Microchannel Fuel Processors as a Hydrogen Source for Fuel Cells in

Article pubs.acs.org/EF

Microchannel Fuel Processors as a Hydrogen Source for Fuel Cells in Distributed Energy Supply Systems G. Kolb,* S. Keller, M. O’Connell, S. Pecov, J. Schuerer, B. Spasova, D. Tiemann, and A. Ziogas Institut für Mikrotechnik Mainz GmbH (IMM), Carl-Zeiss-Straße 18-20, D-55129 Mainz, Germany ABSTRACT: Microstructured fuel processing reactors have been designed and tested for the production of hydrogen from fossil and renewable fuels for fuel cell applications. For airborne power generation, a prototype reactor for the partial dehydrogenation of kerosene was developed, which produced more than 100 L/min hydrogen. A system is under development for the integrated reforming of methanol in a high-temperature proton exchange membrane (PEM) fuel cell, which will serve portable applications. To reduce the emissions of trucks during night time, a 5 kW auxiliary power unit (APU) was constructed for power generation during break intervals of the driver. A 250 W liquefied petroleum gas (LPG)-based fuel processor/fuel cell system has been developed for recreational vehicles, which is now commercially available.

1. INTRODUCTION Future sustainable and distributed energy generation will rely in many cases on fuel cell technology. Especially for stationary applications of the smaller scale, for mobile and portable power generation systems, a compact hydrogen supply is required.1 This could be achieved by storage of compressed or liquefied hydrogen. However, the volumetric and gravimetric storage density of hydrogen is much lower compared to liquid energy carriers, such as fossil fuels or renewable biofuels. To extract the hydrogen from these energy carriers, fuel processing is required. A fuel processor is normally composed of the reformer reactor itself and, depending upon the fuel and the type of fuel cell connected to it, devices for removal of the carbon monoxide, which is a main byproduct of reforming. The main requirements for a reformer are full conversion of the fuel and minimization of impurities (ideally 0 ppm) in the reformate. The reforming reaction is heterogeneously catalyzed in most cases, but homogeneous reformers operated at temperatures exceeding 1000 °C2 have been reported in the literature as well as plasma reformers, which apply nonequilibrium plasma to convert the fuel homogeneously.3,4 The catalytic conversion of hydrocarbons or higher alcohols by steam reforming, partial oxidation, or combinations of these usually produces reformate, which has a composition corresponding to the equilibrium of two side reactions of reforming, namely, water-gas shift and methanation. CO + H 2O ↔ CO2 + H 2

becomes obvious from Figure 1b that the catalyst produces reformate, the composition of which is close to the equilibrium of the water-gas shift. In particular, coke precursors, such as ethylene, damage not only the CO cleanup devices downstream of the reformer (see below) but also the fuel cell membrane electrode assembly (MEA) especially when proton exchange membrane (PEM) fuel cells are applied. Other impurities, such as formic acid, damage the fuel cell even more and, therefore, must not exceed a concentration of a few parts per million (ppm). The removal of carbon monoxide, which is mandatory especially when low-temperature PEM fuel cells are connected to the fuel processor, is achieved especially in small and mobile applications by catalytic conversion through water-gas shift and preferential oxidation, while membrane separation or pressure swing adsorption are more suited for systems of larger scale. Microchannel reactors, owing to the improved heat and mass transfer in small channels, provide significant potential for the reduction of the system size, system complexity, and catalyst costs, among others.5,6 Scaled-up microchannel reactors containing thousands of channels create less scale-up issues compared to conventional (fixed bed) systems. They allow for a careful heat management during start-up and normal operation, which increases the service life of the catalyst. At IMM, the development of complete fuel processors for alcohol and hydrocarbon fuels is ongoing, applying microstructured plate heat-exchanger technology. Complete fuel processors are under development using steam reforming, partial oxidation, and partial dehydrogenation reactions on a large variety of fuels, such as methanol,7,8 ethanol,9,10 natural gas, liquefied petroleum gas (LPG),11 gasoline,12 kerosene, and diesel.13 The system size varies between 100 W and about 20 kW thermal power of the hydrogen produced. All development

