Accepted Manuscript Multiple steady states in the oxidative steam reforming of methanol Jung Hyeon Kim, Young Shin Jang, Dong Hyun Kim PII: DOI: Reference:
S1385-8947(18)30091-3 https://doi.org/10.1016/j.cej.2018.01.075 CEJ 18395
To appear in:
Chemical Engineering Journal
Received Date: Revised Date: Accepted Date:
31 October 2017 4 January 2018 13 January 2018
Please cite this article as: J.H. Kim, Y.S. Jang, D.H. Kim, Multiple steady states in the oxidative steam reforming of methanol, Chemical Engineering Journal (2018), doi: https://doi.org/10.1016/j.cej.2018.01.075
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Multiple steady states in the oxidative steam reforming of methanol
Jung Hyeon Kim, Young Shin Jang, and Dong Hyun Kim* Department of Chemical Engineering Kyungpook National University Daegu, 41566 Republic of Korea
*Corresponding author Email:
[email protected] Abstract In the oxidative steam reforming of methanol (OSRM), the partial oxidation of methanol (POM) is commonly assumed to be the main reaction between oxygen and methanol. This assumption was experimentally tested with a commercial Cu/ZnO/Al2O3 catalyst, and it was found that, in fact, the combustion of methanol (COM) is the main reaction between oxygen and methanol in the OSRM. Furthermore, POM was not an independent reaction but a series of COM and steam reforming of methanol (SRM) reactions. In the presence of oxygen, COM was the main reaction, and, as oxygen was depleted, the SRM started and produced hydrogen. Multiple steady states were observed during the temperature cycling of the reactor between 453 and 583 K. The multiplicity was caused by the COM and was analyzed in terms of the mass and heat transfer between the catalyst particle and the surrounding gas phase. Although the reactor temperature was kept below 573 K, the upper steady state of the catalyst particle (of the two steady states) was estimated to occur at temperatures above 880 K because of the high heat of combustion and the heat transfer resistance between the catalyst and the surrounding gas phase. When the catalyst was in the upper steady state, the combustion rate was so fast that the COM was completed at the reactor inlet and SRM was the main reaction in the reactor. In the lower steady state, on the other hand, the combustion rate was low, and the COM was the main reaction in the reactor.
Keywords: Oxidative steam reforming of methanol, Multiple steady states, Hydrogen production, Methanol combustion, Partial oxidation of methanol, Catalyst deactivation.
1. Introduction Hydrogen can be easily produced by the steam reforming of methanol (SRM). SRM: CH3OH + H2O → CO2 + 3H2
ΔH = 49.5 kJ/mol (at 298 K)
(1)
One liter of methanol can produce 74 moles of H2. This is considerably greater than 33.6 moles in 1 liter of liquid hydrogen at 20 K or 19.7 moles in 1 liter of compressed hydrogen at 700 bar in fuel cell cars. Because methanol is easy to transport and readily reformed to produce H2, methanol can be regarded as a hydrogen carrier.
The steam reforming reaction is endothermic, so the reaction heat must be supplied to the reaction. The steam reformer usually contains a heat exchanger to supply the reaction heat to the catalyst bed. On the other hand, if oxygen is added to the SRM feed, the heat can be generated in situ by oxidizing a part of the methanol; this is the oxidative steam reforming of methanol (OSRM). OSRM: CH3OH + 0.8 H2O + 0.1 O2 → CO2 + 2.8 H2
ΔH ≈ 0 kJ/mol (at 298 K)
(2)
Because OSRM does not require external heating, OSRM reactors have an advantage of simple and compact design compared to steam reforming reactors. This advantage is particularly suitable for mobile and onboard hydrogen production systems in fuel cell cars.
The OSRM is not an independent reaction. Rather, it is a combined reaction of the exothermic reaction between O2 and methanol and the endothermic SRM. In the OSRM literature, the exothermic part of OSRM has frequently been assumed to be the partial oxidation of methanol (POM) [1–16]. POM: CH3OH + 1/2 O2 → CO2 + 2 H2
ΔH = -192.5 kJ/mol (at 298 K)
(3)
For this reaction, Langmuir–Hinshelwood mechanisms, in which the reaction route consists of the formation of HCHO by the partial oxidation of methanol, followed by decomposition of the intermediate HCHO to CO and H2 and CO oxidation by the water gas shift reaction or by adsorbed O species, have been proposed [2,7, 9, 11,14]. According to these mechanisms, H2 is produced directly, and H2O is only formed by the oxidation of the produced H2. Thus, H2 should be observed before any H2O in the reaction products. However, this has not been experimentally verified. In contrast, the formation of H2 is initiated after a period of pronounced H2O formation [17– 19]. In addition, because CO2 is formed from CO in the proposed mechanisms for POM, a CO concentration above the level of the water gas shift equilibrium should be observed in the POM products. However, only very low levels
of CO, less than the equilibrium concentration, have been detected in the product [12, 19, 20]. For POM to be an independent reaction, CO2 and H2 should be produced in the stoichiometric ratio shown in Eq. (3) at all O2 conversions. The stoichiometry has only been confirmed at complete O2 conversions and has not been experimentally observed for fractional O2 conversions. Before the O2 conversion is complete in the OSRM, the formation of H2O always precedes that of H2, and, consequently, the combustion of methanol (COM) has been concluded to be the reaction between O2 and methanol during the OSRM [17–21]. COM: CH3OH + 3/2 O2 → CO2 + 2 H2O
ΔH = -676.5 kJ/mol (at 298 K)
(4)
In fact, POM is a combination of COM and SRM reactions:
1 3 CH 3OH O 2 CO 2 2H 2 O 3 2 2 CH 3OH H 2 O CO 2 3H 2 3 1 CH 3OH O 2 CO 2 2H 2 2
(5)
In this combination, H2O is formed first by COM, and, subsequently, H2 is formed by SRM with the produced H2O. In this case, POM is not an actual reaction; instead, it is a combination of these two reactions. If POM occurs, the products should be produced and observed according to the above reaction stoichiometry at all O2 conversions. In this study, the product distribution in the OSRM at fractional O2 conversions is examined to identify the reaction between O2 and methanol, whether it is the POM or COM. This is important in the analysis of heat transfer from the exothermic oxidation reactions to the endothermic SRM reactions in the catalyst bed because the combustion heat is 3.5 times larger than the heat produced by the POM.
