Development of a RhZrO2 catalyst for low temperature autothermal reforming of methane in membrane reactors

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Development of a RhZrO2 catalyst for low temperature autothermal reforming of methane in membrane reactors L. Marra, P.F. Wolbers, F. Gallucci, M. van Sint Annaland ∗ Chemical Process Intensification, Department of Chemical Engineering and Chemistry, Eindhoven University of Technology, Box 315, Eindhoven, The Netherlands

a r t i c l e

i n f o

Article history: Received 5 August 2013 Received in revised form 21 October 2013 Accepted 23 October 2013 Available online xxx Keywords: Rh/ZrO2 catalyst SMR ATR Kinetics Membrane reactor Microreactor

a b s t r a c t A Rh-based catalyst for low temperature hydrogen generation in membrane microreactor applications has been developed and characterized. A RhZrO2 catalyst with 1.4 wt% Rh was prepared by incipient wetness impregnation and was tested for both methane reforming and autothermal reforming at temperatures interesting for membrane reactor applications (i.e. temperatures below 700 ◦ C and steamto-carbon ratio of 2). The kinetic parameters to describe the reaction rate of both methane steam reforming (SMR) and auto-thermal reforming (ATR) over the RhZrO2 catalyst have been determined using a 1D heterogeneous packed bed reactor model to properly account for mass and heat transfer resistances. The experimental results demonstrate that the RhZrO2 catalyst is extremely active for ATR and resistant to coke formation at much lower temperatures and steam-to-carbon ratios compared with conventional Ni-based catalysts. This makes the new catalyst especially suitable for integration in a Pd-based membrane microreactor with a maximum allowable operation temperature of about 650 ◦ C dictated by the membrane stability. © 2013 Elsevier B.V. All rights reserved.

1. Introduction Hydrogen represents the ideal energy carrier for fuel cell applications, however an efficient process for ultra-pure (CO-free) hydrogen production is of utmost importance [1–3]. Traditional ways to convert natural gas (the main source for hydrogen worldwide) to pure hydrogen include the application of several reaction and separation steps run in sequence. These steps include desulfurization of the hydrocarbon feedstock and steam methane reforming (SMR) for syngas production, which is carried out at high temperatures (over 900 ◦ C), followed by CO cleanup, such as water–gas shift (WGS) and preferential oxidation (or by pressure swing adsorption), and hydrogen purification. The global efficiency of the traditional steam reformer plants, calculated as the ratio between the net heat value of the hydrogen stream produced and the total fuel feed, is typically within the range of 65–85%, depending on plant size. The drawbacks of conventional systems where the reformer is a multi-tubular fixed bed reactor are a significant fuel consumption to heat the reactor, the high heat transport resistance of packed bed reactors, which causes a large temperature decrease in the packed bed (because of the endothermic reaction) and a corresponding decrease of methane conversion

∗ Corresponding author. Tel.: +31 40 2472241. E-mail address: [email protected] (M.v.S. Annaland).

(because of the thermodynamic equilibrium) and the complex control of the many heat integration units [4]. The different operating conditions and the required inter-stage cooling make the process of such systems quite inefficient for low scale applications where excess heat cannot be used efficiently as in larger plants. A membrane reactor has been proposed as a promising concept for process intensification where reaction and separation are combined to overcome the equilibrium limitations. The expected benefits of this new technology applied to hydrogen production are the following: (1) increase in the overall efficiency of the process due to integration and intensification; (2) production of hydrogen with very high purity due to the infinite perm-selectivity of the membrane for this gas component; (3) large reduction in process units such as heat exchangers; (4) reduction of the reformer operating temperature (<700 ◦ C); (5) easier recyclability due to the reduction of components and materials in the reactor. The lower reforming operating temperature and the lower steam-to-carbon ratio (SCR) required by membrane reactors are the main boosters of the global efficiency increase compared with conventional systems [5]. An efficient reaction, which can be realized in the membrane reactor, is the auto-thermal reforming (ATR) of methane, which has attracted considerable attention in the last two decades [6–8]. In ATR the heat required by the reforming reaction is supplied by internal oxidation; consequently, there is no need to supply heat from an external source. The overall chemical reactions taking place

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in the ATR of methane include steam reforming (Eq. (1)), water gas shift (Eq. (2)), and total oxidation (Eq. (3)). The energy generated from the oxidation reaction and WGS is used for the SMR. CH4 + H2 O  CO + 3H2 CO + H2 O  CO2 + H2

H r = 206.10 kJ/mol H r = −41.15 kJ/mol

CH4 + 2O2 → CO2 + 2H2 O H r = −802.30 kJ/mol

(1) (2) (3)

