ARTICLE pubs.acs.org/IECR
Development Work on a Microstructured 50 kW Ethanol Fuel Processor for a Small-Scale Stationary Hydrogen Supply System G. Kolb,*,† Y. Men,† K.-P. Schelhaas,† D. Tiemann,† R. Zapf,† and J. Wilhelm‡ † ‡
Institut f€ur Mikrotechnik Mainz GmbH (IMM), Carl-Zeiss-Strasse 18-20, 55129 Mainz, Germany Rosetti Marino SPA, 48122 Ravenna, Italy ABSTRACT: Development work for a compact hydrogen supply system with a thermal output of 65 kW applying ethanol as fuel was started by IMM in collaboration with Rosetti Marino, an Italian plant engineering company. The system concept is comprised of an ethanol steam reformer operated at 10 bar, water-gas shift for the reduction of carbon monoxide in the reformate, a pressure swing adsorption for separating the hydrogen out of the reformate, and numerous heat exchangers and coupled evaporators/ catalytic burners as balance-of-plant. The work included all aspects of the fuel processor, namely, development of a Rh/Co catalyst for ethanol steam reforming, the verification of its long-term durability in a 1000 h test, static and dynamic simulation work applying ASPEN plus and ASPEN Dynamics, setup of a control strategy, and finally the sizing and design of the reactors and the development of a full 3D-CAD model of the fuel processor.
1. INTRODUCTION For future sustainable energy generation, biofuels such as ethanol are prominent liquid energy carriers for transportation purposes owing to their high volumetric power density. The production of hydrogen from the steam reforming of alcohols could favor the use of hydrogen as an alternative fuel, removing the difficulty of storage and distribution.1 Alcohol fuels such as methanol and ethanol are sustainable and require lower operating temperature of the reformer compared to liquid hydrocarbon fuels. From an environmental point of view, the use of ethanol is preferred because ethanol could be considered as a renewable raw material which can be easily obtained from biomass. In addition, the bioethanol system has the significant advantage of being nearly CO 2 neutral since carbon dioxide produced is consumed for biomass growth, thus offering a nearly closed carbon loop. 2 The decentralized generation of hydrogen from ethanol in a fuel processor3 may be regarded as a viable option for future filling stations of vehicles running either on fuel cell technology or on internal hydrogen combustion. The utilization of the catalytic wall reactor concept 4,5 for the ethanol steam reforming reaction has been proposed by Liguras et al. 6 2. CATALYST DEVELOPMENT Ethanol steam reforming is an endothermic reaction C2 H5 OH þ H2 O f 2CO þ 4H2
0 ΔH298 ¼ þ 256 kJ=mol ð1Þ
It has the disadvantage of numerous side reactions to ethylene, acetaldehyde, and methane by the unselective decomposition of the ethanol feed.3 Apart from these reactions, water-gas shift and methanation of carbon monoxide will take place in the reformer reactor itself according to the equilibrium of these reactions at the reformer temperature. r 2010 American Chemical Society
Although much work has been carried out on the catalysis of ethanol steam reforming, previous studies were mainly dealing with catalyst development for conventional fixed-bed reactors. Nickel7,8 and cobalt9-11 are frequently used in fixed bed reactors. However, these catalysts suffer from relatively low activity which is a minor issue as long as fixed catalyst beds are applied because the production costs of the reactor are lower compared to catalytic wall reactors. In monolithic reactors or catalytic wall reactors, the amount of catalyst that can be coated onto the reactor walls is limited.3 Therefore, development work dealing with monolithic reactors focuses in most cases on noble metal catalysts such as rhodium.12,13 The catalytic activity of nickel catalysts could also be improved by switching to partial oxidation of ethanol.14 A more