Experimental Investigation of a JP8 Fuel Processor: Autothermal

Ind. Eng. Chem. Res. 2010, 49, 1577–1587

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Experimental Investigation of a JP8 Fuel Processor: Autothermal Reformer and CO-Cleanup Train Federico Barrai and Marco J. Castaldi* Department of Earth and EnVironmental Engineering, Henry Krumb School of Mines, Columbia UniVersity, New York, New York 10027

This paper presents an experimental investigation of a fuel processor consisting of a JP-8 autothermal reforming (ATR) reactor and a surrogate-fed CO-cleanup train. The CO-cleanup train, comprising a water-gas shift (WGS) and two preferential oxidation (PROX) reactors, was tested as an integrated reactor train. A finnedwall ATR reactor was examined for light-off behavior and for steady-state product distribution, upon which the CO-cleanup train was designed. The thermal and chemical transient analysis during catalyst ignition indicated that the fuel undergoes deep oxidation to CO2 and H2O until 80% of the catalyst bed is ignited, followed by a significant rise in synthesis gas production. The WGS and PROX reactors were tested individually with the objective of identifying operating regimes for maximum CO removal. The PROX reactor train, consisting of two identical reactors connected in series, reduced the CO concentration from 1% to less than 6 ppm. The PROX-1 and PROX-2 reactors were compared in order to elucidate the CO conversion and selectivity loss observed for PROX-1 at T > 250 °C and for PROX-2 at temperatures between 120 and 145 °C, suggesting that the CO conversion decrease follows different controlling mechanisms for the two reactors. Finally the CO-cleanup train was tested as three reactors in series, illustrating the critical effect that the CO conversion in the water-gas shift reactor has on the downstream PROX reactors. The CO-cleanup train was operated at the maximum conversion, demonstrating the capability to decrease the CO concentration from 8% to single-digit ppm level. Introduction Much of the recent work on fuel processors is focused on the employment of surrogate fuels, biofuels, gasoline, and diesel1-8 and to a lesser extent on JP8 fuel.9-14 Since JP8 is the fuel selected by the Department of Defense (DoD) and North Atlantic Treaty Organization (NATO) to be the exclusive battlefield fuel,15 a better understanding of its properties and behavior is paramount for auxiliary power applications. Alternatively fuel processors can be utilized for bio-oil upgrading16 as well for combustion enhancement in internal combustion engines.17 A fuel processor can be built in a variety of configurations depending on the required product quality. The first processing step is accomplished by the reforming reactor, which can be composed of one of the following: (i) the partial oxidation (POX) reactor, which operates exothermically; (ii) the steam reforming reactor (SR), which operates endothermically; (iii) the autothermal reformer (ATR), which ideally combines POX and SR within the same catalyst bed to ensure thermally neutral operation.18 The H2/CO ratio in the product gas depends upon steam-to-carbon (S/C) and oxygen-to-carbon (O/C) ratios. For S/C ) 2 and O/C ) 1 the ATR reactor typically yields H2/CO ) 4 with a fresh catalyst19,20 and lower values for an aged catalyst.21,22 Subsequent processing steps are targeted toward synthesis gas conditioning (adjustment of H2/CO ratio) or CO removal. If high purity H2 is required to serve as a feed to a low temperature PEM fuel cell, CO has to be removed thoroughly in a CO-cleanup train. This reactor train consists typically of a water-gas shift (WGS) reactor and one or more preferential oxidation (PROX) reactors. This work presents an experimental investigation of a novel finned-wall ATR reactor. An analysis of the temperature and * To whom correspondence should be addressed. E-mail: mc2352@ columbia.edu. Tel: 212-854-6390. Fax: 212-854-7081.

product distribution during catalyst ignition with JP8 fuel is included, which to our knowledge has not been reported in literature yet. Additionally a CO cleanup train, consisting of WGS and PROX short contact time (SCT) reactors has been tested in series and as individual reactors. Autothermal Reforming. ATR reactors convert fuel into a synthesis gas mixture rich in H2, by utilizing the reaction enthalpy provided by exothermic oxidation reactions to run endothermic ones. The overall process is simplified in the following reaction scheme, showing the POX and SR reactions (equations 1 and 2, respectively). C11H21(gas) +

11 21 O f 11CO + H2 2 2 2 ∆Ho ) -966.8 kJ/mol (1) 43 H 2 2 ∆Ho ) +1694 kJ/mol (2)

C11H21(gas) + 11H2O(gas) f 11CO +

While an autothermal operation is ideally achieved by adjusting the relative amounts of fuel, oxygen, and water, this is never attained in practice. The enthalpy provided by the partial oxidation of the fuel must allow for unavoidable heat-losses, which are typical for high temperature units, as well as a selectivity penalty due to competing reactions. Hence the reactor is required to operate with an excess air input compared to the thermally neutral case. In particular, the conversion to H2 is going to be altered by parallel and series reactions, such as the methanation, steam, and thermal cracking and water-gas shift reactions.8,23 Flytzani-Stephanopoulos and Voecks8 found that cracking reactions play an important role in ATR, and their observation was confirmed by recent modeling and experimental work by Gould et al.,2 and Dorazio and Castaldi.23 The latter

