation tests where H2/N2 mixtures were used. This is because the rate of H2 permeation is influenced by external gas

Based on Sievert’s law, a plot of hydrogen flux against partial pressure difference was obtained as illustrated in Figure 3a,b using a surface area of 1.65 × 10−3 m2 whereby the slope represents the permeance of hydrogen. From Figure 3a,b, it is evident that the slope (permeance) increases with H2 composition, temperature and pressure. This is in agreement with the dissociative adsorption, diffusion, recombination and desorption mechanism of hydrogen on the surface of a Pd membrane.6 A nonlinear regression was performed on the data yielding an activation energy of 22.7 kJ/ mol, and a pre-exponential factor of 4.56 × 10−6 (mol·m)/m2·s· kPa0.5. The activation energy obtained in this work is slightly higher than those reported in the literature23,24 and can be attributed to negligible mass transfer resistance across the membrane, which results in higher flux, unlike the lower activation energy reported in the literature, which indicates that the bulk diffusion and/or surface process is the rate limiting step.25 There is also the possibility that other processes such as those involving the metal alloy used in the membrane, can contribute toward the slightly higher activation energy. 3.2. WGSR in Membrane Reactor. The water gas shift reaction was performed in the membrane reactor varying the pressure between 150 and 250 psi (10.2−17 atm) at 500 °C using reformate gas with a flow rate of 275 mL/min and steam/ CO molar ratio of 2.7. The results are shown in Figure 4 in terms of CO conversion, H2 yield, H2 selectivity and H2 recovery. The result indicates an improvement in CO conversion from 70% to 80% when the pressure is increased from 150 to 200 psi (10.2−13.6 atm). This is because increasing the gas pressure on the retentate side results in an increase in the driving force for hydrogen permeation across the membrane, thereby allowing more CO to be converted to product according to Le Chatelier’s principle. Increasing the E


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Figure 5. Effect of catalyst weight and flow rate on WGSR performance in membrane reactor using reformate gas steam/CO ratio of 2.7 and 3Ni5Cu/CeZrCa catalyst at 500 °C at 200 psi. KP = Xeq2/(1 − Xeq)(2.7 − Xeq).

parameters varied include (i) catalyst weight, (ii) feed flow rate and (iii) steam/CO ratio. The results obtained thereof are presented in Figures 5 and 6. 3.3.1. Effect of Catalyst Weight and Feed Flow Rate in Membrane Reactor. The effect of varying the weight of catalyst was investigated using 0.25, 0.5 and 1.0 g of 3Ni5Cu/CeZrCa catalyst under similar conditions of 500 °C, 200 psi and steam/ CO ratio of 2.7. The flow rate of reformate gas feed was also varied between 161 and 275 mL/min keeping the steam/CO ratio constant to evaluate the influence of feed flow rate on the WGSR. Figure 5 illustrates the results in terms of CO conversion, H2 yield, selectivity and recovery. It is observed

phase mass transfer limitations such as the presence of other species, in this case, CO, CO2 and CH4, which tend to compete with hydrogen molecules for adsorption on the membrane surface, thereby reducing the amount of hydrogen permeating through the membrane. This is similar to results reported by Liguori et al.26 An optimum operating pressure is therefore chosen based on a balance of CO conversion, H2 yield, selectivity and recovery; in this case, 200 psi (13.6 atm) was selected for subsequent parametric and kinetic studies. 3.3. Parametric Study in Membrane Reactor. In this section, several parameters were varied to determine the effect on the performance of the WGSR in a membrane reactor. The F