In OSRM reactors, O2 is converted first, and, on complete O2 conversion, the remaining methanol reacts with water to produce hydrogen and CO2 [4, 10, 19, 20]. The oxidation produces heat in the front section of the catalyst bed. The heat is transferred downstream and is used as the reaction heat for steam reforming. During the OSRM, the reactor temperature can be considerably non-uniform because of the poor thermal conductivities of catalyst beds (0.1–0.7 W/(m K)) [22, 23]. When the generated heat is mainly carried by the gas flow, the gas temperature increases significantly, and a hotspot is formed at the reactor inlet [24, 25]. At a feed temperature of 170C, a
hotspot of 575 C has been observed at the inlet of an OSRM reactor made of a 46-mm internal diameter (ID) and 47-cm long stainless-steel pipe [24]. The temperature difference along the radial direction of the catalyst bed is also significant. For packed-bed reactors for SRM, temperature differences as large as 40 K have been estimated along the radial direction in a 4-mm ID reactor, and, to obtain near isothermal conditions, the reactor diameter should be reduced to 0.3 mm [26]. In addition, differences of 10–35 K between the bed and the furnace were observed during OSRM in a reactor (6-mm ID), even though the catalyst in the bed was highly diluted with inert particles to enhance the dissipation of the heat generated in the catalyst particles [19, 20]. It is difficult to analyze the data obtained from a reactor in which the temperature differences along the radial and axial direction are significant. To reduce the temperature differences, the reactor diameter must be as small as possible, and the reactor should be highly conductive to allow the rapid transfer of heat. In this study, a small diameter (ID = 2.2 mm) copper tube was used as the reactor to flatten the temperature profile in the catalyst bed.
Although the gas temperature in the reactor is isothermal, the catalyst temperature can be considerably higher than the surrounding gas phase temperature because of the large amount of heat produced by the oxidation reaction and the heat transfer resistance between the catalyst particle and the fluid. As the catalyst temperature increases, the oxidation rate increases, generating more heat. This, in turn, increases the catalyst temperature and further increases the rate. This cooperative effect cannot continue indefinitely because the mass transfer resistance of the reactant from the surrounding gas to the catalyst surface takes effect and limits the reaction rate [27]. Because of the mass transfer limitation, the reaction rate or the heat generation rate of the catalyst particles typically shows an S-shaped dependence on the catalyst temperature [28]. The heat generated in the catalyst particle is removed from the particle by the heat transfer to the surrounding gas. Furthermore, the removal rate is proportional to the temperature difference between the particle and the surrounding gas and shows, in general, a linear dependence on the catalyst temperature. In the steady state, the heat generation rate should be equal to the removal rate. The S-shaped heat generation curve and the straight heat removal line may meet at more than one catalyst temperature. This gives rise to multiple steady states. In this paper, we report and analyze the multiple steady states of the OSRM.
2. Experimental 2.1. Reactor and activity measurements The reactor was made of a small diameter copper tube (3.175-mm outside diameter (OD), 2.175-mm ID). The catalyst was a commercial SRM catalyst (Synetix 33-5). The properties of the catalyst can be found in Ref. [29]. The catalyst bed in the reactor was made of 50 mg of the catalyst particles (0.25–0.3 mm diameter) diluted with 150 mg of glass beads of the same size. The bed length was around 5 cm. Thermocouples were attached to the outside wall of the reactor at the positions of the inlet, the middle, and the outlet of the catalyst bed. The three temperatures differed by no more than 3 K because of the fast thermal conduction along the copper tube, and the highest temperature was observed at the inlet of the catalyst bed where the exothermic reaction took place. It was not possible to measure the catalyst bed temperature directly because of the small reactor diameter. The outside wall temperature in the middle of the reactor was taken as the bed temperature.
The feed composition was 15% CH3OH, 25% H2O, 3% O2, and 57% He, and the total flow rate was 100 mL (STP)/min. A bubbler containing liquid CH3OH and another bubbler containing liquid H2O were used to saturate the two separate He carrier flows with CH3OH and H2O, respectively. The temperature of each bubbler was controlled by a circulating temperature bath. To examine the effect of O2 on the reaction rate, the O2 flow rate was varied in the range of 3–9 mL/min. Before reaction, the catalyst was reduced with 10% H2/He flow (50 mL/min) at 523 K for 2 h. The reactor inlet pressure was monitored to measure the pressure drop along the reactor, and the reactor outlet was vented. The inlet pressure depended on the methanol conversion and increased from 1.05 atm for no reaction to 1.15 atm for complete conversion because of the increase in the molar rate of the reaction.
An electric furnace was used to control the reactor temperature. The furnace temperature was programmed to run temperature cycles between 453 and 583 K at a rate of 0.2 K/min. In preliminary experiments, the temperature rate of 0.2 K/min was sufficiently slow to yield steady state results for a stable catalyst. It took 21.7 h to complete one temperature cycle. The effluent of the reactor was analyzed on-line with sampling at 10-min intervals using a gas chromatograph with two thermal conductivity detectors (HP 5890A). A Porapak-Q column was used to separate CO2, CH3OH, and H2O, and a Carboxen column (Supelco) was used to separate O2, CO, and CO2.
2.2. Characterization of the catalyst For the characterization of the catalyst in the reactor, the feed to the reactor was switched to a He flow of 100 mL/min, and the reactor was cooled to room temperature under He flow. Then, the reactor was dismounted from the experimental OSRM system and quickly installed in the characterization system.
To examine the oxidation state of the catalyst, we carried out temperature programmed reduction (TPR) of the catalyst using 50 mL/min flow of a 10% H2/N2 mixture at a temperature rate of 10 K/min. Before TPR, the catalyst was purged with an N2 flow of 50 mL/min at 523 K for 2 h to remove adsorbed species. The TPR was carried out from the room temperature to 673 K. The reactor effluent was monitored for the mass number of 18 for H2O with a mass spectrometer (Pfeiffer Vacuum QME200). To obtain the TPR results for the fully oxidized catalyst, the reduced catalyst after a TPR run was reoxidized with 50 mL/min of a 10% O2/N2 mixture at 553 K for 2 h.
The Cu0 surface area of the catalyst was measured by N2O chemisorption [30]. 2 Cu(S) + N2O → Cu2O(S) + N2
(6)
The catalyst was reduced at 523 K for 3 h under a 10% H2/N2 flow of 50 mL/min. The reactor temperature was lowered to 353 K under an N2 flow of 20 mL/min. At 353 K a series of 20 L N2O pulses were injected via a sixport valve at 3-min intervals into the N2 stream flowing through the reactor. For each injected N2O pulse, the mass number of 44 for N2O in the reactor effluent was monitored using a mass spectrometer (Pfeiffer Vacuum QME200). At the start, the initial 3–4 N2O pulses were completely consumed in the reactor, and the corresponding N2O peaks at the exit were not detected. After this, the N2O peak at the reactor exit begun to grow with successive injections as the reduced Cu species in the catalyst were oxidized by N2O and depleted. The areas of the successive exit N2O peaks became constant when the reaction shown in Eq. (6) was complete. The amount of N2O consumed by the catalyst was calculated as the total injected amount subtracted by the total amount eluted from the reactor. A value of 1.46 × 1019 Cu atoms/m2 was used to estimate the Cu metal area of the catalyst [30].
3. Results and discussion 3.1. Multiple steady states In the temperature programming of the reactor furnace between 453 and 583 K, the methanol conversion and the reactor temperature were measured at 10 min intervals. The temperature cycle was repeated five times. The results of the first cycle and the fifth cycle are shown in Fig. 1. The results of other cycles are omitted to avoid overcrowding the data in the figure. The methanol conversion of the first cycle and the fifth cycle show a similar trend with respect to the temperature programming, but the conversion of the first cycle at the same conditions was higher than the corresponding conversion of the fifth cycle. A decrease in the conversion as the temperature cycle repeated was also observed for other cycles (not shown). This indicates a slow deactivation of the Cu catalyst during the OSRM, which has been reported previously [19, 31].