10.1021/ie901735x  2010 American Chemical Society Published on Web 01/13/2010

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Ind. Eng. Chem. Res., Vol. 49, No. 4, 2010

have pointed out the importance of WGS and cracking reactions in autothermal reforming of tetradecane. The thermal imbalance between exothermic and endothermic reactions generates an axial temperature profile which has been reported by several sources.8,18,19,24 Because the exothermic reactions are much faster than the endothermic ones and take place in the upstream portion of the reactor, the endothermic reactions are unable to locally utilize all of the energy released to ensure thermal neutrality. This results in a high reactor temperature upstream and a temperature decrease in the downstream portion of the reactor. The endothermic reactions rely on heat transported downstream by convection and conduction, as well as radiation from the reactor walls. ATR of fuels such as higher hydrocarbons and particularly JP8 poses additional challenges because: (i) coke formation and prereactor ignition need to be suppressed,25 which can be accomplished by carefully controlling the fuel evaporation temperature and by introducing a suitable amount of steam;26 (ii) the need for clean vaporization of the fuel is paramount to avoid the formation of residue and carbon deposits;25,27 (iii) sulfur tolerant catalysts and carriers are required.28,29 Although considerable literature regarding fuel processing and in particular the ATR process has been published,2,5,8,12,30-33 complete fuel processors have received limited attention in the literature, most of which has pertained to integrated reactorheat exchanger systems, designed to operate as a stand-alone unit. Qi21 and Kolb34 reported successful testing of fuel processors in reactors integrated as heat exchangers. While this approach is of great value, due to the inherent complexity of such systems, it prevents independent analysis of the reactors. Roychoudhury et al.19 have tested individual components of a fuel processor, addressed reactor-specific issues and proved the interconnectability of the individual reactors. For the present study, the employed fuel is JP8. JP8 is a middle distillate, principally made of kerosene and blended with additivesssuch as icing, corrosion, and static inhibitors.35 Initial boiling point and final boiling point for JP8 are typically 155 and 290 °C, containing a range of alkanes, cycloalkanes and arenes in the C8-C16 range. By specification JP8 contains up to 3000 ppm of sulfur originating from fuel-bound fossil sulfur and sulfur-containing additives.36 However JP8 sulfur concentrations found in literature are typically below 1500 ppm. The average of concentrations found in literature is 714 ppm with a standard deviation of 414 ppm,9,10,12-14,22,36-39 reflecting a large variability in sulfur content. Concerns about the effects of sulfur in a JP8 fuel processor were not addressed by this research. Water-Gas Shift. The WGS reactor is downstream of the ATR. It can be designed to achieve a required H2/CO ratio,40 or designed for bulk CO removal. The WGS reaction shifts an oxygen atom from water to CO, according to the following mildly exothermic reaction: CO + H2O(gas) h CO2 + H2

∆Ho ) -41.2 kJ/mol (3)

If CO removal is desired, the WGS is designed for the highest possible conversion. Because of the exothermic nature of the WGS reaction, higher CO conversions are favored at lower temperatures. At these conditions, however, the WGS reaction may reach kinetic limitations.41,42 H2 concentration is lower in autothermal reforming than steam reforming. Thus the lower H2 concentration in the ATR reactor shifts the WGS equilibrium toward higher CO conversion in the WGS reactor, making a single-stage approach feasible.40 This is due to the effect that lower H2 partial pressure has on equilibrium via Le Chatelier’s

principle. As reported herein, a CO concentration of approximately 1% is the upper boundary to ensure acceptable PROX operation. The WGS has to operate close to thermodynamic equilibrium and at temperatures as low as possible in order to get a CO concentration of approximately 1% as the PROX-1 feed. Preferential Oxidation of CO (PROX). The last processing step is accomplished by the PROX reactors, designated for decreasing the CO concentration from approximately 1% to ppm-level. The maximum CO concentration allowed into the anode of a fuel cell that employs Nafion as electrolyte is nominally set at 10 ppm.29,43-46 A higher CO content irreversibly poisons the Pt anode, suppressing fuel cell performance. This specification requires the CO-cleanup train to operate at a very high conversion: assuming that the CO concentration at the ATR exhaust is 10%, the CO-cleanup train must remove 99.99% of the incoming CO in order to produce 10 ppm or less in the output stream. The PROX reactor has been investigated extensively in the past decade.45-51 The following reactions take place in the PROX reactor: the first is the CO oxidation (reaction 4), which is the desired reaction. The competing reactions, which diminish selectivity to CO oxidation, are H2 oxidation (5), reverse WGS reaction (6), methanation reaction (7). 1 CO + O2 f CO2 2 1 H2 + O2 f H2O(gas) 2 CO2 + H2 h H2O(gas) + CO

∆Ho ) -283 kJ/mol ∆Ho ) -242 kJ/mol

(4) (5)

∆Ho ) 41.2 kJ/mol

(6) CO + 3H2 h CH4 + H2O(gas)

∆Ho ) -206 kJ/mol (7)

Experimental Details Individual reactors were tested to evaluate the performance of each reactor independently and to determine potential target operating windows. The sequence was chosen to follow the process flow of a fuel processor: ATR, WGS, PROX-1, PROX2. After the identification of an appropriate operating window on the ATR, the ATR product distribution was reproduced as a surrogate feed to the WGS reactor, to ensure precise control of feed conditions and for ease of testing. Next the PROX-1 was tested with the same gas composition as the WGS product distribution. Finally the PROX-2 was tested in series with the PROX-1 reactor. The objective of the two-stage approach is to convert the bulk of the CO in the first stage, and perform the conversion down to ppm-level in the second stage. Next the WGS reactor was connected to the PROX reactor train, and the integrated CO-cleanup train was tested. Catalysts. The ATR catalyst was spray-coated on a finned wall reactor by BASF Catalysts Inc. The catalyst is a commercial catalyst with a double layer of washcoat with a total loading of 3.25 kg m-3 of precious metal. The precious metals are a mixture of Pt, Rh, and Pd. Both the WGS and the PROX reactor tests were conducted on short contact time (SCT) screens. The WGS tests were performed using a platinum-based catalyst with a 1.9% loading, whereas PROX tests were conducted on 1.2% Pt (the loading is on a weight basis, per unit metal+support). For both WGS and PROX catalysts Pt dispersion was 47%. A 5% tetraamine platinum nitrate solution was used as a Pt precursor. The γ-Al2O3 support was stabilized with 8.2% La.


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Figure 2. Details of the ATR catalytic fins reactor: (A) technical drawing of the upper half-shell of the ATR reactor, displaying fin thermocouple placement; (B) cross section of ATR reactor in the surrounding of the central fin, illustrating the placement of fin and skin thermocouples; (C) internal view of the ATR reactor (lower half-shell), displaying the catalyzed fins.