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3.3.2. Effect of Steam/CO Ratio in Membrane Reactor. The effect of steam/CO ratio in the range of 1.7, 2.6 and 3.7 was studied for the WGSR in the membrane reactor using reformate gas at a high pressure of 200 psi (13.6 atm) and 500 °C, and the results are presented in Figure 6. The plots reveal a drastic increase in CO conversion from 34% to 90% and H2 yield from 22% to 52% upon increasing the steam/CO ratio from 1.7 to 2.7. This is due to the stoichiometry of the WGSR, such that increasing the amount of excess reactant (steam) leads to a higher CO conversion to products. A further increase in the steam/CO ratio to 3.7 yielded no additional improvement in conversion but a gain in H2 yield from 52% to 100%. The gain in hydrogen yield can be attributed to the steam reforming of methane present in the reformate feed according to eq 14. CH4 + H 2O ⇆ CO + 3H 2

(14)

This is apparent in the product distribution of the product stream in which the amount of methane in the product stream is significantly lower than that in the feed composition. It is evident that a higher amount of steam would be required in the PBTR to attain a similar performance to that achieved in the membrane reactor. The overall performance of the membrane reactor with various steam/CO ratios is therefore better than that of the PBTR using the 3Ni5Cu/CeZrCa catalyst. This indicates that the efficiency of the WGSR can be improved by the incorporation of a membrane reactor along with a highly active and stable catalyst, such as the one developed in this study while utilizing a low steam/CO ratio of 2.6, thereby saving the amount of steam which would otherwise be required in a conventional reactor to attain the same performance. It is observed in Figures 5 and 6 that there is an induction period before the catalyst reaches peak performance in the membrane reactor runs. This is likely to represent the time required for hydrogen transport through the membrane to reach steady state conditions. 3.4. Results of Kinetic Experiments. The reaction kinetics for the WGSR using reformate gas over 3Ni5Cu/ CeZrCa was studied in three different systems; (i) membrane reactor at 200 psi (13.6 atm), (ii) high pressure packed bed reactor at 200 psi (13.6 atm) and (iii) packed bed reactor at atmospheric pressure. The catalyst particle size used for all reactions was 0.8 mm. Two parameters were varied, namely, (i) feed flow rate and (ii) reaction temperature. This was achieved by changing the flow rate between 161 and 389 mL/min using a constant catalyst weight of 0.5 g and water flow rate between 0.16 and 0.37 mL/min to maintain a H2O/CO ratio of 2.6 at reaction temperatures of 400, 450 and 500 °C (673, 723 773 K). This resulted in a variation of W/FA0 (catalyst weight time) between 1.1 and 2.67 (g-cat·hr)/mol representing the residence time for the reaction. These are shown in Table S1 (Supporting Information). The actual flow rates of reformate gas used in the experiments are given in Table S2 (Supporting Information). All the operating conditions used were to ensure that mass transfer and heat transfer limitations that could affect the kinetic data were eliminated. These are described below. 3.4.1. Investigation of Heat and Mass Transfer Limitations. In the kinetic study of any reaction, data collection can only be considered intrinsic in the absence of heat and mass transfer resistance. In this section, the interparticle and intraparticle mass and heat transfer resistances and the effects

Figure 6. Effect of steam/CO on product distribution in membrane reactor using reformate gas and 3Ni5Cu/CeZrCa catalyst at 500 °C at 200 psi.

that increasing the weight of catalyst from 0.25 to 0.5 g causes an increase in CO conversion from 79% to 90% when the reformate flow rate is 275 mL/min and only a slight increase from 90% to 93% for reformate flow rate of 161 mL/min. This indicates that at a lower residence time (higher flow rate), the significance of the amount of catalyst used is more evident. Increasing the catalyst weight beyond 1.0 g has a negligible effect on the resulting CO conversion for both flow rates. At both flow rates, the equilibrium conversion of 90.7% (indicated by the dashed line) under the prevailing conditions is surpassed with 0.5−1 g of catalyst, which is one of the advantages of using a membrane reactor. CO conversion of 84% was obtained at 161 mL/min, which is slightly higher than the 81% conversion obtained at 275 mL/min using a catalyst weight of 0.25 g. This is as a result of a higher residence time attained at lower flow rate allowing more CO molecules to be converted to products. A similar trend is also observed for the H2 yield, selectivity and recovery with a more significant improvement in performance at a lower flow rate. On the other hand, it is observed that H2 yield and selectivity reduced slightly when the catalyst weight was increased from 0.25 to 0.5 g and no significant change was observed between 0.5 and 1.0 g while the hydrogen recovery is found to be independent of the catalyst weight change. This implies that, although increasing the amount of catalyst may be beneficial for improving the CO conversion, this variation does not necessarily improve the amount of hydrogen obtained. Using 1.0 g of catalyst, however, tends to quickly stabilize the H2 yield and selectivity for the duration of the run. Again, in this case negligible CH4 conversion was observed. G


