Stable Steam Reforming of Ethanol in a Two-Zone Fluidized-Bed

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Stable Steam Reforming of Ethanol in a Two-Zone Fluidized-Bed Reactor L. Perez-Moreno, J. Soler, J. Herguido,* and M. Menendez Catalysis, Molecular Separations and Reactor Engineering Group (CREG), Aragon Institute for Engineering Research (I3A), University of Zaragoza, 50009 Zaragoza, Spain ABSTRACT: The oxidative steam reforming of ethanol in a two-zone fluidized-bed reactor (TZFBR) and in a conventional fluidized-bed reactor over a Ni/Al2O3 catalyst has been investigated. Catalyst deactivation has been studied for both contact modes. Coke generation has been verified by several techniques, and a stable performance was obtained in the TZFBR, where coke formation was counteracted with continuous catalyst regeneration. The effects of the main operating variables (steam/ethanol ratio S/E, oxygen/ethanol ratio O/E, temperature, and relative velocity with respect to the minimum fluidization velocity) have been studied. Both stable results and continuous operation without catalyst deactivation were achieved in the TZFBR in a wide range of S/E values, showing its viability for carrying out this process. High hydrogen selectivity with total conversion of ethanol was achieved even at low S/E and O/E ratios.

1. INTRODUCTION Nowadays, natural gas steam reforming, partial oxidation, and coal gasification are the main processes for hydrogen production. However, the use of fossil fuels to produce hydrogen involves large releases of carbon dioxide. Hydrogen production using biofuels such as ethanol, therefore, appears to be an interesting alternative because of their low toxicity, availability, ease of handling, and renewability. Hydrogen production from bioethanol is attractive because it is the most widely available biofuel and allows renewable production of hydrogen.1 Three processes can be used for producing hydrogen from ethanol: (i) steam reforming of ethanol (SRE), (ii) partial oxidation (POX), and (iii) oxidative steam reforming of ethanol (OSRE). The steam reforming process produces a high H2/COx ratio but with the disadvantage of its high endothermicity and the deactivation of catalyst due to coke deposition. Reaction (1) is strongly endothermic and produces only H2 and CO2 if the ethanol reacts in the most desirable way. C2 H5 OH þ 3H2 O f 2CO2 þ 6H2 ,

ΔH° ¼ 174 kJ mol1

1 3 C2 H5 OH f CO2 þ CH4 2 2 C2 H5 OH f CH4 þ CO þ H2 2

C2 H5 OH þ H2 O f 2CO þ 4H2

ð2Þ

C2 H5 OH þ H2 O f CH4 þ CO2 þ 2H2

ð3Þ

Other reactions that can also occur are ethanol dehydrogenation to acetaldehyde (eq 4), ethanol dehydration to ethylene (eq 5), ethanol decomposition to CO2 and CH4 (eq 6) or to CO, CH4, and H2 (eq 7). C2 H5 OH f CH3 CHO þ H2

ð4Þ

C2 H5 OH f C2 H4 þ H2 O

ð5Þ

r 2011 American Chemical Society

ð7Þ 3

García and Laborde and Vaduseva et al. have shown that an increase in the temperature produces an increase in the H2 and CO equilibrium concentrations and a decrease in CH4. On the other hand, POX is an exothermic reaction, but it yields a lower H2/COx ratio than SRE. OSRE is a combination of the endothermic SRE and the exothermic POX. Consequently, OSRE needs lower energy and results in an adequate H2/COx ratio without external consumption of energy, reducing the quantities of methane and coke produced and supplying high H2 and low CO yields under optimal operational conditions. In the OSRE process,4 the additional supply of oxygen together with the bioethanol and steam consumes a part of the ethanol in order to produce the necessary heat to maintain reaction (8) with a drop in the H2 yield. It also increases the risk of catalyst sintering. 1 O2 f 2CO2 þ 5H2 , 2 ΔH° ¼ 50:3 kJ mol1

C2 H5 OH þ 2H2 O þ

ð1Þ However, other products such as CO and CH4 are also formed during the process:

ð6Þ

ð8Þ

5

Gallucci et al. have studied a variant of the OSRE process: the autothermal reforming of ethanol using a fluidized-bed membrane reactor where no oxygen is in contact with the ethanol. In this process, part of the produced hydrogen is recovered by means of hydrogen permselective membranes, where it reacts with feed oxygen, obtaining the heat needed for the steam reforming. Special Issue: CAMURE 8 and ISMR 7 Received: August 31, 2011 Accepted: October 11, 2011 Revised: October 3, 2011 Published: October 11, 2011 8840

dx.doi.org/10.1021/ie201968u | Ind. Eng. Chem. Res. 2012, 51, 8840–8848

Industrial & Engineering Chemistry Research

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1.1. Coke Formation. Coke formation in the SRE is mainly due to ethylene produced in ethanol dehydration. In fact, acetaldehyde and ethylene are promoters of coke formation. Possible reactions of coke formation are

2CO ¼ CO2 þ C

ðBoudouard reactionÞ

CH4 ¼ 2H2 þ C

ðmethane decompositionÞ

ð9Þ ð10Þ

CO þ H2 ¼ C þ H2 O

ð11Þ

CO2 þ 2H2 ¼ C þ 2H2 O

ð12Þ

6

Mas et al. obtained the range of conditions for coke formation, assuming that coke was formed in graphitic form from the reactions (9)(12). The possibility of coke formation is higher at low S/E values. Hence, for an S/E value of 1, there is the possibility of coke formation for the entire range of temperatures studied, while for S/E = 4, coke formation is only possible for temperatures lower than 450 K. Several authors have reported coke production using a fixedbed reactor in OSRE. Srisiriwat et al.7 observed coke formation with Ni/Al2O3 catalysts of 2.08 mmol C gcat1 h1 working at 700 °C, S/E = 3, and O/E = 0.26. These authors also used NiCeAl, NiZrAl, and NiCeZrAl catalysts, with coke formation in these cases being 0.57, 0.39, and 0.28 mmol C gcat1 h1 respectively, lower than that for the Ni/Al2O3 catalyst. Biswas and Kunzru8 observed a coke formation rate of 3.3 mol C gcat1 h1 at 600 °C, S/E = 8, and O/E = 0.5 with Ni/CeO2ZrO2 catalysts. Frusteri et al.9 reported 0.25 mmol C gcat1 h1 over Ni/CeO2 at 650 °C, S/E = 8, and GHSV of 40 000 h1. Laosiripojana and Assabumrungrat10 proposed a possible mechanism of coke formation during the SRE. According to their results, at low temperatures coke formation occurs mainly because of the CO and CO2 hydrogenation forming water and carbon. At higher temperatures, coke formation is produced mainly because of the methane, ethane, or ethylene decomposition and the Boudouard reaction. 1.2. Catalysts. Several authors in the past decade have explored the SRE as a way of producing hydrogen from bioethanol.1113 The catalytic activity of different catalysts14 based on Rh, Pd, Pt, Ni, Co, and Cu has been studied, and reviews of the results have recently been published.1517 The low cost of Ni, its high activity to ethanol conversion, and high H2 selectivity, together with its high CC bond rupture activity, make it a good choice as an active phase for the SRE. However, as stated in section 1.1, the polymerization of ethylene produced from the dehydration of ethanol, the Boudouard reaction, and methane decomposition contributes to carbon deposition on Ni/Al2O3 catalysts.18 Thermodynamic analyses have demonstrated that if temperatures lower than 650 °C are used, a S/E molar ratio higher than 2 is required to avoid coke formation.6 In the case of the OSRE, these coke formation problems can be reduced using an appropriate O/E ratio. The most studied catalysts for OSRE have been Ni-based, although Co- and Rh-based catalysts have also been studied for both the OSRE and SRE. 1.3. Reactors. Many authors have studied the SRE using traditional reactors over supported metal catalysts. However, as stated before, this process is energetically demanding and commonly suffers from severe deactivation by coking and sintering. Possible solutions include (i) operating at a high S/E ratio, decreasing coke production, but this involves a significant increase in the

Figure 1. Diagram of the reactor configurations: (a) TZFBR; b) FBR.