0 ΔH298 K = − 40.4 kJ/mol

(1)

3H 2 + CO ↔ H 2O + CH4 0 ΔH298 K = − 253.7 kJ/mol

(2)

The latter reaction is obviously the reverse reaction of methane steam reforming. As demonstrated experimentally for methane steam reforming in a microchannel reactor at different temperatures over a rhodium-based catalyst, the methane conversion decreases with an increasing temperature according to the thermodynamic equilibrium (see Figure 1a), while the selectivity toward carbon monoxide increases in parallel. It © 2013 American Chemical Society

Special Issue: Accelerating Fossil Energy Technology Development through Integrated Computation and Experiment Received: December 10, 2012 Revised: February 12, 2013 Published: February 13, 2013 4395

dx.doi.org/10.1021/ef302039x | Energy Fuels 2013, 27, 4395−4402

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Figure 1. (a) Gas composition and methane conversion (XCH4) for methane steam reforming over a rhodium-based catalyst, with S/C = 2 and VHSV = 150 L of H2 h−1 g−1 of catalyst. (b) Selectivity toward carbon monoxide and carbon dioxide. The dotted lines represent the equilibrium of the water-gas shift reaction under these conditions. reaction itself,20 and therefore, the reaction temperature should be limited to 400 °C.23 However, the thermodynamic equilibrium favors the dehydrogenation reaction above 300 °C,19,29 and therefore, a trade-off is required. Dehydrogenation of single hydrocarbon compounds, such as methylcyclohexane, which could be used as a hydrogen carrier for transportation purposes, has been the subject of numerous studies.19,29,30 Platinum supported by alumina was applied as a catalyst in most cases.30 This catalyst support also showed the highest activity and stability followed by silica and activated carbon for the dehydrogenation of cyclohexane at temperatures exceeding 360 °C,24 while palladium is known to be less active for dehydrogenation reactions.31 The product of catalytic dehydrogenation reactions is free of carbon monoxide, but it contains, apart from the hydrogen, small amounts of light hydrocarbons, especially during an initial induction period.32 Wang et al. observed methane and ethane during dehydrogenation of cyclohexane at reaction temperatures exceeding 360 °C.24 Wang et al. performed dehydrogenation of regular Jet A aviation fuel over a sulfated zirconia/platinum catalyst.24 The strong acidity of this catalyst leads to high activity but also faster deactivation by coke formation. This coke could be removed by air exposure at 500 °C.24 2.1. Experimental for Partial Dehydrogenation of Kerosene. The catalyst samples described below were tested as washcoats, which were deposited onto aluminum foams (45 ppi) with a size of 74 × 33 × 2 mm. First the carrier powders (alumina and mixed alumina/silica, titania, and zinc oxide) were co-impregnated with the precursors by incipient wetness impregnation of the metal salt solution to achieve the desired concentration of active species. After about 3 h of drying at room temperature, the impregnated powders were calcined at 450 °C for 6 h. Then, a suspension with the catalyst powder was prepared according to a method that has been described in a previous paper in more detail.33 The foams were then completely filled with the catalyst suspension, dried at room temperature, and calcined at 450 °C. This procedure was repeated for another 2 times. Powder samples of each catalyst were prepared in parallel for analysis purposes. The catalysts were not reduced prior to the activity tests described below. An overview of the catalyst samples, which are reported in the current paper, is provided in Table 1. The specific surface area was determined by nitrogen sorption and calculated by the Brunauer− Emmett−Teller (BET) method. An energy-dispersive X-ray fluorescence (ED-XRF) spectrometer and atomic absorption spectroscopy (AAS) were used for the chemical analysis. In most cases, the amount of active species calculated from the impregnation procedure agreed quite well with the results from chemical analysis. 2.2.1. Experimental Setup for Partial Dehydrogenation of Kerosene. Partial dehydrogenation of kerosene was performed on a variety of Pt-containing catalysts in a continuously operated test rig. A testing reactor heated by heating cartridges took up the foam samples described. It was operated at 10 bar reaction pressure and 350 °C reaction temperature. The pressure of 10 bar was chosen as a result of preliminary experiments not discussed in the scope of this paper to