detailed overview of catalysts for ethanol steam reforming has been provided by Vaidya et al.15 Ethanol steam reforming in microchannels has previously been studied over nickel, cobalt, and Rh/Ni catalysts,16 Rh/CeO2,17 and Ir/ CeO2 catalysts.18 Only rhodium-containing catalyst formulations showed medium term stability. The catalyst development work of the current paper was focused on the development of a long-term stable catalyst formulation. The catalyst described here were tested as washcoatings in microchannels applying a manual procedure19 and an experimental apparatus described in detail elsewhere.20 The small testing reactors applied had a sandwich design and were externally heated by heating cartridges as described in the same paper.20 Analysis of the product gases was performed by online gas chromatography. The ethanol and ethanol/water mixtures were evaporated in microstructured evaporators developed Special Issue: IMCCRE 2010 Received: March 12, 2010 Accepted: June 2, 2010 Revised: June 1, 2010 Published: July 7, 2010 2554
dx.doi.org/10.1021/ie100602w | Ind. Eng. Chem. Res. 2011, 50, 2554–2561
Industrial & Engineering Chemistry Research
ARTICLE
Table 1. Ethanol Conversion and Selectivity over Bimetallic Catalystsa temperature [°C] 400
500
catalysts
a
S(CH4) [%]
S(CO2) [%]
S(CO) [%]
S(C2H4O) [%]
S(C2H4) [%]
S(H2) [%]
Rh/Al2O3
59.3
5.3
0.6
14.9
20.1
59.1
31.5
Rh/Ni/Al2O3
90.0
33.4
8.2
35.3
22.4
0.8
47.4
Rh/Co/Al2O3
100.0
40.4
36.5
4.5
0
55.4
Rh/Al2O3
100.0
1.7
5.7
24.0
3.5
65.1
48.5
93.5
23.2
33.7
16.6
8.1
0
63.4
Rh/Al2O3
100.0
39.5
51.5
8.9
0
0
65
Rh/Al2O3
100.0
26.5
48.3
25.1
0
0.1
78.0
Rh/Ni/Al2O3 Rh/Co/Al2O3
99.6 100.0
23.7 20.5
50.8 57.0
25.5 22.5
0 0
0 0
80.8 84
Rh/Ni/Al2O3 600
conv. [%]
18,6
Experimental conditions: steam/carbon = 2:1; pressure, 1 bar; VHSV 90 L/(h gcat).
Figure 1. Durability test of ethanol steam reforming at 650 °C over Rh/ Co/Al2O3 catalyst. Shown is the ethanol conversion and the species concentration in the off-gas: pressure, 1 bar; temperature, 650 °C; S/C ratio, 2.0; VHSV 90 L/(h gcat).
in-house,21 which were supplied with energy by heating cartridges. A catalyst containing 5 wt % Rh, 10 wt % Ni and 15 wt % CeO2 had been identified as the best suited catalyst for ethanol steam reforming in previous development work by the authors of the current paper especially in the temperature range below 600 °C.16 Further development work revealed that bimetallic Rh/Co catalyst containing 5 wt % Rh and 10 wt % Co catalyst exhibits better catalyst performance than both monometallic Rh catalysts and Rh/Ni/CeO2 catalysts already at relatively low temperatures (see Table 1). The catalytic behavior of Rh is greatly enhanced by promotion with Co (or Ni, as formerly reported) likely through the formation of bimetallic particles. As for the promoting role of Co, the high selectivity of acetaldehyde was observed over Co-based catalyst, which is quite independent of the type of support, indicating the dehydrogenation nature of Co metal. Thus, in accordance to this reaction scheme, the reaction was suggested to proceed first through the formation of an acetaldehyde intermediate, which subsequently decomposed to produce a mixture of CO and CH4 or reformed with H2O producing CO, CO2, and H2. The role of Rh is mainly to cleave the C-C and C-H bonds of ethanol similar to how it does in the case of hydrocarbon reforming to produce H2 and COx, while Ni and Co addition help to convert ethanol into acetaldehyde by dehydrogenation under the reaction conditions employed as indicated by their high selectivity toward acetaldehyde which has been reported previously.16
Figure 2. 1000 h durability test of the rhodium-cobalt catalyst developed in-house (5 wt % Rh, 10 wt % cobalt) for ethanol steam reforming. Shown is the ethanol conversion and the species concentration in the off-gas: pressure, 10 bar; temperature, 700 °C; S/C ratio, 3.0; VHSV 34 L/(h gcat).