Of the two cycles shown in Fig. 1, the fifth cycle is now examined. The methanol conversion slowly increased with temperature from point S (453K, 0) to point A (530 K, 0.09) and then rapidly increased to B (547 K, 0.76). With further temperature increase, the conversion moved from B to C (586 K, 0.99). On decreasing the temperature from point C, the conversion moved from C to D (517 K, 0.34) and dropped to point E (503 K, 0.026). The complete sequence of the cycle was S → A → B → C → D → E → S.
The conversion on the path from B to C was higher than the conversion on the path from C to D in the temperature range of 547–586 K. For example, the conversion of point (a) at 554 K on the path from B to C was 0.82, while the conversion of point (b) on the path from C to D was 0.72. It tool around 5 h from point (a) to point (b) in the temperature programming. The results for the B-to-C path were obtained under a transient state of the catalyst activity after the catalyst moved from an oxidized state in A to the reduced state in B. This is examined in Section 3.4. The oxidation state of the catalyst on the path from C to D and the state on the path from S to A were investigated by means of TPR and are discussed in Section 3.2. On the other hand, the conversion on path C to D was comparatively stable because the conversion stayed on the path when the temperature increased or decreased. The long-term stability of the conversion is examined in Section 3.4.
The conversion pathways along the paths of increasing temperature and that of decreasing temperature were different, and this gave rise to multiple steady states. In the fifth cycle, two steady states existed in the range of 517 and 530 K. For the first cycle, the temperature range of the multiple steady states is 513–525 K. On the paths S to A and E to S, the O2 conversion was incomplete. In contrast, on paths B to C and C to D, the O2 conversion was 100%. The steady states of the reactor can, thus, be characterized by incomplete O2 conversion and complete O2 conversion. The multiple steady states are analyzed in Section 3.6.
3.2. Oxidation state of the catalyst The oxidation states of the catalyst at the two steady states were investigated. For this purpose, TPR was carried out for the catalyst at a reactor temperature of 523 K for the S to A path (lower steady state) and at 523 K for the C to D path (upper steady state). For TPR, the feed was changed to a He flow of 100 mL/min, and the reactor was left to cool down under the He flow. The reactor was then dismounted from the OSRM reaction system and installed in the TPR system. In Fig. 2, the TPR peaks are shown. The TPR peak of the catalyst at the lower steady state was larger than the peak of the catalyst at the upper steady state. This indicates that the amount of the oxidized Cu in the catalyst at the lower steady state was much larger than the amount in the upper steady state.
After the TPR, the reactor was cooled to room temperature, and the reactor was taken from the system and left open for 5 min to allow some air to seep into the reactor to mimic the air contamination of the catalyst during the movement of the reactor from the OSRM system to the TPR system. The TPR peak of the air-exposed catalyst, shown as the dashed line in Fig. 2, is very similar to the peak of the catalyst in the upper steady state, indicating that the reduction peak of the catalyst at the upper steady state likely originated from the air contamination and the catalyst at the upper steady state was in the reduced state.
The reduced catalyst after TPR was oxidized at 523 K with 50 mL/min of 10% O2/N2. The TPR of the oxidized catalyst is shown by a dotted line in Fig. 2. The oxidized catalyst showed much larger peak than the peak of the catalyst at lower steady state. This indicates that, in the lower steady state where O2 conversion was incomplete, the catalyst was not in a fully oxidized state, and a significant fraction of the Cu species was in the reduced state. It has been shown that, for small O2 conversions during OSRM, Cu2+ is the dominant Cu species, whereas, after complete
O2 conversion, the Cu species was reduced to Cu0 [18–20]. In the lower steady states in the OSRM, the surface Cu species of the catalyst are in contact with O2 and Cu2+ is expected to be formed. However, the state of the interior Cu species is unclear, and they could be in the reduced state.
3.3. Methanol oxidation in OSRM Fig. 3 shows the O2 conversion and ΔH2O flow rate, which is the net change in the H2O flow rate (mL(STP)/min) between the reactor inlet and outlet, with respect to the reactor temperature in the fifth temperature cycle shown in Fig. 1. A positive ΔH2O flow rate represents the net formation of H2O in the reactor. The multiple steady states can also be monitored in terms of ΔH2O flow rate and O2 conversion. The corresponding points for A to E in Fig. 1 are also marked on the plot of the ΔH2O flow rate.
The net formation of H2O is only possible when COM takes place in the reactor because POM does not result in the formation of H2O. If COM alone occurs, the H2O formation rate for the complete O2 conversion would be 3×10-6 mol/s (4 mL(STP)/min) for a feed of 100 mL(STP)/min with a composition of 3% O2, 15% methanol, and 25% H2O. This is shown as the upper dashed line marked ‘Combustion’ in Fig. 3. From the start of the temperature cycle, the ΔH2O flow rate increased toward the ‘Combustion’ line with increasing temperature and reached a maximum around point A in Fig. 3. For a further increase in the reactor temperature, the ΔH2O flow rate decreased rapidly as the SRM began. The O2 conversion was 0.6 at point A, and, as shown, the SRM started before complete O2 conversion.
When the methanol conversion is complete, the ΔH2O flow rate would be -6.7 ×10-6 mol/s (= -9 mL(STP)/min) for the feed. This is shown as the lower dashed line marked ‘Complete conversion’ in Fig 3. The ΔH2O flow rate decreased to the line as the methanol was consumed by SRM. Along the path from C to D, the O2 conversion was complete, and the ΔH2O molar flow rate was negative, indicating that SRM consumed H2O more than the amount produced by COM.
The net changes in the rate of O2 and CH3OH through the reactor are plotted against each other in Fig. 4. Because both components are reactants, the net changes in the reactor are all negative. The experimental data are from the
fifth temperature cycle shown in Fig. 1, and point A in Fig. 4 corresponds to the same points in Figs. 1 and 3. The lines for the reaction stoichiometry of POM and COM are included. As shown, the data on the S-to-A path in Fig. 1 are located on the COM line. For -Δ O2 > 1.8 mL(STP)/min, however, the data deviated from the COM line because of the onset of SRM. This proves that the reaction between O2 and methanol is combustion and that POM is not an actual reaction that takes place in OSRM. Consequently, OSRM is a combination of COM and SRM.
3.4. Transient state and deactivation of the catalyst in OSRM After the experiment shown in Fig. 1, the reactor temperature was increased and fixed at 554 K, and the methanol feed to the reactor was cut off for 30 min. This treatment oxidized the catalyst with the O2 in the feed. The methanol feed was resumed, and the methanol conversion was measured with time. The initial conversion was 0.82, the same as that at the conversion at point (a) in Fig. 1. This conversion is shown by the upper dashed line in Fig. 5. The conversion decreased with time, and after 5 h, the conversion became 0.72, the same conversion of the point (b) in Fig. 1 and marked by the lower dashed line in Fig. 5. This result is shown as Run 1 in Fig. 5. The methanol feed to the reactor was again cut off and then resumed after 30 min. The conversion regained the initial conversion of Run 1 (i.e., the upper dashed line) and decreased with time, similar to the trend in Run 1. This shows that point (a) in Fig. 1 is not a stable steady state but a transient state toward a more stable steady state, point (b) in the figure.
