Experimental and Modeling Analysis of Methane Partial Oxidation

Mar 13, 2009 - under representative operating conditions, at steady state and during light-off. Model analysis (based on a dynamic heterogeneous react...
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Ind. Eng. Chem. Res. 2009, 48, 3825–3836

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Experimental and Modeling Analysis of Methane Partial Oxidation: Transient and Steady-State Behavior of Rh-Coated Honeycomb Monoliths Alessandra Beretta, Gianpiero Groppi,* Matteo Lualdi, Ivan Tavazzi, and Pio Forzatti Dipartimento di Energia, Laboratorio di Catalisi e Processi Catalitici, Politecnico di Milano, Piazza Leonardo da Vinci 32, 20 133 Milano, Italy

The present study consists of an experimental and theoretical study of the performance of Rh-coated honeycomb monoliths for methane partial oxidation. The thermal behavior of Rh-coated honeycomb monoliths was studied under representative operating conditions, at steady state and during light-off. Model analysis (based on a dynamic heterogeneous reactor model that incorporates a kinetic scheme of the process independently developed, and well-assessed correlations for heat and mass transfer) provided a key for interpreting the observed effects. The comprehension of how transport phenomena and surface kinetics affect the reactor behavior leads to the conclusion that the feasibility of small-scale production of syngas via CH4 catalytic partial oxidation relies on thermal management of the short contact time reactor and not the obtainment of high syngas yields (which is not a challenging task). Severe operating conditions (and high surface temperatures) can deplete the catalyst activity and cause unstable reactor operation. Guidelines for optimal reactor design are proposed. 1. Introduction Catalytic partial oxidation (CPO) of CH4 is considered a promising technology for H2 and syngas production in smalland medium-scale applications in the energy and transportation sectors.1,2 Micro power devices, distributed heat and power systems, and on-board fuel processors are some examples. A challenge for process development is to optimize the reactor design and the operating conditions, which requires application of a reliable reactor model, accounting for both surface reaction and heat and mass transfer mechanisms. Recent works have treated the modeling of CH4 autothermal reformers for design purposes, mainly packed-bed reactors. Ramaswamy et al.3 have theoretically analyzed the behavior of a short contact time fixed-bed reactor, packed with spheres. However, several transient and steady-state effects (e.g., thermal reactor response to steam injection) discussed by the authors are strictly associated (for magnitude and trends) with the adopted kinetic scheme, that is, the kinetics by Xu and Froment4 [steam reforming, CO2 reforming, and water gas shift (WGS)], which were developed for Ni catalysts operating at high pressure, and by Ma et al.5 (methane oxidation), which were developed for Pt catalysts. Such a combined kinetic scheme had been originally proposed by de Smet et al.6 for simulating industrial-scale reactors for syngas production via autothermal reforming (that is, combined oxidation and reforming) of methane and later applied in several modeling analyses of partial oxidation and autothermal reforming reactors for fuel cell applications.7-9 Bizzi et al.10 have modeled the partial oxidation of methane over Rh in a fixed-bed reactor and applied the detailed kinetic scheme by Schwiedernoch et al.;11 the authors have analyzed literature data and have explored the effects of operating variables, deriving guidelines for optimal reactor design and operation. Important experimental efforts have been devoted by Schmidt and co-workers12-14 to comprehending the axial resolution of concentration and temperature profiles in foam monoliths, and a recent modeling analysis by Dalle Nogare et al.14 has further * Corresponding author.

assessed the key role of transport phenomena for correctly describing the consumption of reactants; however, available correlations for Sh and Nu numbers in foams were found to be inadequate, and adapted correlations were proposed. Less treated in the literature is the modeling of honeycomb monoliths that, in the perspective of industrial application, present pros and cons with respect to packed beds and foams. Hohn and Schmidt15 and Maestri et al.16 have experimentally and theoretically shown that lower heat and mass transfer coefficients enhance the “resistance” of packed beds to extinction even at extremely high flow rates. Still, the better heat transfer properties and lower heat capacity of honeycomb monoliths result in much shorter start-up time16 and are expected to increase the dynamic response of the autothermal reformer to inlet changes. Also, at flow rates of practical interest, honeycombs perform better than packed beds in terms of conversion, selectivity, and pressure drop. These features make honeycomb monoliths a suitable choice for miniaturized applications, including on-board fuel reforming and distributed power generation. Additionally, honeycomb monoliths offer several advantages for experimental investigation and its quantitative analysis; among them (1) representativeness of the laboratory-scale data; (2) availability in the open literature of well-established methodologies for catalyst deposition in uniform, compact and thin layers (a controlling factor for reproducibility of experimental results); (3) ability to monitor the axial temperature profiles along the channels by means of multiple sliding thermocouples without the need of TC capillaries; and (4) reliability of correlations for heat and mass transfer coefficients even at very low Reynolds numbers. Most of these aspects make honeycombs also preferable to foams for kinetic investigation. The present paper addresses the combined experimental and modeling analysis of steady-state and transient behavior of a partial oxidation pilot-scale reformer, using Rh-coated honeycomb monoliths. This work extends previous investigations reported in the literature (Schwiedernoch et al.11 presented a complete model of methane partial oxidation over monoliths, but validation was limited to steady-state conversion/selectivity measurements; detailed kinetics were applied by Poulikakos and