The stability of Rh/Co bimetallic catalyst with time on stream was examined at 650 °C for 130 h. The Rh/Co bimetallic catalyst showed some slight degradation of hydrogen content in the reformate (see Figure 1). Full conversion was achieved for the entire run. The CO content in the produced H2-rich reformate decreased slightly, while the methane content increased, which impaired the amount of hydrogen formed. The content of CO2 remained unchanged throughout the run. No other byproduct was detected. The catalyst formulation was then further improved. and finally the durability of the Rh/Co-based ethanol steam reforming catalyst, which had been developed in-house, was proven by a 1000 h stability test (see Figure 2) at IMM. This test was performed under realistic conditions of the process envisaged here, namely, at a pressure of 10 bar (see section 3 below), a reaction temperature of 700 °C, and, however, a lower Volume Hourly Space Velocity (VHSV) to maintain the durability of the catalyst. To supply the endothermic steam reforming reaction with energy, ethanol combustion was foreseen in the process described below in section 3. Therefore a similar 1000 h durability test was performed for the platinum-based ethanol combustion catalyst (not shown here). The pressure for ethanol combustion in the process described below is ambient, and therefore tests were carried out at 1 bar. A low reaction temperature of 350 °C was chosen to verify especially the low-temperature activity of the catalyst. The λ-value of the feed mixture was 1.45, while a high 2555
dx.doi.org/10.1021/ie100602w |Ind. Eng. Chem. Res. 2011, 50, 2554–2561
Industrial & Engineering Chemistry Research
ARTICLE
Figure 3. Flow scheme of the fuel processor.
VHSV of 461 L/(h gcat) was set. The catalyst showed stable performance throughout the entire test.
3. DESIGN OF THE ETHANOL FUEL PROCESSOR The only complete ethanol fuel processor, which has not only been modeled but also built, tested, and connected to a fuel cell, is the system described by Aicher et al.22 However, the thermal output of that system amounted only to less than 1 kW, and the reaction conditions chosen were rather autothermal reforming and not steam reforming as in the current case. While thermodynamic analysis23 and efficiency calculations24 of the ethanol steam reforming reaction have been performed, only a few studies deal with process design or static modeling of an ethanol reformer fixed bed reactor.25,26 One outcome of such simulation studies was that a low catalyst utilization below 12% is to be expected in fixed bed reactors.25 The current paper presents the design and detailed static and dynamic modeling of an entire ethanol fuel processor in the thermal power range of 65 kW. At IMM, microstructured plate heat-exchanger technology has been applied for the development of compact fuel processors for portable27 and mobile28-30 applications in the power range of up to 13 kWth. Microchannels bear the advantage of a much better utilization of the catalysts because they are coated onto the walls of the microchannels as described in Section 2. Another advantage is the heat integration and combination of exothermic and endothermic reactions together in one reactor as will be described below. It is obvious that increasing the S/C ratio to values much higher than the stoichiometric value of 1.5 in the reformer feed impairs the system efficiency31 because not all the steam which is generated by the evaporator will be removed from the product. Consequently, energy losses are generated, which get even more severe when the energy losses to the environment are considered. A thermodynamic analysis by Francesconi et al. revealed a fuel processor efficiency in the range of 80% for S/C ratios between 2
and 2.5.24 Another thermodynamic analysis reported an efficiency loss of the fuel processor from 65.5% to 55% when increasing the S/C ratio from 3.2 to 6.5.32 To maintain the stability of the catalyst, the S/C ratio of the process simulations was set to a minimum value of 3. 3.1. System Setup and Implementation of the Static Model. The modeling was performed in a two-stage process. As a first step, a static model was implemented in ASPEN Plus to determine a system design with maximum hydrogen yield and minimum heat losses by hot off-gases. The design point was also fixed in that model. No kinetic expressions were implemented into the static or into the dynamic model described below because kinetics for ethanol steam reforming are rarely available and not at all for the current catalyst. However, the dynamic interplay of the different components in the system is less affected by the kinetics of the reaction (especially because the reactors are always operated in a manner that full conversion is achieved in the reformer and equilibrium conversion in the water-gas shift reactor) but rather by the thermal mass of the components and the effects of changing flow rates originating from the control values of the controllers. Because a pressure swing adsorption (PSA) system was foreseen as the gas purification unit downstream, the water-gas shift reactor, a system pressure of 10 bar was required. Elevated pressure is in principle not favorable for the steam reforming reaction because it produces substantial amounts of methane in the reformer owing to the equilibrium of the methanation reaction. Therefore, a relatively high temperature had to be chosen for the reformer reactor to suppress the methane formation to an acceptable minimum. However, the remaining methane in the reformate was not foreseen to be separated from the hydrogen in the PSA but rather to be utilized as fuel in the current system because the intention was to utilize the hydrogen as an additive for natural gas vehicles. Figure 3 shows a system scheme of the fuel processor along with temperature and mass flows of the process streams. The fuel 2556
dx.doi.org/10.1021/ie100602w |Ind. Eng. Chem. Res. 2011, 50, 2554–2561
Industrial & Engineering Chemistry Research
ARTICLE
Calculation of the fuel processor efficiency according to eq 3 revealed a value of 90.6%. ηfuelprocessor ¼
Figure 4. Gas composition of reformer feed and reformate throughout the process. Process streams correspond to Figure 3.