energy requirements, or (ii) introducing an oxygen stream with a low O/E ratio in order to carry out the OSRE, but with a subsequent decrease in yields and increasing the risk of catalyst sintering. To solve these problems, considerable efforts are being made to develop more active, selective, and stable catalysts. Nevertheless, optimization of the reactor and contact mode between reactants can play an important role in the SRE process. In this way, a possible solution is to separate reaction and catalyst regeneration by using a two-zone fluidized-bed reactor19 (TZFBR). This reactor has been successfully tested in our laboratories for other processes such as the non-oxidative dehydrogenation of propane and butane over Cr2O3/Al2O320,21 and PtSn/MgAl2O4 catalysts.22 The TZFBR is a system designed to counteract catalyst deactivation. In this reactor (Figure 1a), one of the reactants (usually a hydrocarbon) is fed at an intermediate point of the fluidized bed, while a second stream (usually containing an oxidizing reactant) is fed at the bottom of the bed. In this way, two zones with different atmospheres are created in the bed: (1) a reaction zone in the upper part where the desired reaction takes place and coke is formed on the catalyst surface and (2) a regeneration zone in the lower part where the catalyst is regenerated by coke combustion. A continuous circulation of solids between both zones is caused by the bubbles, characteristic of gassolid fluidizedbed reactors. Under suitable operating conditions, a steady state is achieved, with most of the oxygen being consumed in the lower part of the bed. The application of the TZFBR and related reaction systems for several catalytic processes was discussed in a previous review.19 An additional advantage provided by this type of reactor is the improved safety because the reactants are not premixed and usually the oxygen concentration in the gas phase at the point where the hydrocarbon is fed is much lower than that at the entry point of a conventional fixed bed with the same total feed. This work is focused on the study of the feasibility of the use of the TZFBR in the OSRE to obtain hydrogen from ethanol over a Ni-based catalyst, avoiding the problems of catalyst deactivation even at low S/E ratios. The minimization of the S/E ratio in raw material is suitable from both an energetic and a practical point of view. Although the whole reactor acts like an OSRE process, the separation of two feeds should provide SRE behavior in the 8841

dx.doi.org/10.1021/ie201968u |Ind. Eng. Chem. Res. 2012, 51, 8840–8848


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reaction zone and simultaneous catalyst regeneration in the oxidation zone. S/E ratios between 0.3 and 3 and temperatures between 600 and 700 °C are used so that the reactor operates in the coke formation region according to the results obtained by Mas et al.6 In addition, experiments in the TZFBR and in a conventional fluidized-bed reactor (FBR) have been done in order to compare the performances of both types of reactors.

Table 1. BET Surface Area Values for the 1% Ni/Al2O3 Catalyst before and after Reaction in the TZFBR and FBR

2. EXPERIMENTAL SECTION

materials (ethanol and also water) taking into account that the efficiency in the use of water may be essential, e.g., in mobile applications.

2.1. Catalyst Preparation. The catalyst used for the SRE was synthesized through an incipient impregnation method, following steps similar to those described in the literature23 for other catalysts. Preliminary experiments with a commercial steam reforming catalyst with a high Ni content showed serious agglomeration problems in the fluidized bed, which suggested the use of a low Ni content. The nominal composition of the catalyst was 1 wt % Ni. A commercially available γ-Al2O3 powder (SCCa 150/200, Puralox Sasol Germany GMBH; SBET = 198 m2 g1; 100 < dp < 250 μm, and 0.60 cm3 g1 of pore volume) was used as the support material. This support was calcined at 950 °C in a muffle furnace for 1 h (heating rate = 1 °C min1), and the specific area was decreased to 137 m2 g1. The catalyst was prepared by incipient impregnation of the support material with an aqueous solution of the Ni precursor Ni(NO3)2 3 6H2O (99.999%; Sigma-Aldrich). The resulting product was dried in an oven at 120 °C for 24 h. Subsequently, the sample was calcined at 750 °C in a furnace for 15 h with a heating rate of 1 °C min1 (SBET = 133 m2 g1). After calcination, the solid was sieved between 100 and 250 μm. 2.2. Reaction. The experimental system consisted of a 2.8-cmi.d. FBR made of quartz, with a distributing porous plate also of quartz. Inside the reactor, a 0.4-cm-o.d. quartz tube was inserted in order to feed the ethanol at an intermediate point of the bed. The gaseous reagents were fed with mass flow controllers (model 5850 TR, Brooks Instruments). The water was fed through a HPLC pump (Shimadzu; model LC-10AT VP) to a vaporizer, where it was evaporated and heated at 150 °C, and the steam was then mixed with the oxygen gas stream. Ethanol was also fed through a HPLC pump (Shimadzu; model LC-10AT VP), and the pipes were heated to the inlet to the reactor. The gaseous products were analyzed online by a gas chromatograph thermal conductivity detector (Varian, model CP-3800). The liquid products were caught in a condenser placed in a bath with salt and ice and subsequently injected into the same chromatograph and analyzed by a gas chromatograph flame ionization detector. Two different configurations of the reactor were used in the experiments: (a) a TZFBR where the ethanol was introduced through the upper part at the center of the catalyst bed and the oxygen, steam, and nitrogen were introduced through the lower part (Figure 1a); (b) a conventional FBR where all of the reagents were cofed to the reactor at the lower part (Figure 1b). The mass load of the catalyst was 40 g in the TZFBR experiments and 20 g in the FBR ones. In the TZFBR configuration, the movable tube distributor was placed at half-height from the bottom to the surface of the fluidized bed, so there was roughly 20 g of catalyst in each fluidized-bed zone. The following parameters were defined to calculate the performance of the process: fractional conversion of ethanol (XEtOH), selectivity from ethanol to product i (Si), and global yield to hydrogen from ethanol plus water (YH2/(EtOH+H2O)). This last parameter was intended to evaluate the efficiency of both raw