10.1021/ie8017143 CCC: $40.75  2009 American Chemical Society Published on Web 03/13/2009

3826 Ind. Eng. Chem. Res., Vol. 48, No. 8, 2009

Figure 1. Schematic drawing of reactor assembly.

co-workers17-19 to analyze in theoretical works the optimal reactor design for maximum H2 yield; and Veser and coworkers20 have experimentally investigated the performance of “sliced” Pt monoliths and provided a semiquantitative interpretation of detailed data), since well-characterized experimental data are analyzed with a comprehensive mathematical model that incorporates adequate kinetics (previously developed over the same Rh catalyst) and state-of-the-art correlations for heat and mass transfer. The aim of the paper is, on one side, to gain an insight into the kinetic factors and heat and mass transfer properties that influence the performance and, specifically, the thermal behavior of the reactor, and on the other side, to assess the reliability of the reactor model. In this respect, start-up dynamics were specifically investigated, since surface kinetic effects are more stringently probed. It is herein recognized that the monolith thermal behavior controls the reactor stability, and its management represents the key issue for practical application of the process. Solutions for optimal heat management are explored. 2. Experimental Section Experiments of CH4 CPO were performed in an adiabatic reformer over a 4 wt % Rh/R-Al2O3 catalyst supported on 400 and 600 cpsi honeycomb cordierite monoliths (# ) 2.2 cm, L ) 1.7-2.7 cm). Rh/R-Al2O3 powders were prepared by incipient wetness impregnation of R-Al2O3 with an aqueous solution of Rh(NO3)3; catalyst deposition was realized by dipping the monolith into a slurry of the catalytic powders, followed by blowing of the excess slurry. The catalyst loading (typically 300-350 mg) was estimated by weight difference before and after coating. The catalyst layer thickness (7-15 µm) was estimated by assumption of a washcoat density of 1500 kg m-3, on the basis of independent measurements. The reactor used in this work is a stainless steel cylinder with an internal quartz lining, which prevents C formation. The internal layout (shown in Figure 1) consists of the catalytic monolith between two FeCrAlloy foams, kept at proper distance, that act as thermal radiation shields and flow mixers. The reactor is equipped with several sliding thermocouples that measure the radial inlet gas temperature and the axial temperature profile. Thermal insulation was obtained by wrapping the static mixer, the lines upstream of the reactor, and the reactor with a very thick layer of quartz wool taping. A continuous gas analyzer (ABB AO2000) acquires the dynamics of the product distribution, while GC analyses are performed at steady-state conditions. 3. Mathematical Model The behavior of the CPO reformer was analyzed by a dynamic 1D, heterogeneous, single-channel model of the adiabatic reformer, previously developed16 and validated;21 the mass,

energy, and momentum balance equations are reported in Table 1. As expected, calculated pressure drops were always negligible, under all investigated conditions. Table 2 reports the heat and mass transfer correlations. Relevant reactor and catalyst parameters are listed in Table 3. The indirect reaction scheme implemented by the model was independently derived, based on a kinetic study in an isothermal annular reactor; it includes oxidations of CH4, CO, and H2, steam reforming of CH4, and the water gas shift reaction and its reverse. Rate expressions and kinetic parameters are reported in Table 4 and are discussed in detail in previous works.22-24 In particular, Donazzi et al.23 have specifically addressed the kinetic role of CO2 reforming, which was found to be negligible within the kinetic scheme of CH4 CPO over Rh. 4. Results and Discussion 4.1. Steady-State Temperature and Concentration Profiles: The Paradox of “Direct” Syngas Formation. Figure 2 reports the measured axial temperature profile under reference operating conditions (10 Nl/min, O2/CH4 ) 0.54, preheated feed). In this and all the experiments, measurements were obtained by sliding along central monolith channels both the frontal thermocouple and the back-end thermocouple; differences were detected. In general it was observed that, within the catalytic portion, the frontal TC tended to underestimate the highest inlet reactor temperatures. The back TC likely provided a more reliable measurement of the temperature profile within the catalytic portion but tended to overestimate the inlet gas temperature. The observed differences are in line with the expected errors of the TC signal due to some conduction along the TC itself. The conduction errors in TC measurements are well-known errors; they affect the TC output signal when the tip and the upstream TC body are subject to important temperature gradients, which is exactly the situation that occurs in the present experiments once the frontal TC tip just enters the catalytic monolith (causing an underestimation of the hotspot temperature), or when the back TC (whose body is heated by the catalytic monolith and the product mixture) is pushed so that the tip is exposed to the inlet feed mixture (whose temperature is then overestimated). Model calculations within the catalytic zone were thus preferentially compared with the back TC measurements. A close match between frontal TC and back TC measurements were always found in the ending portion of the monolith and downstream from the monolith. In particular, both measured profiles showed a few degrees temperature drop at the exit of the catalytic monolith; this was associated with a “cup-mixing” effect, downstream of the monolith, due to a non-perfectly flat radial velocity profile across the reactor (as better illustrated in the following, the expectedly higher linear velocity of the central channels is accompanied by higher local temperatures compared to the peripheral channels). Accordingly, the value of measured