All the advantages of membrane reactor technology, applied at small scale for instance in a microreactor, make the overall process more efficient, and may allow avoiding all the mass transfer limitations affecting other reactor configurations. In a microchannel, indeed, the small distance between the catalyst and the hydrogen selective membrane circumvents bulk-to-wall mass transfer limitations (concentration polarization). Advances in micro-catalytic reactors depend crucially on the catalyst and membrane development. A very active (and stable) catalyst, and a highly permeable (but perm selective) Pd membrane, are indispensable in order to achieve the required membrane area per unit volume of catalyst which match the H2 formation rate [9,10]. In particular, the catalyst properties, e.g. morphology, activity and stability, and the way to deposit and activate the catalytically active sites are inseparable from the choice of the reactor geometry. In a microchannel the amount of catalyst which can be deposited is very low, thus as consequence, the optimization of the catalytic activity must be done as a function of channel diameter and membrane permeability. The state-of-the-art in catalytic reforming for the production of hydrogen mainly concerns the high temperature range (typically above 900 ◦ C). Commercial catalysts have been designed to face most constraints of the process at these conditions. Nickel-based catalysts are traditionally employed for methane reforming reactions especially due to their low cost, but require high temperatures and a high steam-to-carbon ratio to avoid carbonaceous deposits. Moreover, the conventional Ni-based catalysts, e.g. Ni supported on ␣-Al2 O3 [11], seems unsuitable for the ATR process because it requires additional peripheral components, to prevent deactivation of the catalyst [12]. Additionally, the main advantages of membrane reactors, namely operation at much lower temperatures and lower steamto-carbon ratio, pose great challenges for conventional catalysts for SMR and ATR, because at these conditions the catalysts are much less active and more prone to catalyze carbon formation. Thus, the integration of membrane technology in the auto-thermal reforming process requires the development of novel catalysts by exploring the catalytic behavior of new metal nanoparticles and new oxide supports, with a high activity and a good selectivity at much lower temperatures than in conventional catalytic reformers. Qi et al. [13] reported that Rh supported on CeO2 and/or ZrO2 is a good candidate for the reforming of hydrocarbons because of its high activity and stability in the ATR reaction compared to other precious metals and supports. Ligthart [14] investigated the catalytic activity of Rh for the SMR reaction, with different metal loadings on several oxides supports. Yokota et al. studied the dry reforming (using methane and carbon dioxide) on Rh supported on various metal oxides [15]. In this work we have developed a RhZrO2 catalyst for autothermal reforming at low temperatures (below 700 ◦ C) and demonstrate its high activity and stability. First, the synthesis and characterization of the RhZrO2 catalyst is described. A thermogravimetric analysis of this catalyst is carried out and the results are compared with a conventional NiAl2 O3 catalyst, which show the bad performance of the conventional catalyst for low temperature reforming and the superior stability and low carbon formation of the new catalyst. Afterwards, a stability test of the RhZrO2 catalyst

is carried out and the results of a detailed study on the kinetic rates of SMR and ATR of methane over this catalyst is described. An experimental demonstration of the high and stable performance of this catalyst at lower temperatures (550–700 ◦ C) and lower steam-tocarbon ratio (SCR of 2) is given and its suitability for micro-reactor applications is assessed. 2. Catalyst preparation and characterization methods 2.1. Incipient wetness impregnation of Rhodium in a Zirconia support The supported Rh catalysts have been prepared by pore volume impregnation using aqueous solutions of Rh(NO3 )3 ·nH2 O (Riedel de Haën, purity 99.9%) of appropriate concentration. The support material, Zirconia (GIMEX, type RC-100 with 99.74% ZrO2 and 0.13% TiO2 ) has been sieved into a fraction of 125–250 ␮m. Prior to impregnation, the support has been calcined in a gas mixture of 20 vol% O2 in N2 at a flow rate of 100 ml/min, while being heated at a rate of 2 ◦ C/min to the final temperature followed by an isothermal period of 4 h. The impregnated support has been dried for 3 h in air and at 110 ◦ C overnight before further treatment. The estimated metal loading is about 1.6 wt% Rh. The catalyst precursor has been calcined at 600 ◦ C. 2.2. Catalyst characterization methods Brunauer–Emmett–Teller analysis (BET) provided precise Specific Surface Area (SSA) and pore volume evaluation of materials by nitrogen multilayer physisorption at −195 ◦ C, measured as a function of relative pressure using a fully automated Micromeritics TriStar 3000 BET apparatus after outgassing the sample for 3 h under vacuum at 150 ◦ C. Elemental analysis: The metal loading has been determined by inductively coupled plasma atomic emission spectroscopy (ICPAES) analysis, performed on a Goffin Meyvis SpectroCirus ccd apparatus. A solution of 5 M (NH4 )2 SO4 in H2 SO4 was employed to extract Rh from the ZrO2 -containing catalysts. An amount of sample was stirred in the acid under heating until a clear solution was obtained. Transmission Electron Microscopy (TEM): TEM images were acquired on a FEI Tecnai 20 transmission electron microscope at an acceleration voltage of 200 kV with a LaB6 filament. Typically, a small amount of grinded sample was reduced at 500 ◦ C and passivated in 1 vol% O2 in He for 2 h before being suspended in pure ethanol, sonicated and dispersed over a Cu grid with a holey carbon film. TEM images were recorded using a 1k × 1k Gatan CCD camera at different magnifications. From the electron micrographs, the metal nanoparticle diameters were determined from the projected area of the particles assuming that the particles are spherical. The particle size distribution has been determined from analysis of around 100 (for systems with low contrast i.e. relatively small Rh particles supported on oxides with similar atomic number) up to 300 particles (e.g. Rh/ZrO2 and aged systems) from at least three different micrographs. Thermo gravimetric analysis (TGA) is a very commonly used technique for studying reactive gas–solid systems. For this work a TGA system was used to measure the mass change of the catalyst during reforming reactions, with which (including a cross check with a combustion step) the amount of carbon formation was determined. In the setup, 50 mg of RhZrO2 catalyst were placed in a porous quartz sample holder, which was connected to a balance with a platinum wire. For the TGA experiments, first the catalyst was oxidized at 500 ◦ C for 1 h in 3 vol% O2 in N2 and subsequently reduced at 700 ◦ C for 2 h in 20 vol% H2 in N2 . Cooling and heating

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steps were carried out in nitrogen. SMR was performed at a temperature of 700 ◦ C with a feed containing 20 vol% CH4 and SCR of 2 (for SMR) at a total pressure of 1.4 bar. The total feed flow rate was 0.02 mol min−1 . Steam was supplied by evaporation of deionized water in a Controlled Evaporator Mixer unit in combination with a liquid-flow controller (Bronkhorst) and gas flow rates were controlled by mass flow controllers (Brooks). All tubings were kept at 175 ◦ C after the point of steam introduction to avoid condensation. The weight of the sample, during the reaction, increased due to carbonaceous deposits (because of CH4 decomposition and due to the Boudouard reaction). The deposited amount of coke was measured by weighing the sample before and after its complete oxidation.