steps required for building a fuel processor have been addressed, i.e., reactor and system simulation and design, catalyst development, catalyst stability against impurities in the fuel, and setup/testing of complete fuel processors. In addition, the fabrication technology required for the production of the fuel processors has been addressed, namely, catalyst coating by screen printing into the microchannels and laser welding for sealing the reactors.14,15 Below selected examples of this development work will be presented to provide an overview of the scope of activities.

2. PARTIAL DEHYDROGENATION OF KEROSENE Partial dehydrogenation of kerosene (expressed here in a simplified manner as a C11H24 hydrocarbon)

C11H 24 → C12H 22 + H 2

(3)

represents an alternative processing route to compressed or liquefied hydrogen storage and fuel processing, with the potential to achieve reduced system complexity, provided that the dehydrogenation reaction does not considerably change the combustion characteristics of the fuel.16 The reaction is endothermic, and the reaction enthalpy amounts to 126.0 kJ/mol in the case of undecane. Therefore, an external heat supply is common in industrial dehydrogenation reactors,17 while process intensification through wall-coated reactors has been previously proposed for the reaction.18 However, partial dehydrogenation in structured reactors has been the topic of few publications to date,19 and none of them demonstrated the hydrogen production at larger than laboratory scale. Industrial C10−C14 paraffin dehydrogenation is operated at slightly elevated pressure between 2 and 3 bar and in the temperature range between 400 and 500 °C.20 The conversion is typically in the range of 13%, and the selectivity toward monoalkenes amounts to 90% at more than 1 month catalyst lifetime. Industrial dehydrogenation catalysts, which are applied for the dehydrogenation of light alkanes are either chromium oxide or platinum/tin formulations.17,21,22 The platinumbased catalysts require high dispersion of the noble metal, which, in turn, also favors the deactivation of the catalyst by carbon formation.23 However, it is assumed that tin reduces the activity of platinum, while it maintains its long-term stability against coking and sintering22,24 and prevents consecutive dehydrogenation reactions.20 The effects of tin on platinum are not completely clarified.25 Additives, such as potassium, decrease not only the deactivation of Pt/Sn catalysts but also their dehydrogenation activity.26 Resini et al. have summarized the role of alkali promoters on partial dehydrogenation catalysts,27 while the beneficial effect of Mg addition to Pt/Sn catalysts for partial dehydrogenation has been demonstrated by de Graaf et al.28 They reduce the surface acidity, potentially promote hydrogen spillover, reduce coke formation, and maintain platinum accessibility once coke has been formed. The reactions that lead to catalyst deactivation are in many cases more temperature-sensitive than the dehydrogenation 4396


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17. Because the degree of conversion of each individual compound was very low and shifts from one compound to the other may have occurred, it was decided not to analyze the kerosene product. The hydrogen yield was calculated by the following equation:

Table 1. Chemical Composition as Determined by XRF and AAS and Specific Surface Area of the Catalyst Samples Prepared target content of active species (wt %) Pt Pt Pt Pt Pt Pt

a

5 5 5 5 5 5 Mg 2

carrier material

elemental analysis via XRF (wt %)

AI2O3 ZnO TiO2 SiO2 SiO2/AI2O3 AI2O3

Pt: 5.08 ± 0.20 Pt: 6.09 ± 0.30 Pt: 4.72 ± 0.14 Pt: 3.32 ± 0.10 Pt: 5.72 ± 0.17 Pt: 4.93 ± 0.15, Mg: 1.65 ± 0.08a Pt: 4.92 ± 0.15, K: 2.8 ± 3.14a Pt: 4.56 ± 0.14, Na: 1.02 ± 0.05a