Ind. Eng. Chem. Res., Vol. 48, No. 8, 2009 3827 Table 1. Model Equations Gas Phase Balances species mass balance equation

av ∂ωi G ∂ωi )- Kmat,i(ωi - ωi,wall) ∂t Fgε ∂z ε

energy balance equation

av 1 ∂Tg G ∂Tg )- h (T - Ts) ∂t Fgε ∂z ε Fgcˆp,g g

momentum balance equation

(

-

)

G2 dTg 1 G2 1 G2 dp - 2 ) + 2 a f Fg 2F 2 v Fg p dz Fg Tg dz g Solid Phase Balances

species mass balance equation

0 ) avFgKi,mat(ωi - ωi,wall) +

(

)

NR

∑ν

eff

i,jrj

j)1

MWiFsξ

kaxeff ∂2Ts ∂Ts 1 1 ) avh (Tg - Ts) + + ∂t Fscˆp,s(1 - ε) Fscˆp,s ∂z2 cˆp,s(1 - ε)

energy balance equation

(∑[ NR

]

-∆HjR) rjeff ξ

j)1

Boundary Conditions reactor inlet (z ) 0)

ωi,z)z1 ) ωi,feed

-kaxeff

∂Ts ∂z

|

reactor outlet (z ) Lrct)

Tg,z)z1 ) Tfeed

(

Pz)z1 ) Pfeed

∂Tg ∂z

|)

) σεs Tg4 - Ts4 z1

|

)0 z)Lrct

-kaxeff

z1

∂Ts ∂z

|

(

|)

) -σεs Tg4 - Ts4 z2

z2

Initial Conditions

ωi(z, 0) ) 0

Tg(z, 0) ) Tg,feed

Table 2. Transport Correlations for Honeycomb Monoliths

Nu ) 2.977 + [8.827(103Z*)-0.545] exp(-48.2Z*) Sh ) 2.977 + [8.827(103Z*)-0.545] exp(-48.2Z*) Z* )

f)

z for Nu dcellRePr

14.3 Re

Z* )

z for Sh dcellReSc

dcellG Re ) µε

temperature downstream from the monolith was an important reference for model simulations, thermodynamic evaluations, and estimation of the reactor thermal efficiency ∆Texp/∆Tad, in which ∆Texp [) T(zout) - T*(z ) 0), with T*(z ) 0) measured in the absence of reaction by replacing the catalytic monolith with an inert one] is the measured temperature increase along the reactor and ∆Tad is the calculated temperature increase for

Ts(z, 0) ) Troom Table 3. Reactor Geometrical and Physical Properties Fcat (kg m-3) Fmonolith (kg m-3) cˆp (kJ kg-1 K-1) ks (W m-1 K-1) drct (m) front inert bed (m) back inert bed (m) catalytic bed length (m) washcoat (µm) ε

1500 3800 0.865 3 0.022 0.5 0.5 0.017-0.027 7-15 0.78-0.81

an adiabatic reactor with the same inlet temperature and the same performance (conversion and selectivity). In the case of the experiment reported in Figure 2, thermal efficiency amounted to 0.98. Figure 2 also reports the calculated solid-phase and gas-phase temperature axial profiles; note that the value of the inlet gas temperature (an input datum for the simulation) was provided by the independent measurement of a fixed radial thermocouple, placed several centimeters upstream from the catalytic monolith. It is clear that the measurements of the sliding thermocouples are qualitatively and quantitatively in agreement with the predicted gas-phase temperature (with the presence of a rounded

3828 Ind. Eng. Chem. Res., Vol. 48, No. 8, 2009 Table 4. Kinetic Scheme and Parameter Estimatesa ri (mol gcat-1 s-1)

ki873K (mol atm-1 gcat-1 s-1)

Eact (kJ/mol)

1.030 × 10-1

92

1.027 × 10-1

92

6.239 × 10-2

25

1.276 × 10-2

62

2.638 × 103

62

1.938 × 101

76

CH4 Total Oxidation: CH4 + 2O2 f CO2 + 2H2O

ktotoxPCH4

rtotox )