Figure 1. A simplified process flow diagram of the experimental apparatus for the ATR reactor.

A 0.21 molar lanthanum nitrate solution was used as a La precursor. The SCT screens have 30 wires per linear inch, each wire is 0.008 in. thick, and each element has a diameter of 3/4 in. The screens are stacked into the reactor to achieve the necessary bed length, to attain the targeted catalyst weight. The space velocity in this work is reported in gas hourly units, at standard temperature and pressure (293.15 K and 101.3 kPa). For the SCT reactors the volume employed for this calculation is that of the entire stack of screens, whereas for the ATR reactor the flow volume between the fins is used. Autothermal Reforming Reactor Test-Rig. The ATR testrig, depicted in Figure 1, is comprised of a fuel delivery system, a reactor, and the gas analysis section. Deionized water and fuel are introduced by means of two in-house designed syringe pumps. In order to ensure continuous feed delivery, each pump was equipped with a backup tank, such that the main tank could be refilled while maintaining reactant flow. The flow is regulated via needle valves (Swagelok SS-SS4) and calibrated rotameters (Gilmont No. 12 for JP8 and Aalborg 082-03A for water). The water is vaporized to 400-500 °C in a furnace. The fuel is preheated approximately to 200 °C in a 1/16 in. tube coaxial to the steam pipe and vaporized upon contact with superheated steam at 450 °C. The fuel is injected countercurrent to the steam flow into a mixing chamber. Air is preheated to 300 °C and introduced to the JP8/steam mixture prior to reactor inlet and its flow is monitored with a calibrated mass flow meter and controlled with a needle valve. The ATR reactor consists of interdigitated catalytic fins designed in conjuncture with ATK/GASL Inc., as illustrated in Figure 2 and is equipped with 36 K-type thermocouples. The 24 fin-thermocouples (Omega KMQSS-040U-12) are positioned in three rows, of which 12 thermocouples are distributed along the central fin and 6 are distributed each along two lateral fins (Figure 2). The 12 skin-thermocouples are spot-welded above the fin thermocouples in the center line and measure the temperature of the outer surface of the reactor. The data acquisition system is equipped with 24 input channels, hence only 24 temperature readings could be recorded at any one time. A set of experiments was conducted with the 24 fin thermocouples connected to the data acquisition system in order to measure radial temperature gradients, from which it was determined that radial temperature gradients are negligible.

During subsequent tests only the central fin and skin thermocouples were employed. The design of a catalytic-wall reactor was motivated by the desire to observe the catalyst light-off process and to measure temperature profiles and product gas evolution during ignition and propagation of the reaction front. The ATR reactor, insulated with ceramic-fiber blankets, was preheated with air and steam from the feed delivery system, up to the desired light-off temperature. Next the fuel and air were adjusted and set to their nominal value ensuring temperatures below 900 °C. Assuming a fully developed velocity profile in the system, the travel time from the location where the air was mixed, to the entrance of the catalytic section of the reactor was approximately 40 ms. An in-house analysis has shown that JP8 has an average H/C ratio of 1.9, and its chemical structure can be formulated as C11H21, which is confirmed by others.2,10,52,53 The fuel was determined to contain 0.1 ppm sulfur. The JP8 fuel employed for the ATR tests was acquired by Gage Products Co (batch number: 40021-B#31664 JP-8). CO-Cleanup Train Test Rig. The CO-cleanup train test rig consists of a gas delivery manifold, that mixes N2 (UHP grade), H2 (high purity), CO (UHP), CO2 (UHP), and air (industrial grade) provided via calibrated mass flow controllers (Aalborg GFCS-010378) to generate the simulated reformate feed. A multispeed syringe pump (Braintree Scientific/Razel BSP-99M) is used to introduce controlled amounts of liquid H2O into a steam generator. The steam generator is equipped with N2 sweep gas to promote a steady flow of water vapor. Inlet and outlet gas temperature measurements taken by K-type thermocouples are installed directly before and after the catalyst bed. In-house heat exchangers and interstage coolers are used. A schematic of the test rig for the integrated CO-cleanup train is shown in Figure 3, illustrating the feed delivery section and the reactor/gas analysis section. Downstream of the WGS, air is introduced via a mass flow controller and fed to the PROX-1 reactor. Interstage air is introduced to the PROX-2 reactor via a Gilmont-1060 rotameter. Gas Analysis. A gas chromatograph (GC) equipped with a thermal conductivity detector (Agilent MicroGC-3000A) is employed for gas analysis. The light species O2, H2, CO, N2, and CH4 are separated in the molecular sieve and CO2 and other light hydrocarbon species are separated on the PLOT-U/Q column. Bulk water in the product stream is removed in a cold trap (cooled with ice water at 0 °C); a Nafion dryer was used to remove trace moisture prior to sending the gases to the GC. A nondispersive infrared gas analyzer was employed for ppm-

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+ xPROX-2 )/xPROX-1 . It represents a measure λcumulative ) 2(xPROX-1 O2 O2 CO of the total amount of air that needs to be fed to the PROX train. Results and Discussion

Figure 3. A simplified process flow diagram of the experimental apparatus for the integrated CO-cleanup train is depicted.

level measurement of CO in the product stream (California Analytical Instruments, model ZRF). Data Analysis. Fuel conversion and water concentration for the ATR reactor were estimated using a carbon balance. The fractional conversion of CO for the WGS and PROX reactors is defined according to the following definition: χCO ) 1 - xCO/ xinCO. The parameter λ describes the O2 stoichiometry of the inlet in , stream to the PROX reactor.45 It is defined as λ ) 2xOin2/xCO numerically identical to the O/CO ratio. A value of λ ) 1 corresponds to the introduction of a stoichiometric amount of O2. CO selectivity for the PROX is defined according to literature45 and is the fraction of the O2 fed to the PROX, that is utilized for CO oxidation: SCO ) ∆xCO/(2∆xO2). A CO selectivity of 50% denotes that a 2-fold stoichiometric amount of air is needed to fully oxidize CO, resulting in λ ) 2. Throughout the paper λ1 and λ2 refer to PROX-1 and PROX-2 reactors. The variable λcumulative is introduced and is defined as total oxygen stoichiometry to CO into the PROX reactor train:

ATR. Ignition and Transient. Upon fuel introduction, the transient to steady state behavior includes two regimes: (i) an ignition delay from fuel introduction to catalyst light-off, (ii) light-off and propagation of the ignited front. Figure 4 illustrates the reactor temperature as a function of time from fuel introduction (t ) 0) to light-off (t ) 5 min) to propagation of the ignition wave upstream (t ) 5-14 min). During the subsequent time the reactor heats up until it reaches thermal steady state. The front-end of the reactor reaches steady state after at t ) 37 min and the back-end requires approximately 1.5 h to achieve its steady state configuration as shown in Figure 6. Skin and fin temperatures are plotted in Figure 4, where the fin temperature is higher than the skin temperature, indicating a net heat flux out of the reactor. The catalyst lights off locally in the downstream part of the reactor bed, where the light-off location is determined to be the first thermocouple reading that departs significantly from other thermocouple readings. This departure was defined to be 60 °C min-1. For an average initial reactor temperature between 260 and 330 °C, the ignition delay between fuel introduction and catalyst light-off is between 3 and 27 min, respectively. For the test presented in Figure 4 the initial reactor temperature was approximately 295 °C. While at the lower initial temperature the catalyst lights off in the rear end of the catalyst bed, for higher temperatures and flow conditions the catalyst can ignite in other locations along the catalyst bed.54-56 As can be observed in Figure 4, the ignited reaction front propagates upstream. A similar behavior was reported during the light-off of the oxidation of propylene on Pt/Al2O3.57 Others have looked into temperature and gas

Figure 4. Axial temperature profiles of the ATR reactor at different times after fuel introduction (t ) 0 min). Test conditions: S/C ) 2, O/C ) 1 and a space velocity of 30 000 h-1. Each plot shows the skin (() and fin (b) temperature profile. The temperature measured at the longitudinal position of x ) 0 in. is a gas phase temperature measured in the ATR reactor plenum.


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Figure 6. Steady state axial temperature profile of the ATR reactor tested for S/C ) 2. O/C ) 1 and a space velocity of 30 000 h-1. The plot shows the skin (O) and fin ()) temperature measurements. The temperature measured at the longitudinal position of x ) 0 in. is a gas phase temperature (∆). The table shows a test data (in mol %) for O/C ) 1 and H2O/C ) 2 in comparison with thermodynamic equilibrium composition calculated at 800 °C. Figure 5. Light-off temperature (°C) and product distribution (% mol) as a function of time (min). The fuel is introduced at t ) 5 min, immediately reflecting in O2 consumption and CO2 generation. Light-off takes place 4-6 min after fuel introduction. Synthesis gas is being detected starting from 17 min (12 min after fuel introduction), upon which CO2 concentration drops. N2 concentration is not shown. Concentrations are on a dry basis.

concentration profiles for steady-state reacting systems with methane as a fuel20,58 Figure 5 shows that during the light-off and propagation of the ignited front, O2 consumption and CO2 evolution are observed immediately after fuel introduction, illustrating that at first complete oxidation is taking place. CO2 concentration rises sharply upon light-off and levels off when half the reactor is ignited. After CO2 concentration peaks at 15.5%, it decreases as CO concentration rises. Since the feed to the ATR has O/C ) 1, synthesis gas generation is expected, but this is not the case. Only traces of CO and H2 are initially recorded with subsequent increase in CO and H2 concentrations. Complete O2 conversion is recorded when the catalyst is fully active. CO and H2 production increase when 80% of the catalyst is ignited, which signals that the overall oxidation mechanism has switched to partial oxidation, also confirmed by a dip in CO2 concentration. Williams et al.7 have reported a CO2 concentration maximum during the light-off of higher alkanes during partial oxidation. However, the CO2 evolution was accompanied by substantial synthesis gas generation, which is not the case in the present work. Complete oxidation during light-off has already been reported from in situ studies of methane partial oxidation.59-62 During ignition the catalyst surface was reported to be changing from an oxidized to a reduced state.61,62 When the catalyst is in the oxidized state, only complete oxidation occurs on the active sites and after O2 is completely utilized the reforming reactions start taking place. Kimmerle et al.62 have reported that when ignition occurs at the axial location corresponding to complete oxygen conversion, a structural change in the catalyst’s oxidation state takes place, forming reduced metal sites, which in turn catalyze reforming reactions. A similar event may be occurring here. In the case of the deep-oxidation of aromatics and volatile organic compounds (VOC), catalyst ignition was reported to be inhibited by the strong adsorption of aromatic hydrocarbons on the Pt sites, which form a carbonaceous overlayer at low

temperatures and inhibit O2 adsorption.63 Actual ignition was demonstrated to occur when O2 penetrates this overlayer and starts a reaction front.64 Figure 4 and Figure 5 show that the catalyst first becomes active in the downstream end. Before fuel introduction the temperature profile in the reactor is constant at approximately 300 °C. After fuel introduction the temperature starts to increase along the reactor, due to oxidation of the fuel (Figure 5). The increasing temperature rise along the bed could promote the desorption of reaction intermediates, facilitating the ignition process by allowing the adsorption of O2 to the surface. A similar mechanism was reported to occur during CO oxidation on Pt. CO is well-known to be strongly adsorbed on the metal sites, inhibiting adsorption of O2 and hydrocarbons.65 On the other hand O2 was reported to inhibit hydrocarbon adsorption on a wide range of catalysts.66 Since in the downstream end of the reactor the temperature heats up the fastest, the selfpoisoning effect is less severe, hence the light-off is most probable in the farthermost downstream portion of the reactor,67 or at the location where local ignition criteria are satisfied.54,55 Steady-State Operation. The ATR reactor was tested with a steam to carbon ratio varied between 1.8 and 2.4 and oxygen to carbon ratio varied between 0.8 and 1.2. After light-off the reactor was allowed to reach thermal steady state. Figure 6 illustrates a longitudinal temperature profile along the ATR reactor, showing both fin and skin temperatures. The highest temperature reading is observed right after the reactor inlet, which is in agreement with literature.20,58 The inlet thermocouple illustrated in Figure 6, is located in the reactor-feed plenum, a noncatalyzed chamber where the feed tubes from the feed delivery system are manifolded and mixed prior to entering the catalyzed fins. The plenum temperature measurement is higher than the feed gas temperature entering the reactor, since it is affected by conduction and radiation from the catalytic fins. A comparison of skin and fin temperatures explains the very high temperature excursion within the bed. At the first fin temperature recording, a large heat loss takes place, as shown by the large difference between fin and skin temperature (defined as ∆Tf-s) of 45 °C. Here the heat of reaction liberated by exothermic reactions is not being utilized by endothermic reactions, resulting in a net temperature rise. Moving toward the reactor center, temperatures drop considerably, showing that endothermic reactions are being initiated. ∆Tf-s ) 0 for the third