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Mass Transport Effects. The Weisz Prater criterion was used to determine the internal pore mass transfer resistance as follows

on the rate of the reaction will be investigated at 773 K, the highest temperature used for the kinetic studies. Heat Transport Effects. According to the Prater analysis, the internal pore heat transfer resistance is given by ΔTmax =

Deff (CCOs − CCOc)ΔHrxn λeff

Cwp,ipd = (15)

hT 2R

< 0.15

(19)

−2

This ratio obtained was 4.7 × 10 indicating that the observed rate is far less than the rate if external resistance was controlling. Thus, the resistance to external mass transfer does not affect the reaction rate. A more rigorous criterion developed by Mears32 was also applied to determine the onset of film mass transfer limitation. r′CO,obs ρb R Cn k CCCO

< 0.15

(20)

The value obtained from the LHS was 0.08, which is far less than 0.15, implying that there was no mass limitation in the film. 3.4.2. Kinetic Model. 3.4.2.1. Power Law Model. The power law model is a simple empirical rate expression that has been widely applied in kinetic studies. The overall reaction used for the development of the kinetic model for water gas shift reaction is given as CO + H 2O ⇆ CO2 + H 2 ΔH 0rxn = −41

(16)

(1A)

kJ mol

An empirical, reversible power law rate model can be written

where ΔTfilm,max is the upper limit of temperature difference between the gas bulk and the pellet surface, Lc is the characteristic length, rCO,obs is the observed rate of reaction, h is the heat transfer coefficient (estimated from the correlation JH = JD = (h/Cpμρ)NPr2/3, where JH is the heat transfer J factor, NPr = Cpμ/λ, λ is the molecular thermal conductivity, Cp is the heat capacity). JD factors are given by the following correlations: JD = (0.4548/εp)NRe−0.4069;29 NRe = dpμρ/μ(1 − εp). kc is the mass transfer coefficient obtained to be 0.183m/s. The heat transfer coefficient, h, obtained was 0.106 kJ/(m2·s·K) and a value of 5.3 K resulted for ΔTfilm,max. Furthermore, a more rigorous criterion developed by Mears32 for determining heat transport limitation during reaction was employed to further confirm the negligible resistance offered by heat transfer on the rate of the reaction. r′CO,obs ρb R CEΔHrxn

(18)

−rCO,obs d p observed rate = rate if film resistance controls CCObkc 6

Lc( −rCO,obs)ΔHrxn h

Deff CCOs

where Cwp, ipd is the Weisz−Prater criterion for internal pore diffusion, ρC the pellet density and RC catalyst radius. The Cwp, ipd value estimated was 0.982, which is slightly less than 1, indicating the potential for mass transfer limitation within the catalyst pores. However, to further prove that the mass transfer limitation has no significant effect on the rate of reaction, the following ratio was employed:

where ΔTmax is upper limit to the temperature difference between the pellet center and surface (K), Deff is DABεp/τ effective diffusivity, m2/s, DAB is bulk diffusivity of component A (CO) in B (oxygen), m2/s, ε is void fraction, dimensionless, τ is tortuosity factor, dimensionless, C AS and C AC are, respectively, concentration at pellet surface and concentration at pellet center, mol/dm3 (assumed to be the same as bulk concentration and zero respectively.27 ΔHrxn is heat of reaction, kJ/mol, λeff is effective thermal conductivity, kJ/(m·s·K). The value of DAB was found to be 1.16 × 10−4 m2/s at a temperature of 773 K using the Brokaw equation.28 The effective diffusivity Deff obtained was 1.01 × 10−5 m2/s. The void fraction, ε, estimated as the ratio of volume occupied by voids to the total bed volume was 0.5 calculated using the well-known equation εp = 0.38 + 0.073[1 + ((d/dp − 2)2/(d/dp)2)], where d and dp are the internal reactor diameter and particle diameter, respectively. The tortuosity factor was taken as 8.30 The effective thermal conductivity λeff was obtained from the wellknown correlation λeff/λ = 5.5 + 0.05NRe. A value of 5.47 × 10−5 kW/m/K for λ, which is the thermal conductivity, was obtained from the Wassiljewa correlation.28 The effective thermal conductivity λeff calculated was 8.096 × 10−4kW/m/K. Substituting these values in eq 15 resulted in a value of 0.1 K for ΔTmax, which shows the uniform temperature in a catalyst pellet. To determine the heat transfer limitation across the gas film, the following correlation adopted from the report in Kumar and Idem30 was employed ΔTfilm,max =

−r′CO,obs ρC R C2

as r′CO = k e(−E / RT )NCOmNH2OnNH2 oNCO2 p(1 − β) β=

NCO2 × NH2 1 × Keq NCO × NH2O

ln Keq =

⎛ 4577.8 ⎞ ⎜ ⎟ − 4.33 ⎝ T ⎠

(21)

(22)

(23)

here r′CO = rate of reaction with respect to CO, mol/(g-cat·hr), k = pre-exponential factor, E = activation energy, kJ/mol, T = reaction temperature, K, R = molar gas constant, 8.314 J/mol/ K, Ni = molar flow rate of CO, H2O, H2, CO2, mol/h, m = order of reaction with respect to CO, n = order of reaction with respect to H2O, o = order of reaction with respect to H2, p = order of reaction with respect to CO2, β = degree of reversibility of reaction, Keq = reaction equilibrium constant. However, in this study, the ratio of H2 to CO2 was not varied and the ratio of H2O/CO was constant at 2.6. Consequently, it was necessary to drop the b, c and d terms because the species for which they are exponents are not linearly independent, thereby leaving the power law expression as follows

(17)

Upon substituting the required terms on the left-hand side of the equation, a value of 0.8 × 10−4 was obtained, which is far less than 0.15. Therefore, based on this result, the interparticle heat transfer resistance is considered to be absent. H


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Industrial & Engineering Chemistry Research r′CO = k × NCOm

method of analysis was used.27 For integral (plug flow) reactor, differential analysis gave us the following equation:

(24)

where k = Ae(−E/RT) is the reaction rate constant. 3.4.2.2. Experimental Rates of Reaction. The results of each set of operating conditions are presented in Figure 7. From the

−r′CO =

dXCO d(W /FCO,0)

(25)

From eq 25, we can see that the slope of XCO vs W/FCO,0 will give us the experimental reaction rate (mol/(g-cat·hr)) calculated at the exit of the exit of the reactor for a given W/ FCO,0. The XCO vs W/FCO,0 graphs were drawn for all the temperatures and reactor systems, resulting in Figure 7. Now, the experimental rates of the reaction were obtained from Figure 7 as the derivatives of the CO conversion vs W/FCO,0 curves after fitting with MATLAB. Here, NCO, NH2O, NH2 and NCO2 were calculated using the following equations: NCO = NCO,0 − X CONCO,0

(26)

NH2O = NH2O,0 − (XCONCO,0)

(27)

NH2 = NH2,0 + (XCONCO,0)

(28)

NCO2 = NCO2,0 + (XCONCO,0)