1 + kadsH OPH2O

σO2

2

CH4 Steam Reforming: CH4 + H2O T 3H2 + CO

rSR )

kSRPCH4(1 - ηSR) 1 + kadsCOPCO + kadsO PO2

σH2O

2

Direct Water Gas Shift: CO + H2O f CO2 + H2

rWGS )

kWGSPH2O(1 - ηWGS)

ηWGS < 1

(1 + kads,H2OPH2O)2

σCO

Reverse Water Gas Shift: CO2 + H2 f CO + H2O

rRWGS ) kRWGSPCO2(1 - ηRWGS)σH2 ηRWGS < 1

H2 Oxidation: H2 + 1/2O2 f H2O

rH2ox ) kH2oxPH2σO2 CO Oxidation: CO + 1/2O2 f CO2

rCOox ) kCOoxPCOσO2

a

surface adsorption

kadsi873K (atm-1)

∆Hads (kJ/mol)

O2 H2O CO

5.461 3.901 × 102 2.114 × 102

-73 -16 -37

σi ) Pi/(Pi + 10-6). ηSR ) KP,SR/Keq,SR with KP,SR ) PCOPH23/(PCH4PH2O). ηWGS ) KP,WGS/Keq,WGS with KP,WGS ) PCO2PH2/(PCOPH2O). ηRWGS ) 1/ηWGS.

maximum in the initial portion of the catalytic zone). The calculated solid-phase temperature, instead, is characterized by an abrupt increase at the very inlet section of the bed, followed by a progressive decrease. This trend is more easily interpreted by considering the corresponding calculated axial concentration profiles in the bulk gas phase and at the catalyst surface, reported in Figure 3. Herein, it is clearly shown that the surface O2 concentration drops to zero at the very entrance of the reactor; the overall consumption of O2, thus, is extremely fast and controlled by external mass transfer. Within an initial portion of the reactor, O2 is completely consumed from the gas phase, while methane conversion proceeds along the whole honeycomb length. When the surface concentration profiles are examined, H2O and CO2 are formed at the inlet and then consumed, while syngas is progressively formed along the axial coordinate. Thus the heat release due to O2 conversion (oxidations of methane, CO, and H2) is concentrated at the reactor entrance, while the heat consumption that drives syngas formation involves the whole catalyst bed; this originates the sharp hot-spot at the inlet section. It is worth noticing that, upon examination of the calculated concentration profiles, a paradox is apparent: the gas-phase

Figure 2. Measured and calculated temperature axial profiles under reference operating conditions (400 cpsi, monolith length ) 2 cm, flow rate ) 10 Nl/min, O2/CH4 ) 0.54, inlet gas temperature ) 273 °C).

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Figure 3. Calculated axial concentration profiles in the bulk phase and at the wall. Conditions were as in Figure 2.

Figure 4. Measured and calculated effect of flow rate under autothermal operation of a 400 cpsi monolith with O2/CH4 ) 0.54. Table 5. Measured and Calculated Effect of Flow Rate on Reactor Performance (Autothermal Operation) flow rate, Nl/min 5 10 16 19 21

CH4 conversion, %

H2 selectivity, %

CO selectivity, %

exp

calc

exp

calc

exp

calc

79.9 84.5 83.6 83.5 82.7

81.5 82.5 80.6 79.6 78.8

91.9 92.7 92.5 92.3 92.0

91.9 91.6 90.7 90.2 89.8

83.1 86.3 87.0 87.3 87.4

85.4 85.4 85.4 85.6 85.7

concentration profiles have in fact the expected trends of a direct partial oxidation mechanism, since syngas is apparently formed in the presence of O2. Instead the complete conversion of O2 at

the surface allows steam reforming to “spread” along the whole bed length, and some mixing between O2 (diffusing from the gas to the surface) and syngas (diffusing from the surface to the bulk) occurs in the inlet portion of the honeycomb channels. The relative amount of H2 and CO at the wall is affected by their consecutive oxidations (and H2 oxidation is much faster than CO oxidation); this is the reason why the H2/CO molar ratio at the wall is always well below the stoichiometric value of steam reforming, while it increases along the axial coordinate toward the thermodynamic value. A flatter (through still growing) trend of H2/CO in the bulk phase is calculated due to the effect of the enhanced H2 diffusion rate within the gas phase. Notably,

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Figure 5. Measured and calculated effect of flow rate with feed preheating. Preheating temperatures were 273 °C (10 Nl/min), 378 °C (16 Nl/min), 395 °C (19 Nl/min), and 408 °C (21 Nl/min). Other conditions were as in Figure 4. Table 6. Measured and Calculated Effect of Flow Rate on Reactor Performance (Preheated Feed) CH4 conversion, %