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Ind. Eng. Chem. Res., Vol. 49, No. 4, 2010 Table 1. WGS Test Data Summary: For Each Investigated Space Velocity Results Corresponding to Maximum CO Conversion Are Presented. CO Concentrations Are Presented on a Dry Basis ˙ V GHSV υ kc χCO χout T (slpm) (h-1) (cm s-1) (cm s-1) (%) (mol %) (°C) 4.94 3.29 2.47 1.65

Figure 7. WGS test data: CO conversion outlet concentrations are plotted vs reactor exit temperature. Conversion is measured for a range of space velocities. The dashed line represents a thermodynamic equilibrium calculation (performed excluding CH4 from the equilibrium calculation). The solid line represents the CO conversion calculated with a 2-phase plug flow model with mass transfer resistances.

temperature location, reflecting that the reactor is adiabatic at this axial location. This signals that endothermic reactions are subtracting heat from the bulk of the reactor. The rest of the downstream portion of the reactor exhibits decreasing temperatures and since ∆Tf-s > 0, heat is being released to the surroundings. The table in Figure 6 shows the ATR outlet product distribution for the given temperature profile and operating conditions as well as thermodynamic equilibrium results calculated at 800 °C. The measured concentration matches thermodynamic equilibrium values at 800 °C. As illustrated in Figure 6 the measured temperature at the exit of the ATR reactor is 650 °C, which is 150 °C lower than the equilibrium temperature calculated for the measured product concentration. This behavior is reported in literature as well.5 This discrepancy could be attributed to the catalyst temperature being considerably higher than the measured fin temperature at a given axial location. Since the fin thermocouples are located in the metal fin 1 mm from the catalyst surface, this is unlikely. Assuming a linear temperature distribution between skin and fin temperature measurements at the ATR inlet (∆Tf-s ) 45 °C), the difference between catalyst surface and temperature at the fin thermocouple location is approximately 2 °C. Water-Gas Shift. The feed concentration to the WGS reactor is representative of the ATR reactor product distribution for S/C ) 2 and O/C ) 1, consisting of H2 ) 26.7%, CO ) 8.3%, CO2 ) 8.5%, N2 ) 31.2%, H2O ) 25.2%. The reactor is tested for four different space velocities (from 13 600 h-1 to 41 000 h-1), and for a range of reactor inlet temperatures (between 260 and 310 °C). Figure 7 illustrates CO conversion vs reactor bed outlet temperature for the WGS reactor. The maximum conversion of 89% is measured for the test at 13 600 h-1 and 290 °C. As expected, the data shows an increase in reaction rate with increasing temperature, until the thermodynamic limit is reached, as well as an increase in conversion, with decreasing space velocity. The curve at the highest space velocity, corresponding to 41 000 h-1, reaches equilibrium conditions at 340 °C with an 84% CO conversion. In Figure 7 the data is accompanied by model results, showing a good agreement with experiments. The WGS reactor was modeled with a two-phase plug flow model with mass transfer resistances. A power-law kinetic expression was used,68 with reaction orders for CO and H2O set at 0.28 and 1.11,

41 000 27 000 20 500 13 600

86.4 57.6 43.2 28.8

3.4 3.0 2.7 2.3

82.3 85.8 87.5 88.9

1.47 1.18 1.04 0.92

335 305 295 290

respectively, and an activation energy of 58 kJ mol-1. The preexponential factor was used as a fitting parameter and set at 1.26 × 106 in order to fit the test data. The kinetic parameters are in good agreement with literature values.69 Halving the space velocity from 41 000 to 20 500 h-1 at 275 °C increases the CO conversion from 69% to 79%, but a further decrease to 13 600 h-1 increases the CO conversion only from 79% to 82%. This behavior can be elucidated by investigating the theoretical impact of space velocity on the reaction rate. A reduction of space velocity from 41 000 to 13 600 h-1 is accompanied by a reduction in the mass transfer coefficient from 3.4 to 2.3 cm s-1. The benefit of a decrease in space velocity is counterbalanced by the negative impact that a lower linear velocity has on mass transfer between the bulk phase and the catalyst surface. This is evident in Figure 7 in the lower GHSV range, reflecting a lower sensitivity to space velocity than expected in the kinetic limit. A comparison between mass transfer and kinetic rate at 300 °C shows that while the apparent kinetic rate at the reactor exit condition is 130 mol m-3 s-1, the mass transfer rate is calculated between 73 and 50 mol m-3 s-1 (for 13 600 and 41 000 h-1, respectively). Since the apparent kinetic rate is larger than the mass transfer rate by a factor of 1.8-2.6, the process is mass transfer limited. The mass transfer rate is calculated assuming complete mass transfer control at the reactor exit condition and mass transfer coefficients are calculated using experimentally obtained correlations from literature.70 Table 1 presents a summary of the WGS tests: for each space velocity the maximum conversion χCO is reported as well as the average linear velocity V across the wire mesh screens and the mass transfer coefficient kc. A space velocity of 20 500 h-1 was chosen as the optimal one, since incrementally decreasing the space velocity beyond this value has a minor effect on conversion, that is, increasing the space velocity by 50% from 27 000 to 41 000 h-1 caused a drop in conversion by ∆χ ) 2.5%, while the same increase from 13 600 to 27 000 h-1 decreases the maximum achievable conversion by only ∆χ ) 1.3%. Preferential Oxidation, First Stage (PROX-1). The PROX performance envelope was investigated by changing the O2 -toCO stoichiometry (λ). This was accomplished by varying the air flow rate corresponding to 0.6 e λ e 4.3 and for each test the reactor gas inlet temperature was held constant while allowing the reactor gas exit temperature to reach steady state. For λ ) 2.4 the feed to the reactor had the following composition: H2 ) 31.7%, CO ) 0.95%, CO2 ) 14.4%, N2 ) 33.8%, H2O ) 18%, O2 ) 1.15%, which is a simulated WGS product distribution with the addition of air. The PROX-1 reactor is tested for a space velocity of 176 000 h-1 and inlet temperature of 150 °C. Test results for a λ sweep of the PROX-1 reactor are presented in Figure 8. The CO conversion is plotted for a range of λ, along with the reactor exit temperature. From λ ) 0.6 up to 2.2 both the temperature and CO conversion increase slowly with increasing λ. The temperature increases by 6-10 °C for every