Such oxidation treatment has been reported to regenerate a finely dispersed CuO phase from the sintered Cu phase [32]. As a result, the active Cu phase is redispersed, and the activity is recovered. To verify the redispersion, the Cu metal area of the catalyst was measured by N2O chemisorption [30]. The Cu metal area was 4.3 m2/g after Run 1 before oxidation, and the area increased to 6.2 m2/g after oxidation before Run 2. The increase in the Cu area confirms the redispersion of the sintered Cu phase by the oxidation treatment.
The methanol conversion of Run 3 started at 0.83 (11 h in Fig. 5), decreased to 0.72 in 5h (16 h), and further decreased to 0.66 in 31 h (42 h). In Run 4, the conversion started at 0.82 (42.5 h), decreased to 0.69 in 5 h (47.5 h), and decreased to 0.55 in 63.5 h (106 h). In Run 5, the conversion began at 0.81 (106.5 h), decreased to 0.64 at 111.5 h, and decreased to 0.51 at 147 h. Between the runs, the same oxidation treatment was applied to the catalyst. The oxidation treatment activated the catalyst, as can be seen by the repeated recovery of the conversion in Fig. 5. The
deactivation after the oxidation, however, appeared to become faster as the run was repeated because the decrease in the conversion in the first 5 h was 0.1 in Run 1 and 0.18 in Run 5. In addition, two stages of deactivation are seen in the figure, in which the first stage of fast deactivation for about 5 h is followed by much slower deactivation. This indicates two different mechanisms for the deactivation of the Cu catalyst. As shown, the initial fast deactivation is ascribed to the sintering of the finely dispersed Cu phase formed by the oxidation treatment. The later stage of slow deactivation is the sintering of the coarse Cu phase after the fine Cu phase has disappeared during the first stage. The Cu metal area at the end of Run 5 was as 3.7 m2/g, smaller than 4.3 m2/g after the initial stage of deactivation in Run 1, evidence of further sintering of the Cu phase in the OSRM.
3.5. Reaction rate of methanol combustion To calculate the heat generation by COM, the rate of COM over the catalyst is needed. For this purpose, the effect of the partial pressure of the component in the reaction on the COM rate was studied. The O2 partial pressure in the feed was varied from 3 to 6 kPa, and its effect on the methanol conversion of the reactor is shown in Fig. 6. As shown, the conversion weakly increased with increasing O2 partial pressure. The effects of methanol partial pressure and H2O partial pressure in the feed on the O2 conversion were negligible, as shown in Figs. 7(a) and 7(b), respectively.
Accordingly, the COM rate in a power law form can be expressed as rA A0 exp(
E A n ) PA RTS
mol O2 /(kg-cat s)
(7)
The pre-exponential factor ( A0 ), activation energy ( E A ), and the power ( n ) of the O2 partial pressure ( PA , kPa) can be determined by analyzing the data in Fig. 6. TS is the catalyst temperature, which can be significantly higher than the gas flow temperature because of the heat transfer resistance from the catalyst to the gas flow and the high heat of combustion.
Two cases were considered in analyzing the data of Fig. 6. The first case assumes negligible heat transfer resistance and, hence, a catalyst temperature TS equal to the gas temperature T f . The second case considers the heat transfer
resistance, and TS was calculated from the heat balance between the heat generation by the combustion reaction and the heat loss by the heat transfer from the catalyst particle to the reactor. The methods to obtain the reaction rate for the two cases are described in the Appendix.
The reaction rate obtained for the first case, in which the reactor temperature is equal to the catalyst temperature, is rA 3.6 1011 exp(
130 kJ/mol 0.4 ) PA RTS
mol O2/(kg-cat s)
(8)
In the second case, the catalyst temperature is also dependent on the heat transfer resistance between the catalyst and the reactor wall. The resistance can be represented by a heat transfer coefficient. Several correlations have been developed for this coefficient, and the correlation by Cybulski et al. [33] was used to estimate the heat transfer coefficient in the data analysis. The correlation gave hP 111 W/(m2 K) for the reactant flow through the reactor at 523 K. The combustion rate obtained with the heat transfer coefficient is rA 3.7 106 exp(
84 kJ/mol 0.14 ) PA RTS
mol O2/(kg-cat s)
(9)
The procedure for the analysis of the data to obtain Eq. (9) are given in Appendix. With a heat transfer coefficient other than hP 111 , the resulting reaction rate for the same data would be different from Eq. (9). With increasing heat transfer coefficient, the temperature difference between the catalyst and the gas flow decreases and eventually becomes zero as hP . Eq. (8) can be obtained in the second case with hP . The lines in Fig. 6 are the calculated conversion with Eq. (9) and hP 111 . The activation energy of 84 kJ/mol of Eq. (9) is considerably lower than 130 kJ/mol of Eq. (8). This is because the catalyst temperature calculated with the heat transfer coefficient is higher than the gas phase temperature. Because the reactor diameter is small, the gas phase temperature has been assumed to be the same as the measured reactor temperature ( T f T0 ). For T0 520 K and hP 111 , the catalyst temperature at the reactor inlet for the feed with 3% O2 was estimated to be 532 K.
Reitz et al. [17] obtained a power-law rate for COM over a commercial low-temperature shift catalyst (BASF K3110) in which the powers of the partial pressures of the components in the rate expression were 0.18, 0.18, and -0.14
for methanol, O2, and H2O, respectively, and the activation energy was 115 kJ/mol. They assumed the same temperature for the catalyst and the gas phase in the reactor.
3.6. Multiple steady states of a catalyst particle at the reactor inlet 3.6.1. Heat removal from the catalyst particle Before we investigate the multiple steady states of the reactor, we consider a single catalyst particle facing the feed at the inlet of the catalyst bed. The combustion of methanol takes place in the catalyst particle. The heat generated in the catalyst particle is transferred to the surrounding gas flow. The fluid–particle heat transfer can be described by
Q f p h p a p (TS T f )
(10)
where T f is the gas temperature, and a P is the outer surface area of the catalyst particle. In addition to the fluid– particle heat transfer, the radiative heat transfer can contribute significantly to the heat removal from the catalyst particle when the catalyst temperature is greater than 700 K. For the radiative heat transfer, as a simple approximation in view of the dilution of the catalyst with inert particles in the bed, the catalyst particle is assumed to be surrounded by inert particles at T f . Then, the heat exchange by radiation between the catalyst particle and its surrounding can be approximated by [34] Qrad a p (TS4 T f4 )
(11)
8 2 4 where is the emissivity of the catalyst particles, and is the Stefan–Boltzmann constant ( 5.67 10 W/m K ).