H2 selectivity, %

CO selectivity, %

flow rate, Nl/min

exp

calc

exp

calc

exp

calc

10 16 19 21

91.4 94.0 92.4 93.4

90.6 93.2 93.8 93.0

94.9 95.4 95.7 95.5

94.7 95.7 95.6 95.9

91.5 93.4 94.2 94.2

91.8 93.7 93.9 94.1

the calculated gas-phase concentration profiles in Figure 3 closely resemble the axially resolved concentration measurements reported by Schmidt and co-workers12 for Pt- and Rhcoated foam monoliths and by Veser and co-workers20 with Pt monoliths at different lengths. Those measurements were associated with the existence of a contribution from a direct partial oxidation reaction; the present results show that the same trends can be generated by an indirect reaction process at the wall, in combination with the effect of interphase mass transfer limitations. Such a paradox has been elucidated in details in a recent paper using a microkinetic description of the CPO process.25,26 These results are fully in line with the conclusion of Dalle Nogare et al.14 on the prevailing role of mass transfer on the consumption of reactants in the oxidation zone of catalytic foams. 4.2. Effect of Flow Rate. The effect of flow rate in short contact time partial oxidation reactors using Rh or Pt catalyst has been discussed already in the literature, based on experimental and/or theoretical evaluations.3,7-9,15,16,18-20 Most works, however, refer to the behavior of packed bed with coated spheres or foams. In the specific case of honeycomb monoliths, experiments on Pt were discussed by Veser and co-workers.20 Herein, data on Rh-coated monoliths were obtained in the range 5-21 Nl/min, with and without preheating. Figure 4 reports the measured temperature profiles at varying total flow rate under autothermal conditions, that is, without any preheating of the feed mixture. Measurements from the back TC only are reported for simplicity. The increase of flow rate shifted the measured temperature maximum downward and widened it; calculations showed that this trend was closely in line with the effect of flow rate on gas-phase temperature. Much higher surface temperatures were predicted at the reactor inlet, and hotspot temperatures grew at increasing flow rate. Such a temperature increase compensated the shortening of contact time and

in fact the outlet conversion and selectivity were poorly sensitive to flow rate and widely controlled by thermodynamics, as shown by the data reported in Table 5. As already discussed by Veser and co-workers,20 in the case of honeycomb monoliths, wherein the effect of linear velocity on Sh and Nu numbers is negligible, the main effect of increasing the flow rate is the increased heat release by the reaction, which results in higher inlet temperatures at the wall. Because of the effect of enhanced convective heat transfer, cooling of the catalyst phase at the entrance and shift of the catalyst hot-spot downward have been predicted for packed-bed reactors.3 According to our simulations, under conditions of practical interest, the solid-phase temperatures of Rh honeycomb monoliths are always characterized by sharp hotspots at the very inlet section. Also, at increasing importance of convection, conduction is less effective in smoothing the surface hot-spot. Convection significantly affects the gas-phase temperature, instead, whose maximum moves along the reactor. Note that a significant effect of heat dispersion can be excluded under the conditions adopted in this work. In fact, thermal efficiency was verified in each experiment. Minor deviations from adiabatic behavior were evaluated only at 5 Nl/min (∆Texp/∆Tad ) 0.95), while thermal efficiency was optimal at higher flow rates (∆Texp/∆Tad ) 1.0). The same effect of flow rate was explored by running experiments with preheated feed streams. Experimental results and model simulations are reported in Figure 5 and Table 6. Because of the increasing efficiency of the heating cartridges at increasing flow rate, the inlet gas temperature also increased from 273 to 408 °C in the range 5-21 Nl/min. Accordingly, at the highest investigated flow rate of 21 Nl/min, the highest axial temperature profile was measured. The combination of the two effects (flow rate and inlet gas temperature) was such that a moderate increase of conversion and selectivity was observed, despite the reduction of contact time. Model simulations confirmed that the combined increase of flow rate and inlet temperature led to extremely severe operating conditions; it is estimated that inlet surface temperature largely exceeded 1000 °C. Notably, since the adiabatic temperature rise is a decreasing and highly sensitive function of methane conversion and syngas yield, an important increase of the inlet temperature does not produce an equivalently important increase of the outlet temperature. This further illustrates the importance of modeling

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Figure 6. Cold start of a 2.5 cm long, 400 cpsi monolith. Comparison between dynamics of the temperature measurements at the front end and at the back end and calculated dynamics of the gas phase temperature.

Figure 7. Dynamics of the product mixture composition during cold start. Conditions were as in Figure 6.

for a complete rationalization of experiments, like those on the effect of flow rate, wherein several complex phenomena are involved. 4.3. Start-up Dynamics. A theoretical analysis of the startup dynamics was previously addressed by the authors.16 In that work, evolution of the temperature profiles of an adiabatic CPO reactor upon injection of a preheated CH4/air feed was investigated. In line with previous results from Ramanathan et al.,27 light-off was predicted to initiate at different locations of the catalytic bed, depending on catalyst activity and operating conditions. It was also shown that the dynamics of the solid

temperature was driven on one side by the rate of methane combustion (heat release) and on the other side by the rate of the heat-solid convective heat transfer (heat removal). Initial experiments in an insulated bed packed with spheres, operating at a total flow rate of 2 Nl/min, clearly indicated that ignition (that is, initial heating of the surface with local formation of a hot-spot) “moved” from the back end to the front end of the reactor at increasing preheating temperature from 350 to 500 °C.21 In this work the light-off of the CPO microreformer was studied on a more representative scale (total flow rates up to

3832 Ind. Eng. Chem. Res., Vol. 48, No. 8, 2009

Figure 8. Measured and calculated light-off at 10 Nl/min, preheated reactor at 505 °C. Measurements were sampled at each second, starting from time 0. Calculations refer to the evolution at each second, starting from time +1 s.