(29)

where NCO,0 = initial molar flow rate of CO (A), mol/h; NH2O,0 = initial molar flow rate of H2O (B), kmol/s; XA = CO conversion, from experimental data. There are typically some experimental errors associated with catalytic reactor experiments, especially those involving membrane reactors and involving components in the feed that may also be present in the product. Also, feeding different phases of feed such as gas and liquid into the reactor may also introduce some challenges and errors as well. In this study, as has been outlined, we have taken steps to ensure that these errors are minimized. 3.4.2.3. Parameters’ Estimation. The power law model parameters were estimated using a multinonlinear regression software (NLREG), which is based on the minimization of the sum of residual squares of the reaction rates according to Gauss−Newton and Levenberg−Marquardt algorithms. The values obtained for the parameters are presented in Table 1. To validate the models, the percentage average absolute deviation (AAD) between the predicted rate from the power law model and experimentally obtained rate was determined.

Figure 7. Conversion versus W/FA0 for all three systems using reformate gas and 0.5 g of 3Ni5Cu/CeZrCa catalyst.

results, it can be seen that in all three systems, increasing the W/FCO,0 generally results in an increase of CO conversion. However, the increase in conversion tends to slow down beyond 1.6 (g-cat·hr)/mol, because of thermodynamic limits imposed on the reaction under the prevailing conditions. It is also evident that increasing the temperature at a particular W/ FCO,0 results in an increase in CO conversion; however, the degree of increase in CO conversion between 450 and 500 °C is higher than that between 400 and 450 °C. At 450 °C, the increase in conversion with respect to catalyst weight time is slower than at other temperatures. This observation is further revealed in the reaction rates represented in Table S1 (Supporting Information) and Figure 7. More importantly, the results in Figure 7 show that the CO conversions obtained from the membrane reactor are higher than those of the packed bed reactor operating at high pressure and atmospheric pressure. To find out the rate of the water gas shift reaction experimentally by integral (plug flow) reactor, the differential

Table 1. Estimation of the Values of the Parameters of the Three Models parameter

MR

PBTR-HP

PBTR-atm

k (s−1) E (kJ/mol) M AAD (%)

3.66 × 106 99.95 0.23 7.39

1.17 × 1019 190.3 1.78 6.97

1.79 × 1013 136.9 1.27 5.33

The AAD is calculated as follows: AAD (%) =

|experimental rate − predicted rate| × 100% experimental rate (30)

A parity plot, as shown in Figure 8, is also used to depict how well the power law model fits the data. From the parity plot, the rates determined using the power law model matches the experimentally determined rate. Thus, the final power law rate expressions for the three systems are given as follows: I


Article


knowledge, this appears to be the first time that kinetic studies on WGSR has been performed on reformate gas from dry reforming of biogas or natural gas in different systems including a membrane reactor and on the 3Ni5Cu/CeZrCa catalyst.

4.0. CONCLUSIONS A palladium based membrane reactor was employed in this study in order to evaluate the potential to minimize the number of process units utilized in an industrial hydrogen production process. Membrane characterization was performed to evaluate the selectivity of the membrane to H2 and then derive a correlation between the amount of H2 permeated and temperature as well as partial pressure. There was an increase in H2 recovery and flux with increasing partial pressure from 100 to 300 psi (6.8−20.4 atm) and temperature ranging from 300 to 500 °C. The H2 permeation activation energy was 4.4 kJ/mol. The membrane reactor was applied to the water gas shift reaction using the highly active 3Ni5Cu/CeZrCaOx HT-WGSR catalyst on reformate gas at 500 °C varying the pressure between 150 and 250 psi (10.2−17 atm). Increasing the reactor pressure from 150 psi (10.2 atm) to 200 psi (13.6 atm) resulted in an increase in CO conversion. Further increase in reactor pressure beyond 200 psi (13.6 atm) did not lead to any significant increase in performance. A comparison of the performance of the membrane reactor to the packed bed tubular reactor operating at high pressure and atmospheric conditions showed a superior performance in the membrane reactor. Parametric studies performed on the membrane reactor by varying the catalyst weight between 0.25−1.0 g, flow rate between 161 and 275 mL/min and steam/CO ratio of 1.7−3.7 showed that increasing the catalyst weight resulted in an increase in CO conversion up to 0.5 g, beyond which no significant change was observed. Reducing the flow rate resulted in an increase in the amount of hydrogen recovered while slightly improving the CO conversion. Varying the steam/CO ratio gave similar results to the packed bed tubular reactor. For the kinetics of the high-temperature water gas shift reaction using reformate gas studied in a packed bed reactor at atmospheric and high pressure as well as in a membrane reactor, the activation energies obtained from the power law model were, respectively, 137.2, 190.3 and 99.8 kJ/mol with good AAD and parity plot.