2.3. Kinetics setup In a dedicated kinetics set-up the catalytic activities for steam and autothermal reforming were determined. A quartz U-shape tubular reactor was filled with catalyst and inert material, and then placed in a oven containing sand (which behave as a fluidized bed) to ensure virtually isothermal operation of the reactor. Typically, 30 mg of catalyst (sieved to 125–250 ␮m) were mixed with inert quartz particles (with particles diameters dp = 1 mm) in a ratio ‘catalyst:quartz’ = 1:5, to obtain a bed height of about 5 mm and to avoid problems in term of high pressure drop. Prior to the activity measurements, the catalyst was oxidized at 500 ◦ C for 1 h in 3 vol% O2 in N2 and subsequently reduced at 700 ◦ C for 2 h in 20 vol% H2 in N2 . Cooling and heating steps were carried out in nitrogen. The composition of the effluent gas was analyzed by online micro-GC (CP-4900 series from Varian b.v.) with two mol-sieve (5 A) columns and a poraplot Q (PPQ) column. For all the experimental results reported the carbon balance was closed below 2% error. SMR was carried out at a temperature range of 500–700 ◦ C with a feed containing 40 vol% H2 O and a steam-to-carbon ratio SCR of 2–3 in N2 at a total pressure of 1.4 bar. The total feed flow rate was set in the range of 0.02–0.08 mol min−1 . ATR was carried out at a temperature range of 500–700 ◦ C with a feed containing 20 vol% CH4 and SCR of 2 and an oxygen-to-carbon ratio OCR of 0.44 in N2 at a total pressure of 1.4 bar. The total feed flow rate was set to 0.069 mol min−1 for temperatures up to 550 ◦ C and 0.3 mol min−1 for temperatures above 600 ◦ C. Steam is supplied by evaporation of deionized water dosed via a HPLC pump. All tubings were kept at 200 ◦ C after the point of steam introduction to avoid condensation. A cold trap was placed after the reactor outlet and before the pressure controller and the GC to prevent water into the analysis system.

3. Thermodynamic and kinetic model 3.1. Thermodynamic analysis A thermodynamic analysis of SMR and ATR reaction has been carried out first, to determine the influence of the key operational variables on the chemical equilibrium, by using an in-house code, based on the minimization of the Gibbs free energy method. In the case of SMR, Seo et al. [16] evaluated the equilibrium conditions at SCR of 1, where carbon generation plays a role. They found that a minimum value of SCR = 1.4 is necessary to prevent carbon formation. Consequently a SCR ≥ 2 was set for further analysis. Fig. 1(a) and (b) shows the CH4 conversion vs temperature for two different SCRs and the H2 , CO, and CO2 yields, respectively. As expected, the reforming process requires very high temperatures (even for the conditions investigated, i.e. high N2 dilution) when the process is carried out in a traditional reactor. Moreover

Fig. 1. (a) Equilibrium conversion of methane for SCR 2 (blue line) and 3 (red line) for SMR, (b) Yield of H2 (black line), CO (red line) and CO2 (blue line) for SCR 2 and (c) yield of H2 (black line), CO (red line) and CO2 (blue line) for SCR 3. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.)

at high temperature the WGS reaction is not effective, with a low CO conversion. Concerning the ATR reaction, the addition of the oxygen into the feed increases the CH4 conversion and the H2 yield, preventing carbon formation also at low SCR. By assuming complete WGS, the auto-thermal reforming reaction can be expressed as: CH4 + O2 + 2(1 − )H2 O  CO2 + 2(2 − )H2

(4)

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TOX (Eq. (3)) is given by:

0.3 388 1.250

rTOX =

0.3 384

1.245

0.3 382

1.240

0.3 380 1.235 0.3 378 1.230 0.3 376 OCR SCR

0.3 374 500

550

60 0

650

1.225 700

750

800

T Temperature (°C) Fig. 2. Overall feed ratios (OCR and SCR) as a function of temperature to reach autothermal operation.

where  represents the oxygen to carbon ratio OCR. The condition which has to be satisfied to ensure auto-thermal operation is that the net heat produced HiT is equal to zero: nc 



vi HiT = 0

(5)

i=1

Thus, for a selected temperature, Eq. (5) can be solved for  using Eq. (4) and the resulting overall OCR and SCR ratios are shown in Fig. 2. Using overall OCR and SCR of approximately 0.38 and 1.24, respectively, and a feed temperature of 650 ◦ C, autothermal operation of ATR can be achieved, provided that complete CH4 conversion and complete water gas shift are realized. 3.2. Kinetic model To properly design a reactor, and a (micro-) membrane reactor in particular, detailed information on the intrinsic reaction rates, i.e. reaction orders, reaction rate constants and activation energy is indispensable. The intrinsic kinetics of SMR has been investigated over the past decades, where different mechanisms and different kinetic models have been proposed for different catalysts by several research groups. The kinetic expressions for SMR (Eq. (6)) and WGS (Eq. (7)) developed by Numaguchi and Kikuchi (NK) [17] for Ni catalyst supported on Al2 O3 have been used in literature frequently.

 rSMR = kSMR

3 · P )/K PCH4 − (PH CO eq,SMR



 rWGS = kWGS

PCO −

2

0.596 PH 2O

PH2 PCO2 Keq,WGS





kTOX,a · PCH4 · PO2 ox 1 + KCH 4

· PCH4 + KOox 2

k

·P

·P

CH4 O2 TOX,b 2 + 1 + K ox · P + K ox · P CH O2 4 CH4 O2 · PO2

(8)

Kjox

SCR

OCR

0.3 386

 (6)