The heat transfer from the catalyst is the sum of the fluid–particle heat transfer and the radiative heat transfer. Q h p a p (TS T f ) a p (TS4 T f4 )
(12)
3.6.2. Heat generation in the catalyst particle Heat is generated by the combustion reaction in the catalyst. The reaction conditions inside the catalyst can be considerably different from the conditions at the outer surface of the catalyst. The O2 concentration inside the catalyst particle is lower than the O2 concentration at the outer surface, C A S , because O2 must diffuse inside the particle to react. Because the heat generated inside the catalyst should be transferred to the outer surface by heat conduction, the temperature inside the catalyst is higher than the temperature at the outer surface of the catalyst, TS .
The effects of the intraparticle diffusion and conduction should be accounted for in calculating the combustion rate of the catalyst.
The maximum possible temperature increase from the outer surface to the center of the catalyst particle can be estimated [27]. Tmax
H comDAeC AS ke
(13)
Here, DAe is the effective diffusivity of oxygen in the catalyst, k e is the effective thermal conductivity of the catalyst, and H com is the heat of combustion reaction for O2, which equals the heat of reaction of Eq. (4) divided by 1.5 (= 450 kJ/mol). DAe can be estimated by
D Ae
p DC
,
DC
1 m2/s 1 / DAm 1 / DKA
(14)
p (=0.47 [29]) is the porosity of the catalyst, and is the tortuosity factor of the catalyst, which was set to 3 [27,
35]. DC is the combined diffusivity of the bulk molecular diffusivity of oxygen in the reaction mixture, DAm , and the Knudsen diffusivity of oxygen, DKA . DKA 97 rP
TS m2/s MA
(15)
Here, rP is the mean pore radius of the catalyst in meters and M A is the molecular weight of the diffusing species ( M A 32 for oxygen). DAm can be estimated by the method in Reid et al.[36]. In this study, for example, at TS =
523 K, DAm 9.110 5 m2/s, DKA 3.1106 m2/s (for rP 8 109 m [29]), and, by Eq. (14), DAe 4.7 107 m2/s. The effective thermal conductivities of various porous catalysts are in the range of 0.16–0.64 W/(m K) [38]. A conservative estimate for Tmax by Eq. (13) with C AS 0.77 mol/m3 (3% O2 at 523 K and 1.1 atm) and ke =0.16 W/(m K) is only 1 K. Consequently, the temperature in the interior of the catalyst is considered to be uniform and the same as the temperature at the outer surfaces of the particles, TS .
The interior concentration profile is obtained as the solution of the mass balance equation in the catalyst. A dimensionless form of the mass balance for an isothermal catalyst particle is d2y dx
2
2 dy 2 yn 0 x dx
(16)
0,
(17)
The boundary conditions are dy dx
y(1) 1
x 0
The dimensionless variables are defined as x r / R p and y C A ( x) / C AS . is Thiele modulus, defined by
Rp
rA (C AS , TS ) p
(18)
C AS DAe
Here, R p is the radius of the catalyst particle and p is the catalyst density (2400 kg m-3 [29]). From the solution of Eqs. (16) and (17), the actual combustion rate of the catalyst particle is
rcom 4R p DAeC AS
dy dx
(19) x 1
If the concentration and temperature of the interior are the same as those of the external outer surface, the combustion rate of the particle is calculated as rA (CAS ,TS ) wp .
The ratio of the two rates ( rcom
/( rA (C AS , TS ) w p )) is the effectiveness factor, . From the solution y (x) of Eqs. (16) and (17), can be determined by [27]
3 dy 2 dx
(20) x 1
For n = 0.14, Eq. (16) is nonlinear and can only be solved numerically. An efficient numerical method for solving Eqs. (16) and (17) and calculating by Eq. (20) has been proposed by Kim and Lee [37]. When the desired accuracy for is not very high, an approximation formula for can be used. Kim and Lee also proposed such an approximation for , which has been shown to be accurate within 1% of the value obtained by Eq. (20) [38]. The approximation formula for the power law rate of Eq. (9) is
1
(21)
0.0633 2 0.0148 2 exp( 0.104 2 ) Eq. (21) is far simpler than Eq. (20) for the estimation of . The actual combustion rate of the catalyst particle can be obtained as
rcom rA (C AS , TS ) wp
(22)
O2 in the gas flow surrounding the particle is transferred to the outer surface of the particle, and the transfer rate is proportional to the concentration difference. The balance between the mass transfer and the reaction rate is
aP kc (C A0 C AS ) rA (C AS , TS ) w p
(23)
Here, k c is the mass transfer coefficient, and C A0 is the oxygen concentration of the gas mixture, which is the same as the O2 concentration in the feed since the catalyst particle is at the reactor inlet facing the feed.
If TS is given, Eq. (23) can be solved for C A S . Because Eq. (23) is nonlinear and implicit, an iterative numerical method is needed to find the solution. The range of C AS in Eq. (23) is [0, C A0 ]. With increasing C AS from 0 to CA0, the left-hand side of Eq. (23) decreases from a p kcC A0 to 0. On the other hand, the right-hand side of Eq. (23) increases from 0 to w p (rA (C A0 , TS )) . Hence, for a given TS , there is only one point in [0, C A0 ] where Eq. (23) is valid. A numerical method for finding a zero of a function such as the bisection method or Newton’s method can be used to find the solution of Eq. (23) [39]. Once C AS is determined for a given TS , the heat generation at TS is calculated using Eq. (24).
G (H com )[ rA (CAS , TS ) wp ]
(24)
3.6.3. Fluid–particle mass and heat transfer coefficients Several correlations of the mass and heat transfer coefficients have been developed [33, 40–48]. The correlations are given in terms of the Nusselt number (Nu = hp d p / k f ), the Sherwood number (Sh = kc d P / DAm ) or the Colburn jfactor (mass and heat transfer) as functions of the Reynolds number (Rep = vd P / ). The agreement among the correlations is poor, particularly at low Reynolds numbers (less than 10) where the estimated transfer coefficients
differ by several orders of magnitude [33, 41, 44, 47]. Because the calculated steady state depends strongly on the values of the transfer coefficients, h p and k c , widely different estimates of the steady states can be obtained depending on the choice of the correlations. Some correlations yielded simulation results incompatible with the experimental data. Among the available correlations, it was found that the mass transfer coefficient kc from the correlation proposed by Petrovic and Thodos [43] and the fluid–particle heat transfer coefficient h p from the correlation by Cybulski et al. [33] produced simulation results in agreement with the experimental data.
The mass transfer correlation by Petrovic and Thodos [43] is
b jM 0.357 Rep-0.359
(25)
where b is the catalyst bed porosity and j M is the Colburn j factor for mass transfer (= Sh/(Rep Sc1/3 ), where Sc =
/(DAm ) ). The heat transfer correlation by Cybulski et al. [33] is Nu = 0.07 Rep
(26)
In Table 1, the transfer coefficients and the physical properties of the feed at 523 K and 1.1 atm are listed. Although the reactor was operated at atmospheric pressure, the inlet pressure was 1.1 atm because of the pressure drop of the reactor. The physical properties were estimated by the methods in Reid et al. [36].
3.6.4. Multiple steady states of the catalyst particle In the steady state, the heat generation, G of Eq. (24), is equal to the heat removal, Q of Eq. (12), of the particles.