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Figure 9. The markers denote CO effluent concentrations plotted against reactor exit temperature: (b) test results in the absence of CO in the feedstream and (∆) PROX-1 test results for a feed containing 5000 ppm CO, 33% H2, 20% H2O, 17% CO2, λ ) 2.4, balance N2, 150 000 h-1 space velocity. The solid line denotes thermodynamic equilibrium CO concentration. Figure 8. PROX-1 test results: reactor hysteresis from cyclical air introduction is shown. CO conversion and selectivity are shown in the upper portion of the plot, reactor temperature rise in the lower portion of the plot. Reactor inlet temperature is fixed at 150 °C. Data for increasing λ is shown with (2) symbols, data for decreasing λ with (b). Selectivity is marked with open symbols (O) and (∆).

10% CO conversion increase. An energy balance across the catalyst bed, assuming constant heat capacity, predicts 12 °C o ∆Hr/(FtotalCpavg)). After temperature increase (∆Tad ) -χCOFCO an initial dip, selectivity increases to its maximum value before light-off takes place. After λ ) 2.2 a steep increase in reaction temperature by 85 °C can be observed, resulting in an exit temperature of 245 °C, accompanied by a sharp conversion rise and complete O2 conversion. For λ ) 2.4 CO conversion reaches its maximum value of 97% and CO selectivity drops from 0.6 to 0.4. An additional increase in λ causes a CO conversion drop to a value of 88% for λ ) 4.4. Next the air flow was progressively decreased. From Figure 8 it can be observed that by decreasing the air flow rate to a value corresponding to λ ) 2.4, the temperature, conversion, and selectivity closely reproduces the data at the same conditions. A further reduction in air flow rate to a value corresponding to λ ) 2.2 results in an increase in CO conversion and selectivity. In the open interval between λ ) 1.8 and 2.4 the reactor operates at two different steady states and the performance plot exhibits a hysteresis. This is a well-documented behavior for exothermic reactors operating adiabatically and at high conversion efficiencies.54,71-73 A further decrease in λ results in a decrease in temperature (as well as conversion) back to the low temperature steady state. According to test results the PROX-1 reactor is capable of a 95-97% CO conversion resulting in a CO exit concentration of 300 and 500 ppm, for λ1 between 2 and 2.4. No methane was detected during PROX experiments, which is consistent with literature.46,50,74 Since the CO conversion drop could be attributed to CO generation via the reverse WGS (rWGS) reaction, an additional test was conducted in the absence of CO as a reactant.49,75 As is shown in Figure 9, CO is first detected for T > 230 °C, and reaches 500 ppm at 280 °C. The measured CO concentration is far from thermodynamic equilibrium, as confirmed by others.76 These rWGS results are compared with the PROX data, showing that CO concentration measured in the PROX product is comparable to the case of the CO-less test. The two data series converge, reflecting that the rWGS reaction is relevant at higher temperatures. The divergence at the lower temperatures could be attributed to the different test conditions: the test in the

absence of CO was conducted in an isothermal reactor, whereas the PROX reactor exhibited a temperature rise of approximately 80 °C, justifying the lower CO concentration of the PROX test compared to the rWGS test in Figure 9. A performance plot for the PROX reactor (such as Figure 8) can be divided in three operating regimes: A low-temperature regime, where an increase in temperature causes an increasing CO conversion, typical of catalyst activation plots (λ ) 0.6-2.2); a midtemperature regime, where the CO conversion is approximately constant (λ ) 2.2-2.4); a high-temperature regime, which exhibits a CO conversion drop (λ > 2.4). Numerous works have focused on understanding the mechanism underlying the conversion decrease in the high temperature regime. Some work has concluded that CO conversion declines because of the onset of the rWGS reaction, which can take place at lower measured temperatures than predicted due to heat transfer limitations.46,50,76-78 Other works have attributed the CO conversion drop to CO and H2 surface coverage, demonstrating that at higher temperatures the ratio between CO and H2 surface coverage decreases with increasing temperature.45,74,75 Alternatively the CO conversion drop has been attributed to mass transfer limitations.79 H2 conversion data does not provide an indication of the possible reason underlying this CO conversion drop, since H2 stoichiometry is identical in both the H2 oxidation and the rWGS reaction (reactions 5 and 6). Both reactions consume 1/2 mol of H2 and produce 1 mol of H2O, hence the selectivity to H2 oxidation reaction does not enable discrimination of the mechanism by which CO conversion declines. In the low temperature regime, low conversion and high selectivity are observed. This is attributed to CO saturating the Pt surface sites80 and inhibiting any other species to be adsorbed on the catalyst surface.45,81 With increasing temperature, however, the CO saturation coverage drops from 100% at 110 °C, to 75% at 155 °C, down to 55% at 205 °C, as shown with XANES operando experiments by Guerrero et al.81 This behavior is attributed to the competitive adsorption of CO, H2, and O2, and explains the steep conversion rise observed above 150 °C as well as the drop in selectivity. Integrated PROX-1-PROX-2. Both PROX-1 and PROX-2 were tested with the identical SCT catalyst beds (16 short contact time wire-mesh screens per reactor), each operated at a space velocity of 176 000 h-1. The strategy adopted is to set the optimal λ in the first reactor and operate PROX-2 at different λ values in order to reach the maximum conversion. Since the first reactor exhibited significant temperature increase, its