 (7)

where ri represents the reaction rate [mol/kg/s], ki the kinetic rate constants, Pj the partial pressures of the gas components and Keq,i the equilibrium constants (for i = SMR,WGS). When steam reforming and water gas shift are combined with methane oxidation, the overall reaction is defined as auto-thermal reforming. The ATR reaction kinetic expression considers the Eqs. (6) and (7) for SMR and WGS and the kinetic equation from Trimm and Lam [18], which describes the combustion of methane in the packed bed reactor. In their analysis they determined the k0 and the Ea for Pt catalyst supported on Al2 O3 . The expression valid for

where represents the adsorption coefficient for component j [1/bar]. The kinetic experiments have been performed at different operating conditions using a relatively high gas velocity and a low amount of catalyst, to ensure a low methane conversion (below 10%). The kinetics reactor consists of porous catalyst pellets in a small packed bed through which the reactants percolate. The question arises whether the conversion is determined by the activity of the catalyst itself (kinetic regime) or is limited by transport phenomena (diffusion regime). The kinetic analysis has to be carried out at particular experimental conditions, to be sure that the mass transfer can be neglected. The effect of intra-particle mass transfer limitations is strongly related to the catalyst particle size and is evaluated by calculating the overall effectiveness factor h for the reactions, which is defined as the ratio of the integrated reaction rates over the radius of the particle and the reaction rate at bulk phase conditions [19]. During the kinetic tests, it has been found that the effect of internal diffusion is negligible for a particle size smaller than 125–250 ␮m as will be demonstrated in the next sections of the paper. Concerning the external mass transfer contribution, the Mears criterion [20], given by Eq. (9), indicates whether the methane conversion is unaffected by the external diffusion limitations. rCH4 cCH4 kc ac

=

kexp ≤ 0.15 kc∗

(9)

where rCH4 is methane reaction rate [mol/(m3 s)], cCH4 the methane concentration [mol/m3 ], kc the mass transfer coefficient [m/s], ac is the external surface-to-volume ratio of the particle [m2 /m3 ], kexp is the experimental kinetic rate constant [1/s] and kc∗ is the intrinsic mass transfer coefficient [1/s]. From the experiments the kexp was determined, by Eq. (10): kexp =

rCH4 cCH4

=

(yCH4 tot CH4 (1 − ε))/mcat (yCH4 p)/RT

(10)

where yCH4 is the methane molar fraction, tot the total molar feed flow [mol/s],  the catalyst density [kg/m3 ], mcat the catalyst mass [kg], ε the bed porosity, CH4 the methane conversion, p the feed pressure, R the universal gas constant [J/mol/K], T the gas temperature in the bed [K]. The intrinsic mass transfer coefficient kc∗ [1/s] is given by: kc∗ = kc ac = Sh

D ac dp

(11)

with D is the binary diffusivity [m2 /s], dp is the particle diameter [m], Sh is the Sherwood number, that for a packed bed reactor is described by the following expression: Sh = 2 + 1.5Sc 1/3 (1 − ε)Re1/2

(12)

where Re is the Reynolds number and Sc is the Schmidt number. In the case of SMR the Mears criterion is satisfied as shown in the following sections. For the ATR reaction, the setup limitations did not allow to work in kinetic regime. 3.3. Reactor model A more detailed analysis has been carried out when the system is operating at diffusive regime. In this situation, concentration and temperature gradients affected by axial dispersion, and mass and

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Table 1 Governing equations for the 1D model used in this work. Component (mass) balance for the gas phase

εg g

∂ωi,g

= −g g

∂ωi,g

∂x (εg g Cp,g + εs s Cp,s ) ∂∂Tt

Pseudo-homogeneous energy balance for both gas and solid

∂t

− dp = 150 dx

Momentum balance (Ergun equation)

g g (1−εg )2 2 dp

ε3 g

+ =

3

∂ωi,g ∂  D + εg ri Mi with ri = εg,p · ∂x g ax ∂x −s g Cp,g ∂∂Tx + ∂∂x eff ∂∂Tx + εg ri HR,i

+ 1.75





j=1

i · rj

− ˛ d4 (T − Tenv ) r

g g2 1−εg dp ε3 g

Constitutive equations for transport properties Re2 ·Pr 2 · g

Re·Pr· g + 6(1−ε )Nu Peax g 2(0.17+0.33·e−(24/Re) )

Effective heat dispersion

eff = bed,0 +

Expression for Peclet number

Peax =

Expression for Nusselt number

Nu = 7 − 10εg + 5ε2g

Effective dispersion

Deff =

Thiele modulus Effectiveness factor

= =

 1−(0.17+0.33·e−(24/Re) )   · (1 + 0.7Re0.2 Pr 1/3 ) + 1.33 − 2.4εg + 1.2ε2g · Re0.7 Pr 1/3 

Table 2 BET analysis of the ZrO2 support and supported Rh catalyst before and after the SMR reaction (SCR 2 and T = 700 ◦ C). Sample

Pore volume (cm3 g−1 )

BET surface area (m2 g−1 )

Calcined ZrO2 1.4 wt% Rh/ZrO2 1.4 wt% Rh/ZrO2 after the experiments

0.2419 0.2188 0.1546

64 57 37

Table 3 ICP analysis to determine the metal loading of Rh in the ZrO2 support. Sample no. #

Cav (RhZrO2 ) (mg ml−1 )

Cav (Rh) (mg ml−1 )

Rh (wt%)

1: RhZrO2 2: RhZrO2

0.130 0.114

1.786 × 10−3 1.526 × 10−3

1.37 1.34

heat transfer (external and internal) start to become relevant. In order to estimate the intrinsic reaction kinetics for the developed RhZrO2 catalyst for ATR, an heterogeneous 1D reactor model was developed. To account for heat and mass transfer limitations the governing equations are listed in Table 1. Note that heat losses through the walls have been accounted for using an overall heat transfer coefficient ˛· 4. Results and discussion 4.1. Characterization of the catalyst Surface properties of the RhZrO2 catalyst have been evaluated with BET, ICP-AES and TEM analysis. Three sets of BET analysis have been carried out, which includes the ZrO2 support and supported Rh catalyst before and after carrying out the SMR reaction at SCR 2 and T = 700 ◦ C summarized in Table 2. From Table 2 it can be concluded that the addition of catalyst by incipient wetness impregnation decreases both the pore volume and surface area about 10% most probably because of structural changes during the successive calcination rather than pore blocking because of the low amount of catalyst loaded. The samples measured after the SMR experiment show clearly a further decrease in pore volume (29% lower than the value before the reaction) and about a 35% decrease in SSA (compared to the second case), indicating a reduction of accessible active sites and a corresponding decrease in the activity of the catalyst. ICP-AES analysis provides information about the Rh loading in the ZrO2 support. The sample consisted of 65 mg of RhZrO2 , which was solved in a 5 ml solution, prepared by 20 g (NH4 )2 SO3 in 30 ml H2 SO4 . The sample is then further diluted in deionized water, to reach a total solution amount of 50 ml. Table 3 shows the results