Pt 5 K 2

AI2O3

Pt 5 Na 2

AI2O3

specific surface area (m2/g)

yield, Y (%) =

201 20 82 323 266 194

nH ̇ 2 nĊ 12H23

(4)

A 100% yield refers to the case that each mole of kerosene releases 1 mol of hydrogen on average, which is close to the desired situation. The yield of light hydrocarbons was not calculated but rather the molar fraction of methane, C2, and C3 components in the hydrogen. While methane is not crucial for the operation of a PEM fuel cell up to a concentration of 5 vol % and more, C2 and C3 are poisons even at the ppm level.1 Higher hydrocarbons than C3 were not observed in the product. 2.2.2. Results from Catalyst Screening for Partial Dehydrogenation of Kerosene. The effect of the catalyst carrier was investigated by preparing and testing samples based on alumina, 10% silica/90% alumina, pure silica, titania, and zinc oxide as carrier material. As shown in Figure 2, all samples showed an initial induction period, which is known for different catalyst formulations under conditions of partial dehydrogenation of hydrocarbons.34 As reported by Chen et al., this period was even more pronounced in the case of the catalysts that were reduced before they were exposed to the reaction conditions.34 Conversion was highest for the sample containing the titania carrier compared to the alumina base case, despite its lower surface area, followed by the mixed silica/alumina sample. The samples containing silica and zinc oxide showed lower activity. Obviously, the activity toward dehydrogenation did not correlate with the catalyst acidity. However, the stability of all other samples was lower compared to the alumina base case. The formation of C2 hydrocarbons was in the same order of magnitude for the samples containing alumina, mixed alumina/silica, and pure silica. On the other hand, the samples containing titania and zinc oxide, which are known to have the lowest acidity, were less selective toward C2 hydrocarbons (see Table 2). The formation of C3 hydrocarbons was suppressed by all other carrier materials compared to the alumina base case and was lowest for the titania- and silica-containing samples (see Table 2). However, because the catalyst stability was regarded as most crucial, the alumina carrier was selected as the most promising solution and all further samples, which are discussed below, were prepared with alumina carrier. Alkaline and earth alkaline additives were investigated as a next step, namely, sodium, potassium, and magnesium, on a 2 wt % level. The

188 197

Analysis was determined by AAS.

suppress the rate of catalyst deactivation and light hydrocarbon formation. Sulfur-free kerosene (sulfur content of 1 ppm) was fed from a pressurized tank with a thermal mass flow controller to the reactor at a flow rate of 60 g/h. The kerosene feed passed through the foam length axis. Owing to the high thermal conductivity of the aluminum foam, isothermal operation all over the foam sample was observed within the accuracy of the temperature measurement (±2 K). The product mixture was fed into a glass vessel to separate the hydrogen and light hydrocarbon products from the dehydrogenated kerosene product. To determine the flow rate of hydrogen and other gaseous products, nitrogen was added as an internal standard to the product flow downstream of the reactor. The gaseous products, namely, methane, C2, and C3 species apart from the hydrogen, were then monitored by online gas chromatography (Varian CP-4900 Micro-GC). The gas chromatograph could not separate C2 and C3 alkanes from the corresponding alkenes, and therefore, these components were lumped as C2 and C3, respectively. The kerosene was analyzed by simulated distillation (ASTM D2887 standard), which revealed a boiling point distribution between 115 and 287 °C. Qualitative gas chromatography−mass spectrometry (GC− MS) analysis showed the presence of alkanes, cycloalkanes, aromatic, and bicycloalkane compounds, with carbon numbers between 5 and

Figure 2. Hydrogen yield as determined over different samples containing 5 wt % Pt, each supported by different carrier materials, with a reaction temperature of 350 °C, reaction pressure of 10 bar, catalyst mass of 1.0 g, and sulfur-free kerosene flow rate of 60 g/h. 4397


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Table 2. Molar Concentration of C2 and C3 Components in the Hydrogen Produced as Determined over Different Samples Containing 5 wt % Pt, Each Supported by Different Carrier Materialsa target content of active species (wt %) Pt Pt Pt Pt Pt