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Figure 10. PROX-1-PROX-2 reactor train test results: PROX-2 reactor outlet measurements are presented as CO concentration (2), CO selectivity ((), CO conversion (b) and temperature (×) vs λ2. PROX-1 was operated at λ ) 2.4 with an inlet temperature of 120 °C.

product gases had to be cooled prior to feeding them to the second stage. This was accomplished by removing a portion of insulation from the transfer line between the two reactors, allowing natural convection with the ambient. Operating conditions for the PROX-1 were held constant at λ1 ) 2.4 and an inlet temperature of 150 °C, allowing the reactor to reach steady state. The PROX-2 reactor was preheated to 120 °C and interstage air is introduced (see Figure 3) in order to obtain λ2 values between 0 and 20. In the previous section it was shown that the PROX reactor decreases CO concentration from 1% to 300-500 ppm when operated in its optimal operating window. To reduce the CO concentration to single-digit ppm level, the second stage PROX has to ensure a CO conversion higher than 97-98%. Figure 10 presents test data for the PROX-1 and PROX-2 train. The CO exit concentration is plotted against λ2. Here λ1 is held constant to a value of 2.4, and λ2 is allowed to vary. Since the smallest increment that could be achieved in air flow rate (of 2 cm3 min-1) is equivalent to λ2 ) 2, this has a drastic effect on CO conversion, from 300 to 6 ppm, after which further introduction of air results in a progressive decrease in the CO conversion. As expected, the CO selectivity decreases accordingly. The temperature rise between catalyst inlet and exit in the PROX-2 reactor was measured from 0 °C up to 25 °C. This test shows that the PROX train can drive the CO concentration down to ppm level in an operating window of λ1 between 2.2 and 2.6 and λ2 between 2 and 6, starting from a feed composition of 1% CO. This results in a cumulative CO conversion for the PROX train of 99.9% and 45% selectivity to CO. PROX-1 and PROX-2 Comparison. As illustrated in Figure 8 and Figure 10, PROX-1 and PROX-2 display a similar behavior, as they both feature three operating regimes: a window at which conversion increases, one at which it reaches its maximum value, and one at which the conversion decreases. Since the two reactors operate at concentrations that differ by 2 orders of magnitude and nearly 100 °C apart, this similarity is not expected. This is especially the case for the CO conversion decrease, which for PROX-1 occurs at temperatures above 250 °C, whereas for PROX-2 it takes place between 120 and 150 °C. It is therefore likely that there are two controlling mechanisms. For PROX-1 at the high temperature regime the mass transfer rate is approximately 10 mol m-3 s-1, whereas the kinetic rate (calculated with kinetic parameters obtained from literature45) is approximately 100 mol m-3 s-1, denoting a mass

Figure 11. Integrated CO-cleanup test data: PROX-1 CO conversion (∆) and outlet temperature (b) as a function of the CO conversion in the WGS reactor.

transfer limited regime. Additionally the application of Mears’ criterion82 indicates that in this regime the PROX-1 reactor is heat transfer limited, justifying the onset of the rWGS reaction at gas phase temperatures below what is expected.75 For the PROX-2 reactor however, both mass transfer and kinetic rate are in the same order of magnitude (approximately 0.1 mol m-3 s-1) and Mears’ criterion indicates negligible heat transfer limitations. These calculations and test data shown in Figure 9 point toward the importance of the rWGS reaction in the high temperature regime of the PROX-1 reactor. On the other hand for the PROX-2 reactor this conclusion is not definitive. Figure 10 shows that the CO conversion drop occurs for reactor temperatures between 120 and 145 °C, as opposed to temperatures higher than 240 °C for PROX-1. At these low temperatures the rWGS reaction is unlikely to play a significant role. Integrated WGS-PROX-1-PROX-2. Finally the WGS and the PROX reactors are operated in seriessas the CO cleanup train. Since the purpose of this unit is CO removal, all reactors are operated in the regime where they achieve the highest CO conversion. The integration of the reactors required interstage heat exchangers, whose performance had to be adjusted to reactor performance. According to the WGS test results reported in Figure 7, the optimal operating condition is selected at a space velocity of 20 500 h-1 and an inlet temperature of 290 °C. In these conditions the WGS product contains 1.04% CO (Table 1), which is a suitable CO feed for the PROX-1 reactor. Figure 11 illustrates the PROX-1 reactor performance as a function of CO conversion in the WGS reactor. This test was performed by increasing the WGS operating temperature in order to gradually decrease the WGS CO exit concentration, until the CO exit concentration is low enough to enable lightoff in the PROX-1 reactor. The air supply to the PROX reactors was kept constant throughout the test, hence it was possible to present λ1 on the secondary horizontal axis of Figure 11. When the conversion of the WGS increases to a value higher than 79%, the PROX-1 reactor lights off and CO conversion Table 2. Integrated CO Cleanup Train Test Summary WGS

PROX-1

PROX-1 + 2

Tin (°C) Tout (°C) GHSV (h-1) λ1+2a

285 310 20 500

145 245 176 000 2.4

125 128 88 000 4

2.45

CO conversion CO selectivity H2 conversion CO out (ppm)

87.5%

96.5% 42% 2.9% 372

98.5% 45% 3% 6

99.43% 39% (PROX) 3.1% (PROX) 8

a

9500

Cumulative λ.