0.73 Re·Sc

+

k∗

0.5 εg +(9.7ε2 )/(Re·Sc) g

g · dp

dp r · D 6 1 − 12 ·tanh(3·) 3·

of the ICP analysis on two set of samples. The amount of Rhodium dispersed in Zirconia is about 1.4 wt%. TEM analysis has been performed to estimate the size of the Rh particles in the zirconia support and to evaluate the distribution of the metal nanoparticles. The sample has been prepared by diluting some catalyst in ethanol, vigorously stirring and then dropping it onto a copper grid, which was dried in air before the analysis. From the images (a–c) of Fig. 3 it is possible to recognize the zirconia particles; in figure c) some Rhodium particles can be distinguished too (within the indicated cycle). Since the amount of Rh particles detected with the microscope was very low, an Energy Dispersive X-ray (EDX) analysis was made to verify the presence of Rh particles in the sample composition. The spectrum of Fig. 3d indeed confirms this. The TGA analysis was carried out in a High Pressure TGA, to evaluate the performance of the RhZrO2 catalyst with respect to carbon formation and other possible deactivation processes during SMR. For this analysis the 1.4 wt% RhZrO2 was compared with a conventional 20 wt% NiAl2 O3 catalyst, kindly provided by HYBRID Catalysis B.V., both at a relatively high temperature. 50 mg of sample was put in a porous quartz crucible (the porosity prevent any mass and heat transfer limitation), which was suspended into the oven and connected mechanically to a micro-balance, able to detect small changes in weight of the sample. The experiment consisted of three stages: (1) a reduction process (to activate the catalyst), where the feed mixture, consisting of 20% of hydrogen in nitrogen, was injected at a total flow rate of 0.02 mol min−1 for 2 h; (2) the SMR reaction at 700 ◦ C, with a feed composition of S:C:N = 2:1:1.35, which was fed at a flow rate of 0.069 mol min−1 for 3 h; the oxidation of the catalyst, which is performed in air at 0.02 mol min−1 for removal of possible carbon deposits during the SMR reaction. Fig. 4 shows the profiles of the weight change vs time of the two catalysts. The red line, related to the Ni catalyst, grown rapidly, indicating that the amount of carbonaceous deposits generated is very high; indeed the ratio gcoke /gcat = 1.57 in less than half an hour. The blue curve corresponds to the Rh catalyst and shows a low carbon formation, with a ratio gcoke /gcat = 0.045 in 24 h. This analysis confirms that the conventional Ni catalyst in indeed unsuitable for reforming at lower temperatures (below 700 ◦ C). 4.2. Stability tests Before performing a kinetic study of the RhZrO2 catalyst, stability tests for the SMR and ATR reactions have been carried out in the kinetics setup, to determine the evolution over time of the catalyst activity. These data are then used for the kinetic tests, ensuring a stable operation and allowing to determine the kinetic rates of the reactions. During these experiments, the SMR reaction operates at 700 ◦ C, for 16 hours. A total flow rate of 0.069 mol min−1 was used with a SCR of 2 (23 vol% of CH4 in N2 flow). Fig. 5(a) and (b) shows the

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Fig. 3. (a)–(c) TEM images of the 1.4 wt% Rh/ZrO2 catalyst and (d) EDX of the same sample.

Weight change [mg]

90 80

Steam [g/h] RhZrO2 sample [mg]

70

Steam [g/h]

conversion of CH4 and the yield of CO, CO2 and H2 as a function of time on stream. The methane conversion decreased by 40% in 24 h and then became stable. The decrease in activity is related to sintering and/or carbon deposition (as shown in the thermo-gravimetric analysis), and a consequently lower availability of active sites on the porous support. The BET analysis made on the sample, after this reaction, confirmed a lower pore volume and SSA (35% below the value observed before performing the reactions). The stability test for the ATR reaction was performed at 650 ◦ C for 6 h. A total flow rate of 0.069 mol min−1 was used with a SCR of 2 and an OCR of 0.44 (23 vol% of CH4 in N2 flow). Fig. 6(a) and (b) shows the conversion of CH4 and the yield of CO, CO2 and H2 as a function of time on stream. No deactivation was observed when the system was at auto-thermal operations, as the CH4 conversion and H2 , CO, CO2 yields all remain constant. This shows that the autothermal reforming of methane over the RhZrO2 catalyst is stable under these conditions.

NiAl2O3 sample [mg]

60 50 40 30 20 10 0 0

5

10

15

20

25

30

Time (h) Fig. 4. TGA results during SMR of RhZrO2 (blue line) and NiAl2 O3 (red line). T = 700 ◦ C and SCR = 2. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.)

4.3. Kinetics results The kinetic analysis of SMR and ATR was carried out in the kinetics setup at the operating conditions detailed in Section 2.3.

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Fig. 5. (a) Methane conversion vs time during SMR reaction (SCR 2 and T = 700 ◦ C) and (b) yield of H2 (green squares), CO (red circles) and CO2 (blue triangles) vs time. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.)

Fig. 6. (a) Methane conversion vs time during ATR reaction (SCR 2 OCR 0.44 and T = 650 ◦ C) and (b) yield of H2 (grey squares), CO (red circles) and CO2 (blue triangles) vs time. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.)