sample

SBET (m2 g1)

1% Ni/Al2O3 catalyst fresh calcined

133

1% Ni/Al2O3 catalyst after 100 h in the TZFBR 1% Ni/Al2O3 catalyst after 40 h in the FBR

127 93

XEtOH ð%Þ ¼

Si ð%Þ ¼

in out 100ðFEtOH FEtOH Þ in FEtOH

100ni Fiout out in mi ðFEtOH FEtOH Þ

YH2 =ðEtOH þ H2 OÞ ¼

FHout2 in ð3 þ S=EÞFEtOH

% loss

4.3 29.7

ð13Þ

ð14Þ

ð15Þ

2.3. Catalyst Characterization. Catalyst samples after reaction in both FBR and TZFBR were taken when the reactor was cooled to room temperature under a flow of inert gas (N2) sufficient to fluidize the bed. BrunauerEmmettTeller (BET) specific surface areas (SBET) of fresh and used samples in both types of reactors were obtained by static N2 adsorption measurements using TriStar 3000 (V6.08 A) equipment on previously degassed samples. The degasification was carried out in two stages: the first at a heating rate of 10 °C min1 until 90 °C, keeping this temperature during 1 h, and the second at a heating rate of 10 °C min1 until 200 °C, maintaining that temperature for 8 h. Thermogravimetric analyses (TGA) were made using TGA/ SDTA 851eSF/1100 °C (Mettler Toledo) equipment with an air flow of 50 mL (STP) min1 in two stages: heating at a rate of 10 °C min1 up to 1000 °C and then maintaining the temperature for 30 min. The surface morphology of fresh and used catalysts was analyzed by transmission electron microscopy (TEM) on a JEOL 2000 FXII (200 kV) unit. The sample was dispersed in ethanol with ultrasound for several minutes and then placed in a copper grill. The temperature-programmed reduction (TPR) experiments were carried out using AUTOCHEMII equipment (Micromeritics). The test conditions were as follows: (1) drying of the sample with an argon flow of 50 N mL min1 until 110 °C (heating rate 10 °C min1) and maintaining the temperature for 30 min; (2) reduction of the sample with a 10% H2/Ar flow until 950 °C (heating rate 5 °C min1).

3. RESULTS AND DISCUSSION 3.1. Catalyst Characterization. The SBET values of the different samples are summarized in Table 1. The BET specific area of the catalyst after 100 h in the TZFBR was very similar to that of the fresh calcined catalyst. This indicates that no textural changes were produced in the catalyst resulting from reaction in the TZFBR. In the case of the catalyst used in the FBR for 40 h, a significant drop in the BET surface area was observed; this can be due to a high degree of coke formation on the catalyst. 8842


Industrial & Engineering Chemistry Research The coke deposition on the catalyst surface after reaction in the FBR was corroborated with the TGA of the samples. Figure 2 shows the results of TGA experiments for the fresh sample and the catalyst after reaction for 12 h in both types of reactors,

Figure 2. TGA of the fresh calcined catalyst and the catalyst after reaction in the TZFBR and FBR.