Figure 9. Measured and calculated light-off at 10 Nl/min, preheated reactor at 330 °C.

21 Nl/min) and reactor configuration (honeycomb monoliths instead of coated spheres), and different procedures of start-up were tested and simulated. 4.3.1. Cold Start. First, cold start was studied. At time 0, the preheated feed mixture (21 Nl/min of a methane/air feed, with O2/C molar ratio of 0.54, at 400 °C) was fed to the insulated reactor (initially at ambient temperature). The whole reactor assembly started to heat up and the temperature increase was controlled by the large heat capacity of the system. The heating dynamics was monitored by two fixed thermocouples, located at the inlet and outlet sections of the honeycomb monolith; the recorded trends are reported in Figure 6. Above 250 °C, the temperature increase at the back end was faster than that at the front end (an indication of some heat release from the catalyst surface, which gradually accumulates at the back end); once the back-end temperature reached about 340 °C, light-off occurred. The local temperature measurements showed that the

back end heated up before and to a lesser extent than the front end. Also, the back temperature passed through a sort of oscillation, reaching a minimum with following fast increase that occurred simultaneously with the front-end temperature increase. Simulations well captured these features: back-end ignition was in fact predicted and the peculiar behavior of the back-end temperature was also reproduced. Figure 6B reports the detailed simulation of the cold-start dynamics (plots refer to the calculated gas-phase temperature) and provides the explanation. Because of the assumption of an exothermicendothermic reaction sequence (wherein the exothermic reactions are faster than endothermic reactions), an initial heating of the back end first occurs and is followed by a local partial cooling. The same condition (solid heating, followed by partial cooling) realizes successively upstream at each axial position; the result is that of a “moving” hot-spot that propagates from the ignition point toward the inlet section. Once this heat wave

Ind. Eng. Chem. Res., Vol. 48, No. 8, 2009 3833

Figure 10. Stable performance of Rh monolith operated at the intermediate flow rate of 10 Nl/min: (a) under autothermal conditions and (b) with preheat.

reaches the front end, a final heating of the whole monolith yields a steady-state profile with a pronounced inlet hot-spot. Concerning the composition of the product mixture, Figure 7 compares the measured and calculated evolution of species concentration at the outlet reactor section. In correspondence with the initial heating of the reactor back end, the outlet concentration of O2 dropped to zero, and within a few seconds partial and deep oxidation products were formed. The concentrations of CO2 and H2O passed through maxima; the observed trends were well captured by the calculated dynamics, which provides an important validation to the indirect kinetic scheme, wherein the oxidation reactions are responsible for the initial bed heating and CO2 and H2O behave as intermediate reacting species. Calculated curves are delayed with respect to experimental trends, in line with Figure 6. 4.3.2. Preheated Catalyst Bed. Since the cold-start procedure is irreducibly associated with slow heating dynamics of the whole reactor assembly (which is of course specific of the present rig design), a different start-up methodology was tested. Specifically, we tried to eliminate the delaying effect of the reactor heat capacity by preheating the whole system at a desired temperature under flowing N2. At time 0, the feed flow was switched from N2 to the methane/air feed. This procedure also allowed us to “move” the ignition point along the bed by manipulating the reactor preheat temperature. Experiments were performed at 10 Nl/min, with O2/C ) 0.54 and gas inlet temperature of about 280 °C. Results are reported in Figures 8