processor is comprised of the combined reformer/burner reactor (RB), the air-cooled water-gas shift reactor (WGS), a combined evaporator/catalytic burner (EVB), and eight heat exchangers. Ethanol/water feed (Feed1) is preheated in the feed preheater (FP) and evaporated in a two-stage evaporation process (EV and EVB). The first stage evaporation is performed in the hot gas driven evaporator (EV), which is fed with the combustion off-gas stemming from the RB, which was already precooled in the air preheater 2 (AP2). The second stage and complete evaporation are performed in the combined evaporator/catalytic burner (EVB), which is fed with an ethanol/air mixture (Mixcomb1). The air of Mixcomb1 is preheated in the air preheater AP1, which is supplied with energy from the EVB. The ethanol required to feed the EVB with energy by combustion is evaporated and preheated in the fuel evaporator (FEV), which is supplied with energy from the hot cooling gas of the WGS reactor (Cool2). The completely evaporated ethanol/water feed (Feed2) is further superheated in the process cooler (PC1) before it enters the RB. The RB is supplied with energy by combustion of the PSA off-gas (Off-gas), which is preheated in the off-gas preheater (OP) and mixed with additional ethanol and air (Mix1). Most of this air is preheated in the AP2. The feasibility of a combined steam reforming and hydrogen combustion as foreseen in the RB has been practically demonstrated for diesel steam reforming and fuel cell anode off-gas combustion in the power range of 13 kWth in previous work performed at IMM.30 The hot reformate leaving RB supplies PC1 with energy. To achieve a direct control of the critical inlet temperature of the WGS reactor, a second process cooler PC2 is switched downstream, which is cooled by a small air flow Air21. The WGS reactor is operated as a single-stage reactor with internal counter-current cooling, which moves the thermodynamic equilibrium of the steam reforming reaction in a favorable direction.33-35 This principle has been successfully demonstrated in the kilowatt range in previous work at IMM.29 The product of the WGS reactor is further cooled in the OP and water removed out of the stream in a condenser. Figure 4 shows the composition of the process streams of the reformer feed and product throughout the process. The reformer feed flow was calculated to 31 kg/h, while the purified product gas flow amounted to 2.6 kg/h or 18.7 N m3/h with a composition of 92% hydrogen and 8% methane. The reformer efficiency was calculated to be 80.4% according to eq 2 ηreformer ¼
LHVðH2 ÞnH2 LHVðfuelÞnfuel
ð2Þ
LHVðH2 ÞnH2 þ LHVðCOÞnCO LHVðfuelÞnfuel
ð3Þ
The size of the reactors was fixed applying the VHSV required to achieve (a) full conversion of the ethanol, (b) equilibrium conversion for water-gas shift and methanation for the reformer, and (c) equilibrium conversion for water-gas shift. The burners were also designed to achieve full conversion of their feed, taking into consideration the equilibrium of reactions taking place. All these VHSV values had been determined experimentally. The gas composition of the reformate determined from the thermodynamic equilibrium calculations of water-gas shift and methanation reactions at the RB operating temperature of 700 °C agreed well with the experimental tests as can be seen comparing Figures 2 and 4. The equilibria were determined by the RGibbs model, which determines the minimum of the Gibbs free energy. In the case of the water-gas shift reaction in the WGS reactor, it was assumed that methane is inert in the reactor, which is a good approximation for the noble metal based catalyst foreseen in this case. The size of the reactors and heat exchangers, their pressure drop, and heat transfer characteristics were calculated utilizing the extensive experience of IMM in setting up such microstructured plate heat exchangers. 