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respectively. It can be observed that the sample after 12 h in the TZFBR presents a behavior very similar to that of the fresh one, which indicates that there is negligible coke presence on it, while the sample used in the FBR presents a high loss of mass (around 23%) because of the oxidation of the coke formed during the reaction and deposited on it. TEM images of samples of the fresh calcined catalyst and the catalyst after reaction in the TZFBR and FBR are presented. The fresh catalyst (Figure 3a) has the typical structure of alumina and the Ni particles dispersed in this alumina. After use in the TZFBR (Figure 3b), the structure of the alumina is maintained and the presence of coke is not observed, neither in nanofibers nor in agglomerated form. Its appearance is very similar to that of the fresh catalyst; i.e., no significant changes took place in the sample with this contact mode. On the other hand, a huge formation of coke was observed in the catalyst after use in the FBR, both as nanofibers and as agglomerates. Although there is a predominance of nanofiber formation, as can be seen in Figure 3c, agglomerated coke can also be seen in Figure 3d. This large coke formation is consistent with the TGA results, explaining also the decrease in the BET surface area in the catalyst after use in the FBR. The TPR analyses of the fresh and used catalysts are shown in Figure 4. The fresh catalyst has mainly one peak at around

Figure 3. TEM images of fresh and used 1% Ni/Al2O3 catalyst: (a) fresh calcined; (b) used in the TZFBR; (c and d) used in the FBR. 8843



Figure 4. TPR studies of fresh and used 1% Ni/Al2O3 catalyst.

900 °C, which corresponds to the NiAl2O4 phase. Many authors using Ni/Al2O3 catalysts have attributed the peak at higher temperatures to the spinel phase, although depending on the calcination temperature and the quantity of Ni in the catalyst, the temperature of this peak can differ remarkably. Alberton et al.24 worked with a Ni/γ-Al2O3 catalyst and obtained a peak at high temperatures (740 °C for a catalyst with 16% Ni content and 850 °C for a catalyst with 8% Ni content) that was attributed to the spinel phase NiAl2O4. The peak of the spinel does not appear in either of the used catalysts, but there are peaks that correspond to the reduction of NiO to Ni. The first peak at about 250 300 °C would correspond with the bulk nickel oxide that weakly interacts with the support. The second one, at around 500600 °C, could be due to the NiO well dispersed on Al2O3. This result is in keeping with the fact that supported Ni catalysts show different reduction behaviors depending on the nature of the Nisupport interaction. So, bulk nickel oxide particles that weakly interact with the support are reduced at around 350400 °C, but when Ni is highly dispersed on Al2O3, the metalsupport interaction decreases the reducibility of the Ni ion to Ni0 and its reduction peak appears at above 500 °C.25 The main difference between the two used catalysts is that the intensity of the second peak is greater in the case of the catalyst after its use in the FBR. This could mean that this catalyst is more oxidized. Because the active phase in the reforming is Ni, the most reduced catalyst is the most active. 3.2. Study of the Stability of the System. Stability is an important characteristic of any catalyst, determining its efficient use in a reaction. In fact, stability is one of the most important questions from an industrial point of view. Deactivation processes cause a reduction in activity. Several authors710 have reported deactivation of the catalyst due to coke formation in the OSRE process, as explained in section 1.1. In this context, the TZFBR could present several advantages with regard to the cofeeding of the reactants (FBR). A study of the OSRE over a 1 wt % Ni/Al2O3 catalyst was carried out in the TZFBR and in a conventional FBR for comparison purposes. With the TZFBR, in all cases stable results and continuous operation without net catalyst deactivation were observed, showing the viability of this reactor for carrying out this process. The temporal evolution of ethanol conversion for one of these cases, working with the same gas streams but with different feeding configurations (Figure 1), is shown in Figure 5. It can be observed that operation with the FBR exhibited a sharp deactivation

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Figure 5. Temporal evolution of the ethanol conversion in the TZFBR and FBR using a 1% Ni/Al2O3 catalyst (lines for visual help). Operating conditions: T = 650 °C, u = 3umf (reaction zone), Wcat = 20 g zone1, O/E = 0.2, S/E = 1.5, and N2/E = 0.3.

Figure 6. Product distribution (as Si values) from both types of reactors over a 1% Ni/Al2O3 catalyst. Operating conditions: T = 650 °C, u = 3umf (reaction zone), Wcat = 20 g zone1, O/E = 0.2, S/E = 1.5, and N2/E = 0.3. *Coke is calculated by carbon mass balance closure.