and 9 for the cases of initial solid temperature of 505 and 330 °C, respectively. The temporal evolution of the axial temperature profiles [T(x) at constant time] was reconstructed by synchronizing and combining the single local temperature dynamics [T(t) at constant x, like those reported in Figure 6], recorded in repeated experiments; experiment by experiment, the thermocouple tip was placed in a different axial coordinate and kept in the same position during the transient test until steady state was reached, when it was then slid along the whole length to check the reproducibility of the reactor behavior. Injection of the feed mixture in the reactor, preheated at 505 °C (Figure 8), caused a front-end ignition. A hot-spot formed at the entrance within 3-4 s, and propagation from the front to the back end took about 10 s. At lower initial catalyst temperatures, the ignition point “moved” in the middle of the reactor; at 330 °C initial solid temperature (Figure 9), ignition occurred at the back end. An initial induction of about 7 s was still observed, after which a rapid temperature increase at the back end occurred; propagation of the heat wave upward to the front end took about 15 s. Figures 8 and 9 report the calculated evolution of the gasphase temperature for comparison. Qualitatively the simulation is very satisfactory, since the effect of the solid initial temperature was well reproduced. At 505 °C initial solid temperature, front-end ignition was predicted and propagation was simulated to occur in about 10 s. At 330 °C, instead some disagreement is evident between simulations and experiments. The model did not predict any induction period and the heating of the catalyst bed was complete within about 15 s. Such a mismatch could be due to several factors: a nonuniform temperature profile of the reactor assembly (incomplete elimination of the heating dynamics as discussed above), a partial loss of activity of catalyst (the experiments reported in Figure 9 followed a large series of repeated tests), but also some inaccuracy in the description of methane combustion kinetics at low temperatures and high reactant concentrations (a condition that, though explored, was not central in the kinetic study in the annular reactor22,23). 4.4. Catalyst Aging. Reproducibility of the reactor thermal behavior during the bulk of tests on the effect of flow rate (described above) was monitored by periodically repeating standard CPO tests at 10 Nl/min. The measured temperature profiles under autothermal conditions and with preheat are compared in Figure 10. Evidently the overall reactor stability was preserved during the experimental campaign. After the modeling results, it is believed that such stability is strictly related to the fact that most of the experimental campaign did not stress the surface catalyst temperature. However, the same modeling results suggest that flow rate may represent a critical issue for catalyst stability, an issue to which the literature has paid little attention. As shown in Figure 5, operation of the reactor at very high flow rates with preheated feed streams is in fact associated with high surface temperatures (especially at the inlet), which expectedly can cause activity losses for the catalyst and performance losses for the reactor. The effect of operation under severe operating conditions was studied by performing repeated runs at 21 Nl/min with preheat. Steady-state and transient results are reported in Figure 11. We observed important changes in both steady-state performance and ignition. At steady state, consistent with previous studies,28 a marked increase of the peak temperature was observed, while the outlet temperature did not vary significantly (Figure 10a), in line with small modifications of the reactor performance in

3834 Ind. Eng. Chem. Res., Vol. 48, No. 8, 2009

Figure 11. Aging of a honeycomb monolith operated at 21 Nl/min, 420 °C inlet gas temperature. Experimental results of repeated light-off and steady-state experiments and model calculation. Simulation of run 4 was obtained by assuming that the surface reactivity was 0.2 times (in the first third of the bed length) and 0.5 times (in the rest of the bed length) the initial activity. Simulation of run 8 was obtained by assuming that the surface reactivity was 0.1 times (in the first third of the bed length) and 0.3 times (in the rest of the bed length) the initial activity.

Figure 13. Calculated effect of the aspect ratio on gas-phase and solidphase temperature. Conditions were as in Figure 12. Figure 12. Calculated effect of inlet extra activity. An activity profile is assumed such that the initial L* length is 3 times more active than the rest of the length. Simulated operating conditions: autothermal operation, 10 Nl/min, O2/CH4 ) 0.54, 400 cpsi, monolith length ) 2 cm.

terms of CH4 conversion and syngas yield. Concerning the ignition, an increase of the light-off temperature was observed. To better understand the experimental evidence, a sensitivity analysis addressed the effect that local changes of the catalyst activity have on the final reactor performance; it showed that progressive heating of the monolith and ignition delay could be described by invoking reduction of the reforming activity and of the oxidation activity, respectively. In order to take into account axial profiles of deactivation, for simplicity we introduced in the model the possibility to correct the surface reaction rates (listed in Table 3) by an “activity” coefficient with step changes across each third of the catalyst bed. The calculated light-off curves and temperature profiles in Figure 11 were obtained by assuming activity coefficients of 1 (run 2), 0.2 (run 4), and 0.1 (run 8) in the first third of the bed; the activity

coefficient amounted to 1, 0.5, and 0.33 in the second and third thirds of the bed. Although simplified, the modeling analysis shows that the observed trends can be explained by assuming a preferential loss of activity in the inlet section of the monolith (which is mostly heated). Since light-off is initiated by the onset of methane oxidation, the increase in light-off temperature was specifically informative of an important loss of oxidation reactivity. At steady state, despite the important reduction of the reaction rates, oxygen consumption was still complete and mass-transfer-controlled (as shown by the unreported calculated axial concentration profiles of the aged catalyst); instead, the loss of reforming rate was reflected in a much lower heat consumption rate, which resulted in much higher surface and gas-phase temperatures. Model simulations of the aged catalysts also show that because of the increased reactor temperatures, outlet methane conversion tends to be insensitive to the activity loss. Similar experimental effects were reported by BitschLarsen et al.,29 studying the poisoning effect on S on methane partial oxidation in Rh foams.