3.2. Dynamic Model and Control Strategy for Start-Up and Normal Operation. The stationary model then served as the basis for the dynamic model as referred to in the arrangement of the reactors and heat exchangers and also as the basis for the physical properties of the process streams. The dynamic model was implemented into ASPEN Dynamics to develop a control strategy, which enables the stable operation of the plant in the case of changing operating parameters such as load changes, while the product quality (hydrogen content) is only slightly affected. Heat exchangers and evaporators were implemented by the ASPEN model HeatX. Figure 5 shows the ASPEN flow scheme of the fuel processor. Ten temperature controllers, four mass flow controllers, and one pressure controller are required to control the system, which is summarized in Table 2. The temperature of the stream Feed2, which leaves EVB, is controlled by the ethanol flow fed to the integrated burner of EVB. The ratio of air Air1 to the ethanol stream Feed1 is also controlled to maintain the temperature of the evaporator/burner reactor. The temperature in the RB reactor is controlled by the portion of cold air fed to the air preheater AP2. Because the amount of PSA off-gas fed back to the reformer/burner reactor RB may vary, the total air flow to RB needs to be controlled additionally. The temperature profile of the WGS reactor is adjusted by two means: control of the reformate (Prod3) inlet temperature by cooling it in PC2 by Air21. Alternatively, a water injection could be used in this case, which would shift the equilibrium of the water-gas shift reaction further in a favorable direction but also increase the load to the condenser downstream from the WGS reactor. The WGS product outlet temperature is controlled by the counter-current cooling air flow Cool1. Furthermore, three controllers will be required to control the preheating of EVB, RB, and WGS either by electrical heating or by hydrogen combustion as described below. Apart from these controllers, a further 13 temperatures must be monitored and need to be connected to alarm indicators. Two levels for alarms are provided for, i.e., high and low alarms. 2557
dx.doi.org/10.1021/ie100602w |Ind. Eng. Chem. Res. 2011, 50, 2554–2561
Industrial & Engineering Chemistry Research
ARTICLE
Figure 5. ASPEN flow scheme of the fuel processor with control loops implemented.
Table 2. Overview of the Controllers for the Fuel Processor controller no. controlled variable stream or device manipulated variable
stream
actuator
description
1
temperature
Feed2
mass flow rate
FuelEt1
pump
temperature control of feed evaporation (EVB)
2/1
ratio mass flow rate
Air1/Hydr1
mass flow rate
Air1
valve
air/hydrogen ratio control for evaporator-burner
2/2
ratio mass flow rate
Air1/FuelEt1
mass flow rate
Air1
valve
air/ethanol ratio control for evaporator-burner (EVB)
(EVB) during startup 3/1
temperature
Prod1
ratio mass flow rate Air12/Air 13 valve
temperature control of steam reforming (RB), split range 1
3/2
temperature
Prod1
mass flow rate
FuelEt11
pump
temperature control of steam reforming (RB), split range 2
4/1
ratio mass flow rate
Air11/Hydr11
mass flow rate
Air11
valve
air/hydrogen ratio control for reformer-burner (RB)
4/2
ratio mass flow rate
Air11/Offgas
mass flow rate
Air11
valve
during startup air/offgas ratio control for reformer-burner (RB)
5
temperature
Prod3
mass flow rate
Air21
valve
temperature control of product gas (inlet WGS)
6
temperature
Prod4
mass flow rate
Cool1
valve
temperature control of product gas (outlet WGS)
7
pressure
Prod5
mass flow rate
Prod5
valve
operating pressure control
8/1
temperature
EVB
mass flow rate
Hydr1
valve
temperature control of evaporator-burner (EVB)
8/2
temperature
EVB
electr. current
-
9/1 9/2
temperature temperature
RB RB
mass flow rate electr. current
10
temperature
WGS
electr. current
during start-up heater temperature control of evaporator-burner (EVB) during start-up
Hydr11 -
The fuel processor could be preheated either by hydrogen combustion from a tank or by electrical heating. After inerting, the evaporator/burner EVB, heat-exchanger AP1, evaporator EVB, heat-exchanger AP2, and the reformer/burner RB are heated by combustion of hydrogen with Air 11. As soon as the temperatures are reached, the pressure is increased, and the heatexchangers PC1, PC2, and OP are heated along with the water
valve temperature control of reformer-burner (RB) during start-up heater temperature control of reformer-burner (RB) during start-up
-
heater temperature control of water-gas shift (WGS) during start-up
-gas shift reactor WGS. Then the WGS cooling Cool1 is switched on, and the remaining heat-exchangers FEV and FP are heated. In the second phase, the hydrogen flows are switched off, and the ethanol is dosed into the system, while the controllers for normal operation are switched on. The dynamic behavior of the process in the case of stepwise load changes shall now be discussed. Figure 6 shows the mass 2558
dx.doi.org/10.1021/ie100602w |Ind. Eng. Chem. Res. 2011, 50, 2554–2561
Industrial & Engineering Chemistry Research
Figure 6. Step changes of the mass flow of Feed1 as set for dynamic modeling.