4.4. SMR kinetics

Assuming the same reaction rate expression as in Eq. (6), the kinetic parameters kSMR,0 and Ea,SMR can be determined from the experimental data assuming an Arrhenius-type law. An activation energy of Ea,SMR = 83.6 kJ/mol and an intrinsic kinetic rate constant

Before starting the SMR kinetic tests, the SMR reaction was carried out for several hours to ensure a stable activity of the catalyst. The reaction was tested at several flow rates (0.02–0.08 mol min−1 ) and at several temperatures (500–700 ◦ C), for SCR of 2 and 3. The conversion of methane at several operating temperatures at SCR = 2 is shown in Fig. 7. The black line indicates the equilibrium conversion, while the markers are related to the experimentally determined conversions at two different feed flow rates: low flow rate at 0.02 mol min−1 (red circles) and high flow rate at 0.08 mol min−1 (blue triangles). It can be seen that at a high flow rate, the conversion is lower as anticipated, due to lower residence times in the catalytic bed. Since the kinetic analysis can be done only if the system operates in kinetic regime, the Mears criterion was applied. A value of kexp /kc∗ = 0.08 has been found for the low flow rate, which is below 0.15 and indicates that the system has no mass transfer limitations for both sets of data. The experimental reaction rate, at low flow rate, is then selected to be compared with the theoretical NK reaction rate slightly modified by de Smet et al., which is valid for a conventional 20 wt% Ni/Al2 O3 catalyst [8]. In their model the SMR kinetics is described by equation 6. From Fig. 8 it is possible to observe that at the same reaction conditions (SCR 2) the RhZrO2 catalyst is considerably more active than the NiAl2 O3 catalyst.

Fig. 7. Methane conversion induring SMR with SCR = 2 at different temperatures and two different flow rates (low flow rate circles 0.02 ml/min; and high flow rate blue triangles 0.08 ml/min). The equilibrium conversion is indicated by the black curve.

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RhZrO2 Experimental SCR 2

1.2

NiAl2O3 Kinetic Model Reaction Rate (mol/kg/s)

1.0

0.8

0.6

0.4

0.2

0.0 450

500

550

600

650

700

Temperature (°C) Fig. 8. Reaction rate of SMR (SCR 2) as a function of temperature: experimental data (red markers) and theoretical curve of de Smet kinetic model (blue line). (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.)

Fig. 10. Comparison between the experimental (dots) and theoretical (lines) reaction rates at different temperatures. The theoretical curve is calculated from Eq. (6), by inserting the Ea and k0 values found for RhZrO2 catalyst.

Fig. 9. Linearized Arrhenius plot. Fig. 11. Parity plot of The experimental reaction rate vs the theoretical one.

= 97 × 103

mol/(kg s bar0.404 ) have been found. The activation

kSMR,0 energy is clearly much lower than the value found for the Ni catalyst (Ea = 106 kJ/mol) (Fig. 9). The determined kSMR is then substituted in Eq. (6) to describe the SMR reaction rates at different compositions; the thus computed reaction rates (lines) for SCR = 2 and SCR = 3, are then compared with the experimental data (markers) in Fig. 10. There is a good correlation between the experimental values and the theoretical ones, which means that the RhZrO2 catalyst can be described quite satisfactorily with the kinetic model by NK, given by equation 6. A parity plot is provided in Fig. 11. 4.5. ATR kinetics The kinetic analysis of ATR was carried out with SCR = 2 and OCR = 0.44, where the stability tests have demonstrated that the RhZrO2 catalyst is quite stable and carbon formation was negligible. Also the ATR reaction was tested at several flow rates (0.069–0.3 mol min−1 ) and at several temperatures (500–700 ◦ C). Fig. 12 shows the experimentally measured methane conversion at different operating temperatures at two different feed flow rates (markers): relatively low flow rate at 0.069 mol min−1 (red circles) and very high flow rate at 0.3 mol min−1 (blue triangles). The equilibrium conversion is indicated by black line. At both flow rates

Fig. 12. Methane conversion in ATR reaction (SCR 2 OCR 0.44) at several working temperatures. Equilibrium conversion (black line). Experimental conversion at High flow rate (blue dots) and Low flow rate (red dots). (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.)

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75 70

H2 Exp

65

H2 Model

60

CO Exp CO Model CO2 Exp

55 50

Yield (%)

45

CO2 Model

40 35 30 25 20 15 10 5 0 500

550

600

650

700

Temperature (°C) Fig. 13. Methane conversion for ATR reaction. Comparison between the experimental data (black dots) and the theoretical curve (red line) obtained with a 1D heterogeneous reaction model. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.)

the methane conversion is too high to guarantee differential reactor operation. It was not possible in the current set-up to further decrease the methane conversion (by increasing the feed flow rate or by decreasing the height of the catalytic bed). Therefore, an heterogeneous reactor model has been adopted to evaluate the kinetic rate constant values for the ATR reaction (Eqs. (6)–(8)). The kinetics by NK [17] was corrected and implemented with the newly determined kinetic parameters from the previous paragraph. In addition, the Trimm and Lam [18] expression for the total combustion of methane was implemented, and refitted to the ATR experimental results for the developed RhZrO2 catalyst. For the kinetic rate constants kWGS and kTOX , the activation energies determined by NK were taken and a good correlation between the experimentally determined reaction rate values and the theoretical curves obtained with the model, is found when the pre-exponential kWGS,0 and kTOX,0 constants are as shown in Table 4 and in Fig. 13. Note that, considering an average particle size of 190 ␮m and an internal particle diffusivity of 5E−5 m2 /s, an effectiveness factor = 0.98 is calculated (see Table 1), indicating that the internal mass transfer limitations can indeed be neglected. Inspecting the figure at temperatures below 550 ◦ C, it is clear that the model overestimates the methane conversion at low temperatures. This can most likely be explained by the fact that the kinetic rate expressions by Trimm and Lam is valid only for temperatures above 553 ◦ C. From Fig. 14 it is evident that there is also a good correlation between the experimental values of the H2 , CO, and CO2 yields at different temperatures, demonstrating that the kinetic model can correctly describe the prevailing reactions. 5. Discussion From the kinetic study it is possible to conclude that the RhZrO2 catalyst seems to be a promising candidate for low temperature ATR, due to the superior activity and stability and extremely low carbon formation compared to conventional NiAl2 O3 reforming catalysts. For application in a membrane microreactor, the activity of this material has to be sufficiently high to match the required conversion rate in the available volume for the catalyst in the microchannel to be tuned to the hydrogen permeation rate, which at the same time has to be designed such that bed-to-wall mass transfer limitations are avoided. So the mass transfer between the

Fig. 14. H2 (green), CO (blue) and CO2 (red) yields vs temperature. Comparison between the experimental data (dots) and the theoretical values (lines) determined with the heterogeneous 1D model. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.)