Figure 8. Parity plot for the reaction rates of all three systems.

For the membrane reactor: r′CO = 3.66 × 106e(−99950/ RT )NCO0.23

(31)

For the high pressure packed bed tubular reactor: r′CO = 1.17 × 1019e(−190300/ RT )NCO1.78

(32)

For the atmospheric pressure packed bed tubular reactor: r′CO = 1.79 × 1013e(−136900/ RT )NCO1.27

(33)

The results show an increase in activation energy from 137.2 to 190.3 kJ/mol when the operating pressure is increased from atmospheric to a high pressure of 200 psi (13.6 atm) with a corresponding increase in the pre-exponential factor. This is probably because with higher partial pressures of CO in the feed mixture, more energy is required to initiate the WGSR and there is the potential for a higher frequency of collision of the molecules. In contrast, at a similar high pressure of 200 psi (13.6 atm), the activation energy required for the membrane reactor is 99.95 kJ/mol, which is lower than that of the packed bed reactor even at atmospheric pressure; this re-emphasizes the benefits of using a membrane reactor. The order of the reaction with respect to CO varies among the different systems with the membrane system giving the lowest order of 0.23 and the high pressure PBTR system having the highest order of 1.78, showing a stronger dependence of the reaction rate on the concentration of CO. The reaction order for the atmospheric PBTR is 1.27, which is similar to other studies in the literature, which show a first-order dependence of WGSR reaction rate on CO. The activation energies obtained are observed to be in agreement with some studies in the literature. Mars33 reported an activation energy of 134 kJ/mol operating at atmospheric pressure, whereas Chinchen et al.34 reported 129.4 kJ/mol as the activation energy. Other studies have reported lower activation energies; for instance, Lei et al.35 reported 63 kJ/mol and Koryabkina et al.36 obtained 79 kJ/mol. The low activation energies obtained from such studies could be as a result of either the low concentrations or pure CO employed as the reactant without considering the presence of other components in the feed material such as CH4, CO2 and H2, which give a typical composition of a real reformate gas. According to Chinchen et al.,34 another reason for the low activation energies reported in the literature for the WGSR is due to the inadequacy of pore diffusion. Few studies, however, have considered the effect of these components, particularly for reformate from coal gasification.37 To the best of our

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ASSOCIATED CONTENT

S Supporting Information *

Experimental kinetic data and actual reformate gas flow rate and mass balance packed bed reactor data. This material is available free of charge via the Internet at http://pubs.acs.org.

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AUTHOR INFORMATION

Corresponding Authors

*H. Ibrahim. E-mail: [email protected]. Ph: 1306-337-3186. Fax: 1-306-585-4855. *R. Idem. [email protected]. Ph: 1-306-585-4470. Fax: 1-306-585-4855. Notes

The authors declare no competing financial interest. J


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ACKNOWLEDGMENTS The authors acknowledge Hydrogen Canada (H2CAN) strategic network for their financial support in collaboration with Natural Sciences and Engineering Research Council of Canada (NSERC).

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Kinetic Study of Hydrogen Production by the High Temperature Water

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