but the TZFBR maintained a stable performance, in spite of the low S/E ratio (1.5; i.e., 50% of the stoichiometric ratio for ESR according to eq 1) and also with a low O/E ratio (0.2). It is indicative that equilibrium between coke formation (in the upper zone of the reactor) and combustion (in the lower zone of the reactor) is reached in this TZFBR configuration. In the case of FBR, deactivation of the catalyst due to coke formation was corroborated by TGA, and a drop in the BET area was also observed, as stated in the previous section. The distribution of products was also influenced by the configuration of the reactor, as can be seen in Figure 6. The selectivity to hydrogen from ethanol obtained with the TZFBR is much higher than that with the FBR. This can be explained by the fact that, in the TZFBR, oxygen is consumed in the lower part of the reactor to burn the coke. Therefore, as ethanol is fed into the middle of the bed, the hydrogen obtained as the product in the upper zone is not in contact with the oxidant to produce water. Moreover, in the FBR, the parallel reaction of ethanol dehydration is more favored probably because of the low activity of the catalyst, which is deactivated by coke in the FBR, while the TZFBR 8844



Figure 7. Effect of the relative gas velocity on the product selectivity in a TZFBR with 1% Ni/Al2O3 catalyst (lines for visual help). Operating conditions: t = 190 min, T = 650 °C, u = 2umf (regeneration zone), and Wcat = 20 g zone1.

allows good activity to be maintained, diminishing the selectivity to the intermediate products (SC2H4 and SC2H6). In the TZFBR, the use of a small stream of oxygen (low O/E ratio) fed to the lower part of the fluidized bed together with some inert gas was enough to continuously remove the coke that was formed in the ethanol reaction zone. In this way, a steady state was achieved where the rate of coke formation in the upper zone of the reactor was equal to the rate of coke removal by oxidation in the lower part of the reactor. Moreover, the process works in a mode different from that of the conventional OSRE because in the reaction zone no oxygen, or only a small part of the inlet stream, is present in the gas phase. This allows work in the reaction zone in a pseudo-SRE mode but reduced S/E ratio work in an autothermal mode and in situ regeneration of the catalyst by coke combustion, allowing a stable process. 3.3. Experiments in the TZFBR. In the TZFBR configuration, when the reactants were fed separately (ethanol at the middle of the bed and oxygen, nitrogen, and steam at the bottom, respectively), the effects of the main operating variables were studied in order to obtain the optimal conditions: temperature (600700 °C), S/E ratio (0.33.0), O/E ratio (00.3), and relative gas velocity ur = u/umf (2.04.0). Ethanol conversion was 100% in all of the experiments. 3.3.1. Effect of the Relative Velocity. The effect of the relative velocity in the upper zone is shown in Figure 7. H2 and CO selectivities increased with the contact time (smaller values of the space velocity in the reaction zone). This result corresponds with lower ethylene selectivity, an intermediate product of ethanol dehydration. In view of these results, it was decided to study the rest of the variables with a relative velocity equal to 3 (i.e., the maximum ur at which SC2H4 and SC2H6 are null). With an increase in the relative velocity, the ethanol dehydration acquires importance as an ethanol-consuming reaction to the detriment of the ethanol decomposition and ethanol-reforming reactions. These two reactions consume ethanol to produce the products of interest in the OSRE (H2 and CO2) and also CO and CH4, whose selectivities decrease with an increase in the relative velocity. 3.3.2. Effect of the O/E Ratio. The effect of the O/E ratio is shown in Figure 8. As expected, a greater value of H2 selectivity

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Figure 8. Effect of the O/E ratio on the product selectivity in the TZFBR with a 1% Ni/Al2O3 catalyst (lines for visual help). Operating conditions: t = 190 min, u = 3umf (reaction zone), T = 650 °C, Wcat = 20 g zone1, S/E = 1.5, and N2/E = 0.5 (O/E).