G (H com )[ rA (C AS , TS ) wp ] = h p a p (TS T f ) a p (TS4 T f4 ) Q
(27)
A solution method is to determine G and Q separately as a function of TS , and then to find TS values for G Q . In this method, Q(TS ) is simply given by Eq. (12), although the calculation of G(TS ) is somewhat involved. A method for calculating G for a given TS has been given in Section 3.6.2. For a range of TS , G(TS ) and Q(TS ) can be plotted, and TS values for G(TS ) = Q(TS ) can be found from the plots.
Fig. 8 shows the heat generation and removal curves for the catalyst particle at the reactor inlet surrounded by the feed stream at T f 523 K. In addition, the relative oxygen concentration C AS / C A0 with respect to TS is also shown. The heat removal line is slightly curved upward because of the radiative heat transfer from the catalyst particles. There are three steady states, denoted S1, S2, and S3 in Fig. 8, where the heat generation and removal coincide. The surface temperature and concentration ( TS , C AS / C A0 ) for the three steady states are (533 K, 0.99), (778 K, 0.66), and (939 K, 0.32) for S1, S2, and S3, respectively. The state S2 is unstable because any disturbance in the temperature causes the state to move to either S1 or S3. In contrast, S1 and S3 are the stable steady states and can be observed in the experiments. At state S1, C AS / C A0 is close to 1, so the state is limited by the reaction. At state S3, on the other hand, C AS / C A0 is small and the state is limited by the external mass transfer.
If T f increases or decreases, the heat removal Q(TS ) moves to the right or left, whereas G(TS ) remains fixed. If
T f decreases from 523 K, Q(TS ) moves to the left, and, when T f is lower than 502 K, the removal curve crosses the generation curve G(TS ) only once, giving a single steady state. This also happens when T f increases beyond 560 K. Thus, the multiple steady states are observed in a limited range of T f . Fig. 9 shows the computed multiple steady states of the catalyst particle with respect to T f . On increasing T f from 450 K, TS increases in the lower steady state until T f 560 K where TS jumps from 634 to 1011 K. For T f above 560 K, TS moves on the upper steady state. On decreasing T f from 560 K, TS decreases on the upper steady state until T f =502 K, where TS drops from 876 to 506 K and returns to the lower steady state. The simulation shows that for T f between 502 and 560 K, two steady states are observable for the catalyst particle.
3.7. Multiple steady states of the reactor The model and simulation for the multiple steady states discussed in Sections 3.6.1–3.6.4 are for a catalyst particle at the reactor inlet. In this section, we investigate how the multiple steady states of the particle at the reactor inlet affect the performance of the whole reactor.
Because of the high catalyst temperature of the upper steady state of the particle, the combustion rate of the upper steady states is very fast. The calculated combustion rate of the catalyst particle in the upper steady state was 3.2 108 to 4.3 108
mol-O2/s. The feed molar rate of oxygen was 2.23 106 mol/s. Because the reaction order
of the COM is low, the reaction may be approximated as a zero-order reaction with respect to the O2 partial pressure and the same combustion rate can be assumed for each catalyst particle in the region of the reactor where the COM takes place. Then, around 50–70 catalyst particles would be sufficient for the complete conversion of the input oxygen by the combustion reaction. Because the reactor contains around 1900 catalyst particles, the combustion is completed in a shallow region at the bed inlet and, in the remainder of the reactor, steam reforming takes place.
On the other hand, in the lower steady state, the combustion rate is slow, and, hence, all or a significant portion of the catalyst in the reactor is used for the combustion reaction. The oxygen conversion of the reactor calculated with Eq. (9) and the heat transfer coefficient of 111 W/m2 is found to be incomplete below a reactor temperature of 550 K.
Accordingly, the main reaction in the reactor depends on the state of the catalyst particle at the reactor inlet. When the catalyst particles are in the upper steady state, SRM is the main reaction and COM is completed at the inlet. When the catalyst particles are in the lower steady state, COM is the main reaction.
The methanol conversion of the OSRM reactor can be estimated from the SRM and COM kinetics. When COM is complete, the feed to the reactor (15 mL/min CH3OH, 25 mL/min H2O, 3 mL/min O2, and 57 mL/min He) becomes 13 mL/min CH3OH, 29 mL/min H2O, 2 mL/min CO2, and 57 mL/min He. This modified feed is used for the subsequent SRM in the reactor. The SRM kinetics for the Cu catalyst has been reported by Lee et al. [29].
rM
kK1 ( PM / PH 2 ) (1 K1 ( PM / PH 2 )(1 K 2 PH 2 )
mol/kg-cat/s
(24)
where k 3.37 1010 exp(
111 kJ/mol 20.1 kJ/mol ) , K1 3.54 103 exp( ) , and RT f RT f
K 2 4.18 10 7 exp(
51.3 kJ/mol ), RT f
PM is the methanol partial pressure (kPa), and PH 2 is the hydrogen partial pressure (kPa). The total rate of CH3OH consumed in the reactor is the sum of that consumed by COM at the reactor inlet and that consumed by SRM in the reactor. The total consumption rate is divided by the methanol feed rate before the COM reaction to yield the overall conversion of CH3OH throughout the reactor. The calculated overall conversion is shown as the red line in Fig. 10 and agrees well with the measured conversion of the third temperature cycle of the OSRM experiment. The dotted line, on the other hand, is the methanol conversion by combustion when all the catalyst particles are in the lower steady state. This also shows a good agreement with the experimental data of the lower steady state. Fig. 10 clearly shows that the steady state of the reactor depends on the state of the catalyst particle at the inlet of the reactor.
It has been suggested that the OSRM or POM rate is much faster than the SRM rate [1, 5, 7, 10, 12, 16]. As shown in this study, OSRM and POM are not single reactions but a series of COM reaction followed by SRM. Accordingly, the rates of OSRM and POM should be interpreted in terms of the rates of the individual reactions. Fig. 10 shows that the OSRM conversion of the reactor can be estimated from the COM and SRM rates.
The calculated temperature range for the multiple steady states does not agree well with the experimentally observed temperature range of 507–530 K shown in Fig. 1. The literature correlations for mass and heat transfer between the particle and gas flow assumed constant physical properties of the gas mixture at the particle surface and in the surrounding gas flow. This is not valid for the upper steady state because the temperature difference between the particle and the gas flow in the upper steady state is estimated to be more than 370 K. In Eqs. (12) and (14), symmetry in the concentration and temperature profiles, as seen from the center of the catalyst particle, is assumed. However, this may not be valid because the conditions at the front section of the particle facing the incoming flow and those at the aft section under the passing flow can be considerably different. The bulk temperature, T f , is assumed to be the same as the reactor wall temperature measured in the experiment. The simplifying assumptions could have attributed to the discrepancy between the experimental results and the simulation. Nevertheless, the
analysis clearly shows that the heat balance between the generation by COM and the loss by the external heat transfer gives rise to the multiple steady states.
The catalyst in the upper steady state is at a very high temperature, which is far above the normal operating temperature range (473–573 K) of Cu based catalysts [49] and will deactivate the catalyst quickly. This can explain the decrease in the methanol conversion as the temperature cycle is repeated, as shown in Fig. 1.