5 5 5 5 5

carrier material

C2 molar oncentration in H2 (ppm)

C3 molar oncentration in H2 (ppm)

AI2O3 ZnO TiO2 SiO2 SiO2/AI2O3

573 160 313 617 562

16300 234 243 5150 1452

Table 3. Molar Concentration of C3 Components in the Hydrogen Produced as Determined over Different Samples Containing 5 wt % Pt, Each Supported by Aluminaa target content of active species (wt %) Pt Pt Pt Pt

5 5 Mg 2 5K2 5 Na 2

carrier material AI2O3 AI2O3 AI2O3 AI2O3

C2 molar C3 molar concentration in H2 concentration in H2 (ppm) (ppm) 573 700 1374 3365

16300 264 1829 5186

Reaction temperature of 350 °C, reaction pressure of 10 bar, catalyst mass of 1.0 g, and sulfur-free kerosene flow rate of 60 g/h.

a

a

Values shown were determined with a reaction time of 30 min, reaction temperature of 350 °C, reaction pressure of 10 bar, catalyst mass of 1.0 g, and sulfur-free kerosene flow rate of 60 g/h.

catalyst had not been aged for 100 h before, as was the case for the current sample. 2.3. Development of a Hot-Gas-Driven Modular Demonstration Reactor for Partial Dehydrogenation of Kerosene. 2.3.1. Reactor Design. The next step of the development work was the design and fabrication of a demonstration reactor of larger than laboratory scale. The energy required for the endothermic dehydrogenation reaction could originate from the combustion of fuel cell anode off-gas and some additional kerosene in a future technical process. The hot combustion gases could supply the dehydrogenation reactor with energy. The reactor was therefore designed as a plate heat exchanger composed of layers of microstructured heating channels (see the left panel of Figure 5), which were fed with hot gas and of reaction layers and carried catalyst coated onto aluminum foam modules. The kerosene entered the reaction module bottom plate and was distributed therein before it passed through the 6 mm thick foam sample (see the right panel of Figure 5), which had both length and width of 73 mm. The foams were cut by wire electrodischarge machining. A total of 304 g of the Pt/Mg catalyst described above was then coated onto 60 foam segments. The distribution plates and the heating channels of the reactor are shown in the left panel of Figure 6. They were prepared from stainless-steel 316 Ti by wet chemical etching, while the frames, which were required to take up the foam, were fabricated by laser cutting. The modules were stapled and sealed by laser welding. Two reactors composed of 30 modules each were fabricated (see the right panel of Figure 6).

alkaline metal addition decreased the activity of the platinum catalysts significantly, as shown in Figure 3, while the magnesium-containing sample still showed similar activity compared to the base case. While the addition of sodium and potassium increased the content of C2 hydrocarbons in the hydrogen (see Table 3), it remained in the same range when magnesium was added. While the stability of the magnesium-containing sample was slightly impaired in comparison to the base case, the content of C3 hydrocarbons could be decreased by far more than 1 order of magnitude, as shown in Table 3. As described above, most of the catalysts under investigation showed rapid deactivation during the first few hours of testing, despite the absence of sulfur in the feed. As a result of further extensive optimization work, a catalyst containing 10 wt % platinum and 1 wt % magnesium could be identified, which showed stable performance for about 30 h at 15% hydrogen yield and 1.4 L of H2 h−1 g−1 of catalyst volume hourly space velocity (VHSV) after some initial deactivation during the first 100 h of operation, as shown in Figure 4. The fluctuations of the hydrogen yield originated from the emptying of the separation bottle, which was required from time to time. The catalyst showed only moderate formation of C2 and C3 components, well below 500 ppm in the hydrogen product after an induction period of less than 10 h. The specific hydrogen production is regarded as comparable to data available from the literature; e.g., Lazaro et al. achieved a hydrogen production rate of 1.7 L of H2 h−1 g−1 of catalyst over their Pt/CNF catalyst during batch operation.23 However, this