was obtained in working without oxygen, but coke selectivity after the selected time-on-stream (190 min) was in this case 18%, while in the rest of the experiments there was no net coke formation. On the other hand, working with an O/E ratio = 0.1, a smaller H2 selectivity was obtained and the production of ethylene and ethane was favored at the expense of a smaller production of CO2. When the O/E ratio was slightly increased to 0.2, ethanol dehydration was not favored. In fact, ethylene was not produced, having high H2 and CO2 selectivities in comparison with the other cases. However, if the O/E ratio was increased to 0.3, there was a slight decrease in the H2 selectivity, which suggests that if more oxygen than that necessary for regeneration in the lower zone is used, the excess is employed in the burning of hydrogen. These results are in good agreement with those of Peela and Kunzru.26 These authors have also observed for OSRE a decrease in the selectivity to H2, CO, and CH4 with an increase in the O/E ratio due to the oxidation reactions of these products in order to form CO2 and H2O. With an increase in the O/E ratio, the selectivity to some undesired products decreases but H2 selectivity is adversely affected. 3.3.3. Effect of the Temperature. The effect of the temperature is shown in Figure 9. At 600 °C, ethanol dehydration was favored. For this reason, the selectivity to ethylene was 30% at this temperature, which implied a decrease in the H2 and CO production. With an increase of the temperature to 650 °C, a significant modification of the product distribution was observed, with the production of ethylene and ethane disappearing and the H2, CO, and CH4 selectivity increasing. An increase in the temperature to 700 °C decreased the CO2 selectivity considerably in favor of an increase in the CO selectivity, maintaining the H2 selectivity. According to these results, an increase in the reaction temperature improves the selectivity to H2 and CO, which suggests that the two main reactions at high temperature are ethanol decomposition and ethanol reforming. Ethylene was only detected at the lowest temperature; i.e., ethanol dehydration was more important than the decomposition and reforming reactions. The increase in the temperature involves an increase in the CO selectivity and a decrease in the CO2 selectivity, which suggests 8845



Figure 9. Effect of the temperature on the product selectivity in the TZFBR with a 1% Ni/Al2O3 catalyst (lines for visual help). Operating conditions: t = 190 min, u = 3umf (reaction zone), Wcat= 20 g zone1, O/E = 0.2, S/E = 1.5, and N2/E = 0.3.

Figure 10. Effect of the S/E ratio on the product selectivity in the TZFBR with a 1% Ni/Al2O3 catalyst (lines for visual help). Operating conditions: t = 190 min, T = 650 °C, u = 3umf (reaction zone), Wcat = 20 g zone1, O/E = 0.2, and N2/E = 1.8 (S/E).

that at higher temperatures the reverse watergas shift reaction is favored. 3.3.4. Effect of the S/E Ratio. This variable is important for the OSRE in the TZFBR because the set objectives in this process are to work at S/E ratios as low as possible, achieving zero net coke formation. The effect of the S/E ratio is shown in Figure 10. It is observed that an increase in the S/E ratio favors the products of SRE (H2 and CO2) while decreasing ethylene and ethane. In fact, these byproducts are zero for S/E equal to or higher than 1.2. When the S/E ratio is increased from 1.5, there is a decrease in the selectivity to CO and CH4. On the basis of the results obtained, for low S/E ratios, the reactions that consume ethanol are not only the SRE and ethanol decomposition, where CO, CH4, H2, and CO2 were formed, but also ethanol dehydration (for example, ethylene selectivity is 30% at S/E = 0.3). For higher S/E ratios, the SRE and decomposition

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Figure 11. Comparison between the results from the TZFBR obtained in this work and those found in the literature.

were favored, with a corresponding decrease in ethylene production. This could be explained by ethanol dehydration being disfavored in the presence of water or by the reforming of the formed ethylene. From S/E = 1.2, the C2 products disappeared and CO and CH4 selectivities decreased, maintaining the upward trend in H2 and CO2. This performance is related to a higher ethanol consumption for those S/E values by the SRE, which produces H2, CO2, and CH4, instead of ethanol decomposition. It is noteworthy that S/E ratios lower than the stoichiometric ratio for the SRE have been used. Comas et al.27 also observed that H2 selectivity increased with an increase in the S/E ratio. In their case, it was necessary to work with S/E values between 1 and 6, while in this work, S/E values between 0.3 and 3 have been used, achieving better results without working with such high S/E values. Comas et al.27 worked in a fixed-bed reactor, obtaining a selectivity to coke of 30% at S/E = 1 and 15% at S/E = 3.3; while in this study, there is no net coke formation in the TZFBR. In summary, when working in a TZFBR under suitable conditions such as those used in the runs shown in Figure 10 and with a low S/E ratio (1.5), a good selectivity to H2 is obtained, maintaining a stable operation without a net coke formation. 3.4. Comparison with the Literature. A comparison of the results obtained in this work for the OSRE in the TZFBR using a 1 wt % Ni/Al2O3 catalyst and those reported in the literature with different catalysts and experimental systems is shown in Figure 11. The global H2 yield from ethanol plus water, calculated according to eq 15, is compared in this figure as a function of the O/E ratio. The majority of studies have used fixed beds with different catalysts, mainly Ni-based: Srisiriwat et al.7 used Ni/Al2O3 catalysts without and with promoters of CeO2, ZrO2, and CeO2ZrO2; Youn et al.28 used Ni catalysts supported in stabilized metal oxide; Comas et al.27 used Ni/Al2O3 catalysts; Biswas and Kunzru8 used NiCeO2ZrO2 catalysts; Li et al.29 used Ni/ZrO2 catalysts; Liberatori et al.30 used Ni/Al2O3 catalysts with the addition of La or Ag. Other authors have used a microreactor as a reaction system, for example, Peela et al.26 using Rh-based catalysts and Huang et al.31 using Co/Al2O3 and Co/Fe/Al2O3 catalysts in a fixed-bed microreactor. For most of the cases represented in Figure 11, deactivation of the catalyst due to coke formation was reported, while with the TZFBR, a continuous and stable system was achieved, allowing at 8846