3.8. Hot spots To visually examine the catalyst at the upper steady state, the OSRM was carried out in a quartz tube reactor (OD: 6 mm, ID: 4 mm) in which 50 mg of the catalyst (0.25–0.3 mm) was diluted with 150 mg of the glass beads. For the reactor temperature of 573 K and 3% O2 (3 mL/min) in the feed, no visible hot spots were observed. However, when the O2 feed rate was increased to 9 mL/min (8.5% O2 in the feed), the catalyst inlet layer glowed red hot. This is shown in Fig. 11(a). The furnace controlling the reactor temperature was opened momentarily to take the pictures. The picture clearly shows that the combustion was instantaneous and completed at the inlet. After 4 and 12 h of operation, pictures of the reactor were taken again, as shown in Figs. 11(b) and 11(c). The glow at the reactor inlet dimmed with time, disappearing after 12 h. The pictures show the deactivation of the catalyst in the OSRM.
Because the combustion heat is used as the reaction heat in SRM, the heat should be transferred to the part of the catalyst bed where SRM takes place. In conventional packed bed reactors, the heat is mainly carried by the gas flow because of the poor thermal conductivity of the catalyst bed. Thus, the gas temperature increases sharply, and a hot spot is formed at the inlet. Even though the gas temperature is kept at the safe operating temperature of the catalyst, this study shows that the catalyst temperature can be very high due to the fluid–particle heat transfer resistance. If the catalyst is not resistant to such high temperatures, the catalyst deactivates rapidly, as shown in this study, and shortens the reactor life.
4.
Conclusions
The oxidative steam reforming of methanol (OSRM) has been investigated to identify its constituent independent reactions. The reaction between oxygen and methanol in OSRM was the combustion of methanol (COM). The partial oxidation of the methanol, which is commonly assumed in OSRM studies, has been experimentally shown to not take place in the OSRM. In the OSRM, the COM took place alone until a substantial amount of O2 had been consumed and, subsequently, the steam reforming of methanol (SRM) took place.
Multiple steady states of the reactor were observed. The steady-state multiplicity of the reactor was linked to that in the catalyst particles at the reactor inlet where the COM was occurring. The multiple steady states of the catalyst particle were estimated by solving the heat and mass transfer balance equations between the catalyst and the surrounding gas phase. When the catalyst particle was in the upper steady state, the catalyst temperature was estimated to be above 880 K, the COM was completed instantly at the reactor inlet, and most of the catalyst bed was used for SRM to produce H2. When the catalyst particle was in the lower steady state, the catalyst temperature was comparable to the reactor temperature and the rate of COM was slow, so the entire reactor was used for the COM and producing the combustion products, CO2 and H2O.
The Cu catalyst was deactivated in the OSRM because of the high catalyst temperature. The high temperature was due to the combined effect of the significant heat of combustion and the heat transfer resistance between the catalyst particles and the gas flow in packed bed reactors. As a result, the lifespans of the OSRM reactors packed with Cubased catalysts may not be sufficiently long to be useful for onboard or mobile H2 production.
Acknowledgements The authors thank the financial support of Research Institute of Advanced Energy Technology at Kyungpook National University.
Appendix The rate equation, Eq. (7), has three kinetic parameters: A0 , E A , and n . rA A0 exp(
E A n ) PA RTS
mol/(kg-cat s)
(7)
Because the O2 conversion corresponding to the methanol conversion in Fig. 6 was as high as 0.7, the continuous increase in the O2 conversion from the inlet to the exit of the reactor should be considered in analyzing the data. The equation for the O2 conversion, X , along the reactor length is
FA0
dX rA (TS , PA ) dW
(A.1)
In Eq. (A.1), it is assumed that the effect of intraparticle diffusion on the reaction rate is negligible, so the effectiveness factor is not included in the reaction rate because it is 1. After the parameters in the reaction rate are specified, this assumption will be tested and verified.
To obtain the conversion at the reactor exit, Eq. (A.1) is integrated with the initial condition at the reactor inlet ( W 0, X 0 ) to the reactor exit ( W WC ). Here, WC is the total catalyst weight in the reactor and, in this study, it
was 5 105 kg. The computed conversion at the exit, X com , is compared with the experimental conversion, Xexp, to find the best estimates of A0 , E A , and m . To calculate X com , however, the three parameters in Eq. (7) must be specified beforehand.
Hence X com depends on A0 , E A and n , and, if correct values are specified for the
parameters, the agreement between X com and X exp will be good.
An objective function for search of the best estimates of the parameters can be defined as S
1 N
N
( X
com, i
X exp,i )2 .
(A.2)
i 1
Here N is the number of experimental data. Because X com is a function of the three kinetic parameters ( A0 , E A , and n ), S is also a function of these parameters. The minimum of S was found using the Nelder–Mead simplex method, implemented by the function fminsearch in MATLAB.
The rate of heat generation is the product of the reaction rate and the heat of the reaction. Because the heat of combustion is very large, even a slow combustion rate can generate sufficient heat to increase the catalyst temperature significantly. The catalyst temperature, TS , can be estimated by using a heat transfer coefficient between the fluid and the particle. The balance equation for the heat generation and the heat removal is (HCOM )wp (rA ( PA ,TS )) hp a p (TS T0 ) ,
(A.3)
where wP is the weight of a single catalyst particle, aP is the external surface area of a catalyst particle, hP is the fluid particle heat transfer coefficient, and T0 is the temperature of the surrounding bulk fluid. When compared to Eq. (27), the radiative heat transfer from the catalyst is not considered in Eq. (A.3) since the catalyst temperature of the data in Fig. 6 was less than 573 K. The left-hand side of the equation is the heat generated by combustion in a catalyst particle, and the right-hand side of the equation is the heat transferred from the catalyst particle into the bulk fluid. For the catalyst particle to be in a steady state, the two sides must be equal. Eq. (A.3) can be rearranged as
TS T0
(H COM ) wP (rA ( PA , TS )) . hP aP
(A.4)
Because, for h p , TS T0 , in this case, the measured reactor temperature can be used as TS in calculating the reaction rate in Eq. (7). The combustion rate equation Eq. (8) was obtained for the case of hp .
When h p is not sufficiently large, the temperature difference becomes non-negligible and TS must be calculated by solving Eq. (A.4). Because of the exponential terms in the reaction rate, Eq. (A.4) is nonlinear and can only be solved numerically. Initially, TS T0 is assumed in the right-hand side (RHS) of Eq. (A.4), and a new TS is calculated using Eq. (A.4). This new TS is substituted into the RHS of Eq. (A.4) to obtain an improved TS . The iteration is repeated until the sequence of TS values converges to the required temperature accuracy. The kinetics of Eq. (9) has been obtained for hP 111 J/(m2 K s).
In Eq. (A.1), the effectiveness factor of the catalyst particles has been assumed as 1.0. With the obtained reaction rate of Eq. (9), the effectiveness factor of the catalyst particle is calculated by Eq. (21) for the data in Fig. 6. The computed factors are in the range of 0.96 – 1.0 and can be assumed as 1.0 in the search for the kinetic parameters.