Figure 3. Hydrogen yield as determined over different samples containing 5 wt % of Pt each supported by alumina, with a reaction temperature of 350 °C, reaction pressure of 10 bar, catalyst mass of 1.0 g and sulfur-free kerosene flow rate of 60 g/h. 4398


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Figure 4. Kerosene conversion and concentration of C2 and C3 components in the hydrogen produced as determined during a medium-term stability test over catalyst containing 10 wt % platinum and 1 wt % magnesium, with a reaction temperature of 350 °C, reaction pressure of 15 bar, catalyst mass of 1.0 g, and sulfur-free kerosene flow rate of 60 g/h.

Figure 5. Modules of the demonstration reactor. (Left) Details of the reaction module. (Right) Heating channels on the reverse side of the bottom plate of the reaction module.

Figure 6. (Left) Parts of one reactor module. (Right) Demonstration reactor fabricated from 30 modules, as shown in Figure 5. The test rig described above was adapted to the higher flow rates required for the demonstration reactor. A high-pressure gear pump from Barmag−Saurer was used for supplying the pressurized kerosene feed, which was preheated in stainless-steel tubes inside an oil thermostat to a temperature of 250 °C. The product flow was cooled by water-cooled tube-in-tube heat exchangers downstream of the reactor, and the dehydrogenated kerosene was separated from the

hydrogen and light hydrocarbons in a polymer vessel, while the flow rate of the nitrogen, which was used as an internal standard, was increased to compensate for the much higher hydrogen flow rates produced. Self-developed gas heaters supplied with electric energy were used to preheat up to 600 L/min air as a surrogate of the hot combustion gases to heat the reactors. The heating gas temperature and flow rate 4399


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were adjusted, as deemed necessary, to provide the heat for the endothermic dehydrogenation reaction and to compensate for the heat losses of the reactors to the environment. 2.3.2. Results from Demonstration Reactor Testing. Up to 16 L/h kerosene was fed into the reactors, and the maximum flow rate of hydrogen produced exceeded 110 L/h at a reaction temperature of 375 °C. This corresponded to a 4.7% hydrogen yield and a VHSV of only 0.36 L of H2 h−1 g−1 of catalyst. Therefore, the hydrogen production rate was much lower compared to the results from catalyst screening, which was attributed to maldistribution of the kerosene over the foam elements. However, the hydrogen flow rate would have been sufficient to power a 100 W fuel cell. It is to the authors’ knowledge the first partial dehydrogenation reactor developed for fuel cell applications of that size presented in the literature.

4. DIESEL FUEL PROCESSOR/FUEL CELL AUXILIARY POWER UNIT (APU) The choice of fuel converted to hydrogen mainly depends upon the application and existing infrastructure. While trucks have diesel fuel readily available, LPG is the right choice for caravans.35 Idling of truck engines was common practice in the U.S., where drivers spend considerable time in the vehicles overnight. By idling, 1.2 billion gallons of diesel were consumed in the U.S. in 2010. However, owing to the low efficiency of idling (stated to be in the range of 10%) and related emissions, this is prohibited in an increasing number of U.S. federal states. Alternatives, such as internal combustion APUs, show low efficiency in the range of 20% as well. Fuel cell technology promises benefits concerning efficiency and noise emissions.5 Diesel-fueled APU systems based on fuel cell technology show great potential in terms of reducing detrimental effects on human health and the reduction of the environmental impact, while their huge market potential has been forecasted36 and the payback period was estimated to exceed no more than 2 years. In the scope of the European project HYTRAN, a diesel APU was developed that was based on microchannel plate heatexchanger technology.35 The complete system is shown in Figure 8.