CO-cleanup train


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Figure 12. Exit CO and H2 concentration of the CO-cleanup train as an integrated system. The WGS reactor is operated at 20 500 h-1 and an inlet temperature of 290 °C. The CO (2) and H2 (O) exit concentrations are plotted against the λ2 value. Plot A contains PROX-2 reactor exit concentrations for λ1 ) 2.4, and plot B contains results for λ1 ) 2.6.

surges from 20% to its maximum value. This is reflected by the PROX reactor exit temperature increasing from 170 to 240 °C. A further improvement of WGS performance results in a slight drop in PROX conversion, from 99 to 97%. While this may seem counterintuitive, this behavior signals that the PROX-1 reactor is operating in the regime at which CO conversion decreases with increasing λ and temperature, as illustrated in Figure 8. Since during this test the air feed to PROX-1 is held constant, a decrease in CO feed concentration to PROX-1 (due to increased WGS performance), results in a higher λ1 value. This is illustrated in Figure 8 where increasing λ results in a lower conversion. Conversely, a decline in WGS conversion results in the PROX conversion to drop from 98% to 20%. Such a performance drop in the WGS reactor has a dramatic impact on the CO-cleanup train product purity, which decreases from ppm-level to an unacceptable CO concentration of 2%. Table 2 presents a representative test result obtained with the CO cleanup train. This corresponds to an overall CO conversion of 99.43%, consisting of 87.5% conversion in the WGS and 98.5% conversion across the PROX train, where PROX-1 provides 96% and PROX-2 provides 98.5% conversion. The PROX-1 selectivity to CO is 43%, resulting in a hydrogen consumption of 3.1%. Overall CO conversion results for the integrated CO-cleanup train are presented in Figure 12 where the CO and H2 outlet concentrations are plotted against λ2 value. The two plots show results of the CO-cleanup train where the PROX-1 reactor was operated at two different λ1 values (Figure 12A is for λ1 ) 2.4 and Figure 12B is for λ1 ) 2.6). The trend of the H2 concentration is expected, as it decreases with increasing λ2 value, indicating a drop in CO selectivity, which is especially evident in Figure 12. Overall the test data in Figure 12 closely matches the trend plotted in Figure 10 for the PROX-1-PROX-2 test. The CO concentration decreases sharply upon introduction of air in the second PROX stage, and it reaches its minimum (6 ppm). Further increasing the air feed causes a rise in CO concentration. The lowest CO concentration was reached for a λ1 ) 2.6 and λ2 ) 5. In Figure 13, the inlet and outlet concentration for all reactors in a JP8fed fuel processor are illustrated in a bar-plot. The plot on the lower portion of Figure 13 depicts the trend of the CO concentrations (in ppm) for each reactor stream. Conclusions The experimental investigation of a JP8 fuel processor outlined in this work incorporates an analysis of a finned-wall

Figure 13. Summary of the test data of the complete fuel processor. The abscissa represents the position along the fuel processor stream. The barplot data is presented in mol % and the CO concentration is presented in ppm.

ATR reactor and a CO-cleanup train consisting of a WGS reactor and a two-stage PROX reactor train. The ATR reactor test showed that during catalyst ignition (at the back-face of the catalyst bed) the fuel undergoes deep oxidation to CO2 and H2O. This was observed during the time from fuel introduction to catalyst ignition, until 80% of the catalyst bed was ignited. Synthesis gas was only detected in significant amounts after this time, upon which the CO2 concentration dropped. The WGS reactor was tested for a range of space velocities and temperatures. The PROX-1 was operated with a CO feed concentration of 1%, whereas the PROX-2, connected in series to the PROX1, was fed with a CO concentration between 200 and 500 ppm (∆Tmax ) 10 °C). The PROX reactor exhibits a conversion and selectivity loss at the high temperature regime. This loss is well documented in literature and is believed to be either governed by a decreasing CO/H2 surface coverage ratio or a heat transfer limitation, which promotes the onset of the reverse water-gas shift reaction. A comparison between the two reactors has shown that the PROX-1 reactor performance is likely limited by the reverse WGS reaction for high CO input concentrations (due to heat transfer limitation). The CO conversion drop in the PROX-2 reactor is likely not attributed to heat transfer limitations, since in the PROX-2 reactor temperatures are measured between 120 and 145 °C. Further, the CO cleanup train was tested as an integrated unit, showing that CO levels can be reduced to single digit ppm levels. However if the WGS conversion falls below a threshold value of 80%, the PROX train becomes ineffective. Acknowledgment The authors thank Bob Farrauto (BASF Catalysts, Inc.) for numerous discussions and ATR catalysts provision, ATK/GASL for the ATR reactor fabrication. Precision Combustion, Inc. of North Haven, CT, is acknowledged for providing the PROX and WGS catalysts.

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Nomenclature -1

-1

Cp ) heat capacity, J mol K Fi ) molar flow rate of species i, mol s-1 GHSV ) gas hourly space velocity, h-1 k ) reaction rate coefficient kc ) mass transfer coefficient, m s-1 SCO ) ∆PCO/(2∆PO2) ) oxygen selectivity to CO, SCT ) short contact time T ) temperature,°C Tad ) adiabatic temperature,°C ∆Tf-s ) difference between fin and skin temperature (°C) xi ) molar fraction of species i Greek Letters λ ) 2PO2/PCO ) oxygen stoichiometry with respect to CO τ ) residence time in reactor, s χi ) fractional conversion of species i Subscripts and Superscripts 1 ) relative to the first stage PROX 2 ) relative to the second stage PROX f ) placement of the thermocouple in the fin bulk s ) placement of the thermocouple in the skin location 0 ) reactor inlet

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ReceiVed for reView July 1, 2009 ReVised manuscript receiVed December 9, 2009 Accepted December 10, 2009 IE901735X

Experimental Investigation of a JP8 Fuel Processor: Autothermal

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