ARTICLE
Figure 9. Volume fraction of hydrogen contained in the process streams during the step changes.
Figure 10. Temperature of the product stream Prod1, controller output, and constant controller set point as determined for the step changes. Figure 7. Temperatures of the process streams as determined for the step changes.
Figure 11. Temperatures of the process streams entering and leaving air preheater AP2 as determined for the step changes. Figure 8. Mass flow rate of hydrogen contained in the process streams as determined for the step changes.
flow of the stream Feed1. From the design point load changes of (20% were set. The time between the changes amounts to 1 h. 2559
dx.doi.org/10.1021/ie100602w |Ind. Eng. Chem. Res. 2011, 50, 2554–2561
Industrial & Engineering Chemistry Research
ARTICLE
Figure 12. CAD model of the future 50 kW fuel processor (left) and side view of components (right) with tubing connections.
Figure 13. CAD model of the future 50 kW fuel processor and PSA unit.
This time is required until a new stable operation is achieved. The time response of the system is determined by two factors: 1. Small time constants for the species flows (instantaneous in the model). 2. Large time constants by storing and releasing heat into and from the components, respectively. Figure 7 shows the temperature changes of the process streams (the same streams as shown in Figure 4) during the load changes. All temperatures do not change significantly, and a maximum deviation of (10 K is observed. The mass flow rates of hydrogen in these process streams are depicted in Figure 8. The new operating points are achieved practically without any delay, while the volume fraction of hydrogen in the process streams is only affected to a small degree (see Figure 9). The maximum deviation of the hydrogen volume fraction amounts to 1% absolute. The controller output and the temperature of the reformer reactor are shown in Figure 10. The first downward step of the reformer feed results immediately in a reduction of the PSA offgas flow fed to the afterburner, which is part of the reformer. Air11 is then automatically reduced by a proportional set point reduction. These measures create in total a temperature drop of the reactor because the portion of Air11, which is dosed through the AP2 (stream Air12), is too small. The controller reacts on this maldistribution and increases the flow of Air12. The reactor temperature is then quickly recovered, and only a slight overshoot of the reactor temperature takes place. The second upward
change creates a different behavior than the first upward change shown in the very left part of Figure 10. This time, the air flow to AP2 is once more too small, and again the temperature drops are similar to the downward change. This, however, happens only because the system has not yet reached the steady state after the former downward step change and hence is in a different state when the upward change happens. A similar explanation can be provided for the much smoother second downward step shown in the right part of Figure 10. The temperatures of the streams, which enter and leave AP2, are shown in Figure 11. While the Air12 stream is always on ambient temperature, the outlet temperature of this stream (Air14) is equivalent to the temperature of the reformer-burner off-gas stream (Flue11) because the latter stream is much larger. The temperature changes of the Flue12 stream, which are as high as (50 K, have an effect on the feed evaporation in EVB1, which will not be further discussed here. 3.3. System Design. Together with Rosetti Marino, a 3D CAD model of the fuel processor was set up, which proved the high compactness of the system (see Figure 12, left). Safety and maintenance issues could be addressed, and specifications for periphery components such as pumps and valves are provided. The entire fuel processor will be put into a pressurized tubular housing to release the mechanical stress from the reactors. The fuel processor will be positioned upright to achieve a minimum footprint of the device. All tubing connections are consequently oriented to the same direction (see Figure 12, right). The maintenance of the fuel processor will be carried out by lifting the tube. A CAD model of the fuel processor along with the pressure swing adsorption unit is shown in Figure 13. The footprint of the system will be very small with 1.5 2.5 m at a height of 3.0 m.