Fig. 15. Schematic of a micro-channel section: l represents the thickness of the catalytic coating [m], L the height of the channel [m], and A the membrane area [m2 ].

catalytic bed and the membrane has to be fast (and avoid concentration polarization), and consequently their distance small, which indicated a small dimension of the channel. Moreover, assuming that the catalyst will be deposited on the wall of the channel (and avoid large pressure drops of micro-packed beds), internal diffusion should be avoided by making the catalytic layer very thin. This means that the amount of catalyst available in the micro-channel is quite low. A schematic representation of the cross-section of a possible configuration of the micro-channel membrane reactor is presented in Fig. 15 where the geometric parameters involved in the analysis are highlighted. The pseudo-first order kinetic constant can be determined from the experimental results shown in the previous sections. Under the operating conditions used in the kinetic analysis and assuming negligible mass transfer limitations, where the 1.4 wt% RhZrO2 catalyst was used to perform the ATR reaction with SCR of 2 and OCR of 0.44 (T = 650 ◦ C, p = 1.4 bar, tot = 2.8 × 10−3 mol/s, Mcat = 30 mg, VPBR = 4 × 10−8 m3 , CH4 = 46%) using equation [10] a kR = 1.8 × 103 s−1 and a reaction time, defined as the reciprocal of the kinetic rate constant,  R = 5.5 × 10−4 s is determined. To verify whether the mass transfer limitations can be

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Table 4 Intrinsic kinetic rate constant and activation energy values for RhZrO2 (determined in this work) and the conventional NiAl2 O3 (used as reference). SMR

RhZrO2 NiAl2 O3

WGS

TOX

kSMR,0 , mol/(kg s bar0.404 )

Ea,SMR , kJ/mol

kWGS,0 , mol/(kg s bar)

Ea,WGS , kJ/mol

kTOX,0a , mol/(kg s bar2 )

kTOX,0b , mol/(kg s bar2 )

Ea,TOX , kJ/mol

9.7 × 104 2.6 × 105

83.6 106.9

17.2 × 102 2.5 × 102

54.5 54.5

56.7 × 105 8.1 × 105

47.7 × 105 6.8 × 105

86.0 86.0

neglected, the reaction time constant  R is compared with the internal and external diffusion time constants: - The characteristic internal diffusion time  int is defined as the time needed for the gas to diffuse through the catalyst layer at the micro-channel wall, and is given by: l2 /D, where l is the characteristic length for diffusion (in this case represented by the thickness of the catalytic coating) and D the effective gas diffusivity. A representative value of l = 10 ␮m and D = 10−5 m2 /s (assuming a washcoat layer porosity over tortuosity of 0.1), give a characteristic diffusion time value of 10−5 s. For these conditions the Thiele modulus (for a flat slab) can be estimated to be equal to 0.134, for which the effectiveness factor is 0.994. - The gas bulk-to-wall diffusion time  bulk-to-wall is defined as the time needed for the gas components to reach the catalyst layer at the channel wall. For the representative value of L = 100 ␮m (height of the channel) and D = 2 × 10−4 m2 /s (evaluated at reaction conditions) a Sherwood number can be estimated equal to Sh, which results in a bulk-to-wall diffusion time of 3.3 × 10−7 s. Concluding, for a channel diameter of 100 ␮m and a catalyst layer thickness of 10 ␮m, the reaction time  R is much higher than both diffusion times  int and  bulk-to-wall , so that indeed the external and internal mass transfer limitations can be safely excluded. In order to quantify the required activity of the catalyst in a membrane micro-reactor, a first order simplified analysis has been carried out in terms of a Fickian mass transfer coefficient and the pseudo-first order kinetic rate constant kR [s−1 ]. For this analysis, the highest membrane flux along the micro-channel should be in the same order of magnitude of the hydrogen production rate: JH2 Amem = RH2 Vcatalyst ≈ 4RCH4 Vcatalyst

(13)

For the selected micro-channel configuration (see Fig. 15) Vcatalyst /Amem = 3l so that RCH4 =

JH2 12l

(14)

The maximum hydrogen concentration in the micro-channel can be estimated at the thermodynamic equilibrium without membranes, that for the conditions considered in this paper is about 68%. By integrating a self-supported thin Pd-Ag membrane, with a permeance in the order of 1.5 × 10−5 mol/(m2 s Pa) (which corresponds to one of the most permeable membranes reported in literature [21]), the maximum flux of hydrogen is approximated at 0.76 mol/(m2 s) and correspondingly the methane reaction rate should at least be 6.34 × 103 mol/(m3 s). Taking the inlet and equilibrium methane concentrations, the required reaction time R2 for an optimal micro-channel reactor design should thus be between 1.2 × 10−4 s and 5.2 × 10−4 s. Comparing the required and the measured reaction times, R2 and  R , it is clear that the required activity of the RhZrO2 should be a factors 2–4 higher than the catalyst developed in this paper, which can be achieved by increasing the Rh loading to 2–6%.