Nomenclature A0
pre-exponential factor of the rate constant (mol kg-1 s-1 kPa-n)
aP
outer surface area of a catalyst particle (m2)
C AS
O2 concentration at the outer surfaces of the catalyst particles (mol m-3)
C A0
O2 concentration in the gas flow (mol m-3)
DAe
effective diffusivity of O2 in the catalyst (m2 s-1)
DAm
bulk diffusivity of O2 (m2 s-1)
DkA
Knudsen diffusivity of O2 (m2 s-1)
dP
catalyst diameter (m)
EA
activation energy (J mol-1)
G
heat generation rate (J s-1)
hP
fluid–particle heat transfer coefficient (W m-2 K-1 )
-1 H com enthalpy change in the combustion of methanol for 1 mole of O2 (J mol )
jM
Colburn j factor for mass transfer (= Sh Rep-1 Sc-1/3)
kc
mass transfer coefficient (m s-1)
ke
effective thermal conductivity in the catalyst (W m-1 K-1)
kf
thermal conductivity of the gas mixture (W m-1 K-1)
Nu
Nusselt number ( = hp d p / k f )
n
reaction order
PA
O2 partial pressure (kPa)
Q
heat removal rate (J s-1)
R
gas constant (8.314 J mol-1 K-1)
Rp
radius of the catalyst particle (m-1)
Rep
particle Reynolds number (= vd P / )
rA
O2 reaction rate (mol kg-1 s-1)
Sh
Sherwood number (= kc d P / DAm )
Tf
gas flow temperature (K)
TS
catalyst temperature (K)
T0
reactor temperature (K)
WC
total weight of the catalyst in the reactor (kg)
wp
weight of a catalyst particle (kg)
Greek symbols
Stefan–Boltzmann constant ( 5.67 108 Wm-2 K -4 )
emissivity of catalyst
b
catalyst bed porosity
p
catalyst porosity
effectiveness factor
gas mixture viscosity (Pa s)
gas mixture density (kg m-3)
p
catalyst density (kg m-3)
Thiele modulus, defined in Eq. (18)
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Table 1. Physical properties of the feed and the mass and heat transfer coefficients of the catalyst particles. Feed: 15% CH3OH, 25% H2O, 3% O2, and 57% He; 100 mL (STP)/min Catalyst particle diameter, d P = 0.27 × 10-3 m Viscosity,
=
2.26 × 10-5 kg/(m s)
Thermal conductivity, k f = 0.143 W/(m K) Bulk diffusivity of O2, D Am = 9.1 × 10-5 m2/s Particle Reynolds number, Rep = 3.0 Mass transfer coefficient, k c = 0.405 m/s Heat transfer coefficient, h p = 111 W/(m2 K)
Figure captions Fig. 1. Methanol conversion with respect to temperature programming from 453 to 583 K (up) and from 583 to 453 K (down) at a rate of 0.2 K/min. The results of the 1st and 5th temperature cycles are shown.
Fig. 2. TPR graphs of the catalyst in the lower steady state (red line), upper steady state (blue line), after air exposure at room temperature for 5 min (green dashed line), and after oxidation at 523 K (green dotted line).
Fig. 3. Net formation of H2O and O2 conversion through the reactor along the fifth temperature cycle in which the temperature was programmed from 453 to 583 K (up) and from 583 to 453 K (down) at a rate of 0.2 K/min.
Fig. 4. The net flow rate of CH3OH with respect to the net flow rate of O2 through the reactor.
Fig. 5. Methanol conversion with respect to time-on-stream after the oxidation treatment applied between the runs.
Fig. 6. Effect of the O2 partial pressure in the feed on the methanol conversion. Line: calculated conversion with the reaction rate (Eq. (9)) and heat transfer coefficient hP 111 W m-2 K -1.
Fig. 7. (a) Effect of methanol partial pressure on O2 conversion. (b) Effect of H2O partial pressure on O2 conversion.
Fig. 8. Heat generation and removal of the catalyst particle with respect to the catalyst temperature. The relative O 2 concentration, C AS / C A0 , is shown for comparison.
Fig. 9. Multiple steady states of the catalyst particle at the reactor inlet. Arrows indicate the direction of change in the catalyst temperature with respect to the gas temperature.
Fig. 10. Comparison of the calculated (line) and experimental (symbols) methanol conversion.
Fig. 11. The catalyst bed in the quartz reactor during the OSRM. The pictures were taken (a) at the start, (b) after 4 h, and (c) after 12 h of the reaction. The feed consisted of 9 mL/min O2, 15 mL/min CH3OH, 25 mL/min H2O, and 57 mL/min He. The feed entered the reactor from the left.
Fig. 1. Methanol conversion with respect to temperature programming from 453 to 583 K (up) and from 583 to 453 K (down) at a rate of 0.2 K/min. The results of the 1st and 5th temperature cycles are shown.
Fig. 2. TPR graphs of the catalyst in the lower steady state (red line), upper steady state (blue line), after air exposure at room temperature for 5 min (green dashed line) and after oxidation at 523 K (green dotted line).
Fig. 3. Net formation of H2O and O2 conversion through the reactor along the fifth temperature cycle in which the temperature was programmed from 453 to 583 K (up) and from 583 to 453 K (down) at a rate of 0.2 K/min.
Fig. 4. The net flow rate of CH3OH with respect to the net flow rate of O2 through the reactor.
Fig. 5. Methanol conversion with respect to time-on-stream after the oxidation treatment applied between the runs.
Fig. 6. Effect of the O2 partial pressure in the feed on the methanol conversion. Line: calculated conversion with the reaction rate (Eq. (9)) and heat transfer coefficient hP 111 W m-2 K-1.
Fig. 7. (a) Effect of methanol partial pressure on O2 conversion. (b) Effect of H2O partial pressure on O2 conversion.
Fig. 8. Heat generation and removal of the catalyst particle with respect to the catalyst temperature. The relative O2 concentration, C AS / C A0 , is shown for comparison.
Fig. 9. Multiple steady states of the catalyst particle at the reactor inlet. Arrows indicate the direction of change in the catalyst temperature with respect to the gas temperature.
Fig. 10. Comparison of the calculated (line) and experimental (symbols) methanol conversion.
(a) start
(b) after 4 h
Catalyst bed
Ceramic fiber plug (c) after 12 h
Fig. 11. The catalyst bed in the quartz reactor during the OSRM. The pictures were taken (a) at the start, (b) after 4 h, and (c) after 12 h of the reaction. The feed consisted of 9 mL/min O2, 15 mL/min CH3OH, 25 mL/min H2O and 57 mL/min He. The feed entered the reactor from the left.
Highlights
Multiple steady states were observed in oxidative steam reforming of methanol (OSRM).
Partial oxidation of methanol did not occur in OSRM.
The reaction between O2 and CH3OH in OSRM was combustion of methanol (COM).
Fluid-particle mass and heat transfer resistances caused the multiplicity.
The Cu-based catalyst was deactivated in OSRM.
Multiple steady states in oxidative steam reforming of methanol
Methanol conversion
1.0 T up T down SRM after COM COM
0.8 0.6 0.4 0.2 0.0 460
480
500
520
540
Reactor temperature (K) SRM: steam reforming of methanol COM: combustion of methanol
560
580