4. CONCLUSIONS AND OUTLOOK The application of microstructured plate heat-exchanger/ reactor technology allowed for the design of a very compact unit. The future steps toward realization of the system will be the construction of the reactors by introduction of the microchannels into the plates by wet chemical etching, coating of the structured plates with catalyst, and sealing of the reactors by laser welding. Finally, the reactors and heat-exchangers need to be assembled along with balance-of-plant components to the fuel processor, and the control hard- and software need to be installed. 2560
dx.doi.org/10.1021/ie100602w |Ind. Eng. Chem. Res. 2011, 50, 2554–2561
Industrial & Engineering Chemistry Research
’ AUTHOR INFORMATION Corresponding Author
*E-mail:
[email protected].
’ REFERENCES
€ (1) Trimm, D. L.; Onsan, Z. I. Onboard fuel conversion for hydrogen-fuel-cell-driven vehicles. Catal. Rev. - Sci. Eng. 2001, 43 (1&2), 31. (2) Comas, J.; Marino, F.; Laborde, M.; Amadeo, N. Bio.ethanol steam reforming on Ni/Al2O3 catalyst. Chem. Eng. J. 2004, 98, 61. (3) Kolb, G. Fuel Processing for Fuel Cells, 1st ed.; Wiley-VCH: Weinheim, 2008. (4) Reay, A. D. Catalytic combustion: Current status and implications for energy efficiency in the process industries. Heat Recovery Syst. CHP 1993, 13, 383. (5) Redenius, J. M.; Schmidt, L. D.; Deutschmann, O. Millisecond catalytic wall reactors: I Radiant burner. AIChE J. 2001, 47 (5), 1177. (6) Liguras, D. K.; Verykios, X. E. Hydrogen from biomass for heat and power production. Chem. Eng. Trans. 2004, 4, 31. (7) Fatsikostas, A. N.; Verykios, X. E. Reaction network of steam reforming of ethanol over Ni-based catalysts. J. Catal. 2004, 225, 439. (8) Coleman, L. J. I.; Epling, W.; Hudgins, R. R.; Croiset, E. NI/MgAl mixed oxide catalyst for the steam reforming of ethanol. Appl. Catal., A 2009, 363, 52. (9) Llorca, J.; Homs, N.; Sales, J.; Ramirez de la Piscina, P. Efficient production of hydrogen over supported cobalt catalysts from ethanol steam reforming. J. Catal. 2002,̂209, 306. (10) Batista, M. S.; Santos, R. K. S.; Assaf, E. M.; Assaf, J. M.; Ticianelli, E. A. High efficiency steam reforming of ethanol by cobaltbased catalysts. J. Power Sources 2004, 134, 27. (11) Lin, S. S.-Y.; Daimon, H.; Ha, S. Y. Co/CeO2-ZrO2 catalysts prepared by impregnation and coprecipitation for ethanol steam eforming. Appl. Catal., A 2009, 366, 252. (12) Wanat, E. C.; Venkatamaran, K.; Schmidt, L. D. Steam reforming and water-gas shift of ethanol on Rh and Rh-Ce catalysts in a catalytic wall reactor. Appl. Catal., A 2004, 276, 155. (13) Toth, M.; Doemoek, M.; Rasko, J.; Hancz, A.; Erdohelvi, A. Reforming of ethanol on different supported Rh catalysts. Chem. Eng. Trans. 2004, 4, 229. (14) Liguras, D. K.; Goundani, K.; Verykios, X. E. Production of hydrogen for fuel cells by catalytic partial oxidation of ethanol over structured Ni catalysts. J. Power Sources 2004, 130, 30. (15) Vaidya, P. D.; Rodrigues, A. E. Insight into steam reforming of ethanol to produce hydrogen for fuel cells. Chem. Eng. J. 2006, 117, 39. (16) Men, Y.; Kolb, G.; Zapf, R.; Hessel, V.; L€owe, H. Ethanol steam reforming in a microchannel reactor. Process Saf. Environ. Prot. 2007, 85 (B5), 413. (17) G€orke, O.; Pfeifer, P.; Schubert, K. Kinetic study of ethanol reforming in a microreactor. Appl. Catal., A 2009, 360, 232. (18) Cai, W.; Wang, F.; Van Veen, A. C.; Descorme, C.; Schuurman, Y.; Shen, W.; Mirodatos, C. Hydrogen production from ethanol steam reforming in a micro-channel reactor. Int. J. Hydrogen Energy 2010, 35 (3), 1152. (19) Zapf, R.; Becker-Willinger, C.; Berresheim, K.; Holz, H.; Gnaser, H.; Hessel, V.; Kolb, G.; L€ob, P.; Pannwitt, A.-K.; Ziogas, A. Detailed characterization of various porous alumina based catalyst coatings within microchannels and their testing for methanol steam reforming. Chem. Eng. Res. Dev. 2003, A 81, 721. (20) Kolb, G.; Zapf, R.; Hessel, V.; L€owe, H. Propane steam reforming in micro-channels - results from catalyst screening and optimization. Appl. Catal., A 2004, 277, 155. (21) Kolb, G.; Tiemann, D. Mikroverdampfer, 14.5.2005, 2006. (22) Aicher, T.; Full, J.; Schaadt, A. A portable fuel processor for hydrogen production from ethanol in a 250 Wel fuel cell system. Int. J. Hydrogen Energy 2009, 19 (34), 8006.