6. Conclusions A Rh catalyst supported on ZrO2 has been prepared by incipient wetness impregnation for use in hydrogen production by autothermal reforming of methane in a (micro-) membrane reactor at significantly lower temperatures and SCRs than in conventional reactors. A chemical and physical characterization of the catalyst shows that the ZrO2 support has a high specific surface area and the Rh nanoparticles are well dispersed. A thermo-gravimetric analysis was carried out to quantify the carbon deposition rate at low SCR during methane reforming on the synthesized RhZrO2 and compare with a conventional NiAl2 O3 catalyst. The TGA results showed negligible carbon deposition on the RhZrO2 at low temperatures and low SCRs making it a suitable catalyst for low temperature ATR. A stability study of this catalyst for SMR and ATR has been carried out, showing in the first case that the Rh is stable after 20 hours of activity, even if the methane conversion decreased about 40% from the initial value. A more stable behavior was observed for the ATR reaction, where the methane conversion was quite stable in time. A kinetic analysis was carried out to determine the kinetics of the reactions. The SMR reaction rate was much higher than for NiAl2 O3 catalysts and could be well described by the kinetic model by NK after fitting the reaction rate constants where a much lower activation energy was found compared to the activation energy for the conventional Ni catalyst. For the kinetic analysis for the ATR, it was found that the reaction rates could be well described by the NK rate expression for the SMR and WGS and the Trimm and Lam expression for the TOX, however with fitting pre-exponential constants and activation energies. The experiments demonstrated the high catalytic activity of the developed RhZrO2 catalyst for the ATR reaction, typically about 7 times higher than the conventional NiAl2 O3 catalyst. The high activity and very stable (no deactivation and carbon formation detected) performance of this catalyst for ATR at low temperatures (550–700 ◦ C) and low steam-to-carbon ratios (SCR of 2) (in strong contrast to traditional NiAl2 O3 catalysts) makes this catalyst very suitable for application in membrane microreactors. Acknowledgements This project is supported by the European Union’s Seventh Framework Program (FP7/2007–2013) for the Fuel Cells and Hydrogen Joint Technology Initiative under grant agreement n◦ 278997. The authors thank Dr. A.J.J. Koekkoek and Dr. H.C.L. Abbenhuis from Hybrid Catalysis BV, Eindhoven, for their role in the elaboration of concept of catalysts used in this work and planning of the relevant joint studies, and acknowledge Lee McAlpine and Joris Garenfeld for the technical support. References [1] J. Ogden, Review of small stationary reformers for hydrogen production, Report to the International Energy Agency (609) (2001). [2] B. Ewan, R. Allen, A figure of merit assessment of the routes to hydrogen, International Journal of Hydrogen Energy 30 (8) (2005) 809–819. [3] M.A. Pena, New catalytic routes for syngas and hydrogen production (1996) 7–57.

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[4] M. De Falco, G. Iaquaniello, A. Salladini, Experimental tests on steam reforming of natural gas in a reformer and membrane modules (RMM) plant, Journal of Membrane Science 368 (1–2) (2011) 264–274. [5] L. Roses, F. Gallucci, G. Manzolini, S. Campanari, M. Van Sint Annaland, Comparison between fixed bed and fluidized bed membrane reactor configurations for PEM based micro-cogeneration systems, Chemical Engineering Journal 171 (3) (2011) 1415–1427. [6] S. Lee, D. Applegate, S. Ahmed, S. Calderone, T. Harvey, Hydrogen from natural gas: part I—autothermal reforming in an integrated fuel processor, International Journal of Hydrogen Energy 30 (8) (2005) 829–842. [7] A. Docter, A. Lamm, Gasoline fuel cell systems, Journal of Power Sources 84 (2) (1999) 194–200. [8] C.R.H. De Smet, M.H.J.M. De Croon, R.J. Berger, G.B. Marin, J.C. Schouten, Design ¨ of adiabatic yxed-bed reactors for the partial oxidation of methane to synthesis gas Application to Production of Methanol and Hydrogen-for-Fuel-Cells, vol. 56, 2001, pp. 4849–4861. [9] A.L. Mejdell, M. Jøndahl, T.A. Peters, R. Bredesen, H.J. Venvik, Experimental investigation of a microchannel membrane configuration with a 1.4 ␮m Pd/Ag 23 wt.% membrane—effects of flow and pressure, Journal of Membrane Science 327 (1–2) (2009) 6–10. [10] T. Boeltken, M. Belimov, P. Pfeifer, T.A. Peters, R. Bredesen, R. Dittmeyer, Fabrication and testing of a planar microstructured concept module with integrated palladium membranes, Chemical Engineering and Processing: Process Intensification 67 (2013) 136–147. [11] J. Zhang, Y. Wang, R. Ma, D. Wu, Characterization of alumina-supported Ni and Ni-Pd catalysts for partial oxidation and steam reforming of hydrocarbons, Applied Catalysis A: General 243 (2) (2003) 251–259.

11

[12] Z. Yuan, C. Ni, C. Zhang, D. Gao, S. Wang, Y. Xie, A. Okada, Rh/MgO/Ce0.5Zr0.5O2 supported catalyst for autothermal reforming of methane: the effects of ceria–zirconia doping, Catalysis Today 146 (1–2) (2009) 124–131. [13] A. Qi, S. Wang, C. Ni, D. Wu, Autothermal reforming of gasoline on Rh-based monolithic catalysts, International Journal of Hydrogen Energy 32 (8) (2007) 981–991. [14] D.a.J.M. Ligthart, R.A. van Santen, E.J.M. Hensen, Influence of particle size on the activity and stability in steam methane reforming of supported Rh nanoparticles, Journal of Catalysis 280 (2) (2011) 206–220. [15] S. Yokota, K. Okumura, M. Niwa, Support effect of metal oxide on Rh catalysts in the CH4 CO2 reforming reaction, Catalysis Letters 84 (2002) 131–134. [16] Y.-S. Seo, A. Shirley, S. Kolaczkowski, Evaluation of thermodynamically favourable operating conditions for production of hydrogen in three different reforming technologies, Journal of Power Sources 108 (1–2) (2002) 213–225. [17] K. Numaguchi, T. Kikuchi, Intrinsic kinetics and design simulation in a complex reaction network; steam-methane reforming, Chemical Enginnering Science 43 (8) (1988) 2295–2301. [18] L. Trimm, C. Lam, The combustion of methane on platinum–alumina fiber catalysts-II, Chemical Engineering Science 35 (1980) 1405–1413. [19] C.S. Patil, Membrane Reactor Technology for Ultrapure Hydrogen Production, Enschede, The Netherlands, 2005, pp. 235. [20] D.E. Mears, Tests for Transport Limitations in Experimental Catalytic Reactors, Industrial & Engineering Chemistry Process Design and Development 10 (4) (1971) 541–547. [21] F. Gallucci, E. Fernandez, P. Corengia, M. van Sint Annaland, Recent advances on membranes and membrane reactors for hydrogen production, Chemical Engineering Science 92 (2013) 40–66.

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