3. INTERNAL REFORMING HIGH-TEMPERATURE PEM FUEL CELL SYSTEM In the scope of a project IRAFC, IMM has and is currently developing the balance-of-plant components for a hightemperature PEM fuel cell stack powered with fuel from internal methanol reforming. The net power output of the system equals 100 W. The basic idea of the project is to convert the methanol by steam reforming inside the fuel cell stack in dedicated foam elements operated at the temperature of the fuel cell. However, such a system, despite drastic simplification measures that have been achieved through the combination of catalytic fuel reforming and electrocatalytic hydrogen conversion, still requires balance-of-plant components. A computer-aided design (CAD) model of the system is shown in Figure 7. The required components are (a) a monolithic

5. LPG FUEL PROCESSOR/FUEL CELL SYSTEM VEGA In close collaboration between IMM and the German company Truma, a fuel cell system was developed, which works with high-temperature PEM technology and a LPG-based microchannel fuel processor, operated under conditions of steam reforming. The system, which is designed for recreational vehicles and decentralized industrial applications, was already tested in field trials in 2009. A total of 200 systems were fabricated and tested until August 2012, when the VeGA system was introduced on the market. Figure 9 shows the current design of the system, which is a battery charger with a net power output of 250 W. It contains fuel cell, fuel processor, balance-of-plant components, fully automated control, power conversion for battery charging, and water recovery to close the water balance. This is crucial in the case of steam reforming, because the water consumption is high. The water production of the fuel cell more than counterbalances the water demand under normal operating conditions. Results from single reformer and system testing have been published elsewhere.11 6. CONCLUSION Fuel processing has significant potential for future power generation in mobile airborne, ground, and naval applications on the short, medium, and when sustainable fuels are applied, also on the longer term. Microstructured and microchannel reactors allow for a compact and integrated design of the components, which is crucial for mobile applications and help to reduce catalyst (noble metal) costs. The practical applicability of an integrated, hot-gas-driven reactor for the partial dehydrogenation of kerosene as a hydrogen source for fuel cell applications could be demonstrated. However, the activity and stability of the catalyst formulation applied require further improvement as well as selectivity toward light hydrocarbons. The application of aluminum foams as catalyst carriers allowed for the maximization of catalyst load per reactor volume, which is crucial, owing to the relatively low catalyst activity achieved thus far. However, the design of the hot-gas-driven reactor was proven

Figure 7. CAD model of the system periphery of an internal methanol reforming PEM fuel cell system developed in the scope of the project IRAFC.

catalytic burner for preheating the fuel cell (indicated as a black box in the background of Figure 7) to operate the temperature [this burner is supplied with methanol during startup by a small evaporator heated by electricity (EVA S/U in Figure 7)], (b) heat exchangers for preheating the cathode air feed by cathode off-gas (not shown in Figure 7) and for cooling the fuel cell through cooling oil (HX-02 in Figure 7), and (c) a combined catalytic burner/evaporator for evaporation of the methanol/ water feed by fuel cell anode off-gas combustion (EVA/AFB in Figure 7). All of these components have been fabricated at IMM. The system is currently under construction. 4400


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Figure 8. (Left) CAD model of the 5 kW diesel fuel processor/fuel cell APU developed in the European project HYTRAN. (Right) Finished APU.