Industrial & Engineering Chemistry Research the same time one to obtain a high H2 yield in working with low O/E ratios compared with those in the literature. Moreover, the operation was carried out with low S/E ratios (0.33), in most of the experiments lower than the stoichiometric ratio, while in most of the cases reported in the literature, S/E ratios equal to or higher than the stoichiometric ratio were used (38). At low O/E ratios, the system is more unstable and is not autothermic, and the coke formation is easier. For this reason, it is observed that in using the TZFBR we have achieved stable operation while also working in the region of instability, except for a ratio O/E = 0, as stated in section 3.3.2. Besides, the values obtained for the global H2 yield are much higher than those obtained by other authors working with much higher O/E ratios.

4. CONCLUSIONS Stable results and continuous operation within the studied range of S/E values without net catalyst deactivation have been observed when using the TZFBR, showing the viability of this reactor for the OSRE process. Under the same conditions, deactivation due to coke formation in the FBR was observed by TGA. The TEM images also show the presence of coke in the catalyst after its use in the FBR. There is a significant drop in the BET surface area in the catalyst after use in the FBR compared with that used in the TZFBR. The H2 selectivity is clearly affected by the operating conditions. A relative velocity (ur) of higher than 2 causes a selectivity decrease. The optimum operation temperature was found in the range 650700 °C. The increase in the S/E ratio favors the products of SRE (H2 and CO2) while decreasing ethylene and ethane. S/E ratios lower than the stoichiometric ratio for the SRE have been used, obtaining a H2 selectivity of up to 73% and a CO2 selectivity of 50%. Better results than those reported in the literature for conventional reactors have been obtained in work with the TZFBR with S/E ratios smaller than the stoichiometric ratio and with a higher formation of reforming products (H2 and CO2) without the formation of byproducts (ethylene and ethane). A stable operation has been achieved without net coke deposition in spite of work in the coke formation region. In fact, the values of the global H2 yield from ethanol plus water obtained with the TZFBR are between 0.6 and 0.9. ’ AUTHOR INFORMATION Corresponding Author

*Tel: +349762393. Fax: +349761879. E-mail: [email protected].

’ ACKNOWLEDGMENT This work has been partially funded by the Ministry of Education and Science (Projects CTQ 2007-63420-PPQ and CTQ-2010-15568). L.P.-M. is thankful for the grant received from the Government of Aragon. ’ NOMENCLATURE Symbols

FBR Fi GHSV mi

conventional fluidized-bed reactor molar flow of product i (mmol min1) gas hourly space velocity (h1) number of C atoms (or of H atoms when i is H2) in a molecule of ethanol

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N2/E ni

inert gas to ethanol molar feed ratio number of C atoms (or of H atoms when i is H2) in a molecule of product i O/E oxygen to ethanol molar feed ratio OSRE oxidative steam reforming of ethanol POX partial oxidation S/E steam to ethanol molar feed ratio selectivity from ethanol to product i as defined in eq 14 Si (%) SRE steam reforming of ethanol T temperature in the fluidized bed (°C) t time on stream (min) TZFBR two-zone fluidized-bed reactor u gas velocity in the bed (cm min1) minimum fluidization velocity (cm min1) umf relative gas velocity defined as u/umf ur Wcat mass load of catalyst in the bed (g) XEtOH fractional conversion of ethanol as defined in eq 13 (%) YH2/(EtOH+H2O) global yield to H2 as defined in eq 15 ΔH° standard enthalpy of reaction (kJ mol1) Superscripts

in out

at the reactor entrance at the reactor exit

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Stable Steam Reforming of Ethanol in a Two-Zone Fluidized-Bed

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