ARTICLE
(23) Fishtik, I.; Alexander, A.; Datta, R.; Geana, D. A thermodynamic analysis of hydrogen production by steam reforming of ethanol via response reactions. Int. J. Hydrogen Energy 2000, 25, 31. (24) Francesconi, J. A.; Mussati, M. C.; Mato, R. O.; Aguirre, P. A. Analysis of the energy efficiency of an integrated ethanol processor for PEM fuel cell systems. J. Power Sources 2007, 167, 151. (25) Giunta, P.; Amadeo, N.; Laborde, M. Hydrogen production from ethanol steam reforming: fixed bed reactor design. Int. J. Chem. React. Eng. 2008, 6, (Article A12). (26) Arteaga, L. E.; Peralta, L. M.; Casas, Y.; Castro, D. Optimal design, modeling and simulation of an enthanol steam reforming reactor. Int. J. Chem. React. Eng. 2009, 7, (Article A59). (27) Kolb, G.; Schelhaas, K.-P.; Wichert, M.; Burfeind, J.; Hesske, C.; Bandlamudi, G. Development of a Microstructured Methanol Fuel Processor coupled to a High Temperature PEM Fuel Cell. Chem. Eng. Technol. 2009, 32 (11), 1739. (28) Kolb, G.; Sch€urer, J.; Tiemann, D.; Wichert, M.; Zapf, R.; Hessel, V.; L€owe, H. Fuel Processing in Integrated Microstructured Heat-Exchanger Reactors. J. Power Sources 2007, 171 (1), 198. (29) O’Connell, M.; Kolb, G.; Schelhaas, K. P.; Schuerer, J.; Tiemann, D.; Ziogas, A.; Hessel, V. The development and evaluation of microstructured reactors for the water-gas shift and preferential oxidation reactions in the 5 kW range. Int. J. Hydrogen Energy 2010, 35, 2317. (30) O’Connell, M.; Kolb, G.; Schelhaas, K. P.; Sch€urer, J.; Tiemann, D.; Ziogas, A.; Hessel, V. Development and evaluation of a microreactor for the reforming of diesel fuel in the kW range. Int. J. Hydrogen Energy 2009, 34, 6290. (31) Ioannides, T. Thermydynamic analysis of ethanol processors for fuel cell applications. J. Power Sources 2001, 92, 17. (32) Benito, M.; Padilla, R.; Sanz, J. L.; Daza, L. Thermodynamic analysis and performance of a 1 kW bioethanol processor for a PEMFC operation. J. Power Sources 2007, 169, 123. (33) Zalc, J. M.; Sokolovskii, V.; L€offler, D. G. Are noble metal-based water-gas shift catalysts practical for automotive fuel processing? J. Catal. 2002, 206, 169. (34) TeGrotenhuis, W. E.; King, D. L.; Brooks, K. P.; Holladay, B. J.; Wegeng, R. S. Optimizing microchannel reactors by trading-off equilibrium and reaction kinetics through temperature management. 6th International Conference on Microreaction Technology, IMRET 6, New Orleans, USA, 11-14 March, 2002; AIChE Pub. No. 164: New Orleans, USA, 2002; p 18. (35) Baier, T.; Kolb, G. Temperature control of the water-gas shift reaction in microstructured reactors. Chem. Eng. Sci. 2007, 62 (17), 4602.
2561
dx.doi.org/10.1021/ie100602w |Ind. Eng. Chem. Res. 2011, 50, 2554–2561