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5. Reactor Design: Activity Profile and Aspect Ratio Experiments and calculations clearly indicate that heat management is crucial for the reactor stability. Experience suggests that surface temperatures below 1000 °C are needed for preventing catalyst deactivation and extremely dangerous autocatalytic overheating. The results of the present study suggest some design solutions for minimizing inlet surface temperatures. Figure 12 shows the expected thermal performance of a honeycomb monolith wherein an inner length L* is present where the catalyst surface activity is 3 times higher than in the rest of the monolith. Since the rate of O2 consumption is mass-transfer-controlled, the rate of heat release is unaffected by the enhanced surface oxidation rate. On the opposite side, a local promotion of the steam reforming rate greatly affects the rate of heat consumption. An impressive beneficial effect on surface temperatures is calculated for L*/L ) 1/3, with a decrease of about 75 °C of the maximum temperature; simulations at varying L*/L show then that the same effective moderation of surface temperature is obtained even if the extra activity is localized at the very reactor inlet (for instance, 2.5 mm over a 20 mm long monolith). In fact, the beneficial effect on the hot-spot temperature depends on the rate of reforming reaction at the very inlet (axial coordinate ) 0). At axial coordinates > L*/L, the catalyst temperature profiles tend to become closer to the reference case of even activity. This is the reason why, at an intermediate coordinate, temperature increases upon passing from L*/L ) 1/3 to 1/4 to 1/8. Figure 13 explores the effect of the reactor aspect ratio (L/D ) length/diameter). It is calculated that the decrease of the aspect ratio at equal reactor volume also favors a considerable decrease of temperatures within the monolith. The effect is strictly associated with the effect of heat conduction though the monolith and the inert shields (back dispersion), which becomes important at decreasing convection. The effect of linear velocity is in fact negligible on the extent of heat and mass transfer coefficients in honeycomb monoliths. A similar effect had been presented by Bizzi et al.10 for packed beds; however, higher L/D were suggested in that work for enhancing syngas yield. We recommend, on the other hand, to take care to avoid severe surface heating. In this respect, lower L/D ratios are expected to favor stability. 6. Conclusions An experimental and theoretical investigation allowed us to gain insight into the transient and steady-state behavior of a CPO adiabatic reformer with an Rh-honeycomb monolith. Heat and mass transfer phenomena largely control the inlet wall temperatures, which, even when air is used at only slightly overstoichimetric O/C ratios, can easily overcome 1000 °C at sufficiently high flow rates and preheating temperatures. Operation under such severe conditions rapidly damages the Rh catalyst; outlet reactor performance (composition and temperature of the product mixture) is not a sensitive probe of catalyst aging, while ignition temperature and axial temperature profile are greatly affected by decreases in oxidation and reforming reaction rates, respectively. Stable performances and thermal behavior were found for operation at intermediate flow rates (10 Nl/min). Start-up dynamics are largely controlled by the initial bed temperature; at initial room temperature, the honeycomb ignition (cold start) always proceeds with a back-end mechanism (whose characteristic feature is partial oscillation of the outlet bed temperature, which is coherent with the wrongway behavior analyzed by Dudukovic and co-workers3), while

preheating of the catalyst bed at temperatures higher than 350 °C moves the ignition point progressively upstream of the catalyst bed. At 500 °C preheating temperature, front-end ignition occurs. Reaching a steady state is more rapid than in the back-end mode, since the heat-wave propagation is favored by the convective heat transfer.27 The capability of the model to capture the observed transient and stationary effects makes us confident in further addressing optimization of reactor design; activity profiles that enhance the inlet catalyst activity, and suitable aspect ratios (enhancing conduction over convection) offer room for reducing inlet catalyst temperatures. Acknowledgment Funding from MIUR-Rome (PRIN2006) is gratefully acknowledged. Nomenclature A ) reactor cross section (m2) av ) specific area per unit volume (m-1) cp ) specific heat (J kg-1 K-1) dcell ) hydraulic diameter of the monolith channel (m) drct ) reactor diameter (m) f ) friction factor G ) superficial mass flow rate (kg m-2 s-1) h ) heat transfer coefficient (W m-2 K-1) kax ) solid thermal conductivity (W m-1 K-1) kaxeff ) effective solid thermal conductivity (W m-1 K-1) Kmat,i ) mass transfer coefficient (m/s) Lrct ) reactor length (m) p ) pressure (Pa) rj ) j reaction rate (mol kgcat-1 s-1) t ) time (s) T ) temperature (K) z ) reactor axial coordinate (m) Greek Letters ε ) bed void fraction εs ) emissivity ∆HjR ) heat of reaction (J/mol) ξ ) volumetric catalytic fraction F ) density (kg/m3) σ ) Stefan-Boltzmann constant ω ) weight fraction Subscripts and Superscripts g ) gas phase s ) solid phase eff ) effective

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ReceiVed for reView November 11, 2008 ReVised manuscript receiVed January 21, 2009 Accepted January 27, 2009 IE8017143