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(1) Kolb, G. Fuel Processing for Fuel Cells, 1st ed.; Wiley-VCH: Weinheim, Germany, 2008. (2) Brueggemann, P.; Seifert, P.; Meyer, B.; Mueller-Hagedorn, M. Chem. Prod. Process Model. 2010, DOI: 10.2202/1934-2659.1444. (3) Bromberg, L.; Cohn, D. R.; Rabinovich, A.; Alexeev, N.; Samokhin, A.; Ramprasad, R.; Tamhankar, S. Int. J. Hydrogen Energy 2000, 25, 1157−1161. (4) Czernichowski, A.; Czernichowski, J.; Czernichowski, P.; Wesolowska, K. Reformate gas production from various liquid fuels. In International Symposium and Workshop on Fuel Cells and Hydrogen for Aerospace and Maritime Applications; Hamburg, DE, Sept 16−17, 2004; pp 121−130. (5) Specchia, S.; Specchia, V. Ind. Eng. Chem. Res. 2010, 49, 6803− 6809. (6) Kolb, G. Chem. Eng. Process. 2013, DOI: 10.016/j.cep.2012.015. (7) Kolb, G.; Schelhaas, K.-P.; Wichert, M.; Burfeind, J.; Hesske, C.; Bandlamudi, G. Chem. Eng. Technol. 2009, 32 (11), 1739−1747. (8) Kolb, G.; Keller, S.; Tiemann, D.; Schelhaas, K. P.; Schuerer, J.; Wiborg, O. Chem. Eng. J. 2012, 207−208, 388−402. (9) Men, Y.; Kolb, G.; Zapf, R.; Hessel, V.; Löwe, H. Process Saf. Environ. Prot. 2007, 85 (B5), 413−418. (10) Kolb, G.; Men, Y.; Schelhaas, K. P.; Tiemann, D.; Zapf, R.; Wilhelm, J. Ind. Eng. Chem. Res. 2011, 50, 2554−2561. (11) Wichert, M.; Men, Y.; O’Connell, M.; Tiemann, D.; Zapf, R.; Kolb, G.; Butschek, S.; Frank, R.; Schiegl, A. Int. J. Hydrogen Energy 2011, 36, 3496−3504. (12) Kolb, G.; Baier, T.; Schürer, J.; Tiemann, D.; Ziogas, A.; Ehwald, H.; Alphonse, P. Chem. Eng. J. 2008, 137 (1−3), 653−663. (13) O’Connell, M.; Kolb, G.; Schelhaas, K. P.; Schürer, J.; Tiemann, D.; Ziogas, A.; Hessel, V. Int. J. Hydrogen Energy 2009, 34, 6290−6303. (14) O’Connell, M.; Kolb, G.; Schelhaas, K.-P.; Wichert, M.; Tiemann, D.; Pennemann, H.; Zapf, R. Chem. Eng. Res. Des. 2012, 90, 11−18. (15) Klotzbücher, T.; Braune, T.; Arlt, A. G.; Kolb, G. Laser Tech. J. 2010, 5, 28−31. (16) Jaenker, P.; Felix, N.; Wolff, C. Device for the generation of hydrogen gas by dehydrogenation of hydrocarbon fuels. WO2007/ 033641, 2007. (17) Virnovskaia, A.; Jorgensen, S.; Hafizovic, J.; Prytz, O.; Kleimenov, E.; Hävecker, M.; Bluhm, H.; Knop-Gericke, A.; Schlögl, R.; Olsbye, U. Surf. Sci. 2007, 601, 30−43. (18) Tschudin, S.; Shido, T.; Prins, R.; Wokaun, A. J. Catal. 1999, 181, 113. (19) Kariya, N.; Fukuoka, A.; Utagawa, T.; Sakuramoto, M.; Goto, Y.; Ichikawa, M. Appl. Catal., A 2003, 247, 247−259. (20) Saeedizad, M.; Sahebdelfar, S.; Mansourpour, Z. Chem. Eng. J. 2009, 154, 76−81.

Figure 9. Fuel cell/fuel processor system VeGA developed by Truma. The system works fully automated with touchscreen control.

to be applicable but requires further improvement concerning its isothermity and flow distribution. Other applications are closer to commercialization, such as diesel fuel cell APUs for trucks, especially for the U.S. market. With the market introduction of the LPG fuel cell system, VeGA microchannel reactors have become a mass product readily available for end-users.

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REFERENCES

AUTHOR INFORMATION

Corresponding Author

*Telephone: 0049-6131-990341. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Parts of the work presented here were funded by Airbus Deutschland GmbH and the German Ministry of Economics and Technology in the scope of the projects ELFA-BREZEN (as part of the national aerospace research program) and LPGAPU. Other work was funded by the European Joint Technology Initiative (JTI) for Hydrogen and Fuel Cells in the scope of the project IRAFC and by the European Commission in the scope of the European project HYTRAN. 4401


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