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Stable hydrogen production by methane steam reforming in a two-zone fluidized-bed reactor: Effect of the operating variables
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 7 8 3 0 e7 8 3 8
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Stable hydrogen production by methane steam reforming in a two-zone fluidized-bed reactor: Effect of the operating variables L. Pe´rez-Moreno, J. Soler, J. Herguido*, M. Mene´ndez Catalysis, Molecular Separations and Reactor Engineering Group (CREG), Arago´n Institute of Engineering Research (I3A), University of Zaragoza, 50009 Zaragoza, Spain
article info
abstract
Article history:
The oxidative steam reforming of methane in a two-zone fluidized-bed reactor (TZFBR) was
Received 4 March 2013
investigated over a Ni/Al2O3 catalyst. The effects of the main operating variables (tem-
Received in revised form
perature, steam/oxygen ratio, steam/methane ratio and relative velocity with respect to
19 April 2013
the minimum fluidization velocity) have been studied. A comparison has been made with
Accepted 21 April 2013
results given in the literature in terms of hydrogen yield. Despite working with very low
Available online 18 May 2013
steam/methane ratios, high values of hydrogen yield at both high methane conversion and at steady state were obtained in the TZFBR.
Keywords:
Copyright ª 2013, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights
Oxidative steam reforming
reserved.
Methane Ni/Al2O3 catalyst TZFBR
1.
Introduction
Methane is the main component of natural gas (NG) and landfill gas as well as being a by-product of oil refining and chemical processing. It has enormous potential value as a source of clean fossil energy. Natural gas reservoirs are large and widespread throughout the world, mainly in the Middle East and Russia, although other areas also have their fair share. The estimated reserves at the end of 2006 amounted to approximately 6300 trillion cubic feet [1]. In recent years, hydrogen has attracted great interest as a clean fuel for combustion engines and fuel cells [2]. Among all the potential sources for obtaining hydrogen, natural gas, which has methane as its main component, is considered a good option because it is clean, abundant and can be easily converted into hydrogen [3].
Steam reforming of methane (SRM) is the main industrial route for obtaining hydrogen and synthesis gas [4]. About 50% of the worldwide hydrogen demand is satisfied by steam reforming of methane. However, in this process there are two main problems: a) the reaction is endothermic and requires external heating; b) nickel based catalysts require the use of high steam/methane ratios (3e3.5) in order to avoid the formation of coke that causes catalyst deactivation [5]. In order to solve these problems, several alternatives have been studied. The reduction of the steam/methane ratio decreases the heat requirements although it may increase coke deposition. The introduction of oxygen with the feed adds exothermic reactions such as combustion or partial oxidation. This process is called oxidative steam reforming of methane (OSRM). It avoids some of the problems found in conventional steam methane reforming but has some drawbacks.
* Corresponding author. Tel.: þ34 976762393; fax: þ34 976762142. E-mail address: [email protected] (J. Herguido). 0360-3199/$ e see front matter Copyright ª 2013, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.ijhydene.2013.04.122
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CH4 þ 2O2 /CO2 þ 2H2 O 1 CH4 þ O2 /CO þ 2H2 2
DH0298 ¼ 803 kJ=mol
DH0298 ¼ 36 kJ=mol
(1) (2) (3)
Other reactions that can also occur are the water gas shift reaction (WGS) (Eq. (4)) and methane dry reforming (Eq. (5)) CO þ H2 O4CO2 þ H2 CH4 þ CO2 /2CO þ 2H2
DH0298 ¼ 41 kJ=mol DH0298 ¼ 247 kJ=mol
(4) (5)
Although oxidative steam methane reforming avoids the need for the large heat supply that characterizes conventional steam methane reforming, this process presents several drawbacks: 1) the formation of hot spots in the initial part of the bed [6]; 2) low activity of the nickel based catalysts due to the oxidation of the nickel metallic species; 3) selectivity loss; and 4) coke formation in the final part of the bed due to the lack of oxygen [7]. The use of Ni-supported catalysts in oxidative steam methane reforming has been extensively studied [3,8e11]. It has been reported that Rh and Pd based catalysts are effective for eliminating the formation of hot spots [12,13] whereas Nibased catalysts tend to form them [11,12,14]. However, noble metals such as Pt, Pd and Rh are not suitable components due to their limited availability and their high price. For this reason, two strategies can be proposed: a) the development of Ni-based catalysts modified with low quantities of noble metals [15] or b) the study of processes with Ni-based catalysts but using different types of reactor in order to minimize the problems. The high temperatures associated with methane steam reforming also favour coke formation. The deactivation of nickel catalysts by carbon formation is a significant problem in methane reforming because of Ni surface fouling, catalyst pore blockage and disintegration of the support material [16]. Thermodynamically, the most probable reactions for carbon formation are the following: CH4 4C þ 2H2 ðCH4 crackingÞ 2CO4C þ CO2 ðBoudouardÞ
DH0298 DH0298
¼ 122:3 kJ=mol
(6)
¼ 125:2 kJ=mol
(7)
CO þ H2 4C þ H2 OðCO reductionÞ DH0298 ¼ 84:0 kJ=mol
1.1.
100 98
Methane conversion (%)
DH0298 ¼ 206 kJ=mol
96 94
C H 4 /H 2 O /O 2
92
2/1.25/1.2 5 2/1.5/1 2/1.67/0.8 3 2/1.875 /0 .625 2/2/0.5
90 88
90
80
SH2 (%)
CH4 þ H2 O/CO þ 3H2
increased because of the preferential formation of CO at high temperatures. This fact is consistent with the WGS equilibrium (Eq. (4)) because CO2 generation is favoured at low temperatures due to the exothermicity of the reaction, while the reverse reaction (RWGS) yielding CO is favoured at high temperatures. At 900 C, total conversion of methane can be achieved, yielding syngas with a composition close to that given by Eq. (1) (75% H2 þ 25% CO). Methane and steam conversions are very high but not total, even at 700 C. The hydrogen expected in dry gases amounts to 70 vol.%. In our study, a simulation of the thermodynamic equilibrium was done using the Hysys v7.0 Aspen Tech program with the Gibbs reactor option. Methane conversion and selectivity to the main products of the process have been studied in relation to the operating variables: temperature, steam/oxygen ratio and steam/methane ratio. In Fig. 1 the variation of methane conversion and selectivities ðSH2 and SH2 =SCO Þ with the temperature are shown. The methane conversion increased with the temperature, this
70
60
1.1
SH2/SCO
In oxidative steam methane reforming, oxygen must be provided to the reactor together with the steam in order to combine steam methane reforming (endothermic, Eq. (1)) with two exothermic reactions: methane combustion (Eq. (2)) and the partial oxidation of methane (Eq. (3)). In this way the global process may be autothermal.
1.0
0.9
(8)
Thermodynamic equilibrium
Bion et al. (2010) studied the effect of temperature on the gas composition at thermodynamic equilibrium for methane steam reforming working with H2O/CH4 ¼ 1 [17]. They observed maximum hydrogen production at around 700 C. Above this temperature, the H2 concentration no longer
0.8 700
750
800
850
900
T (ºC)
Fig. 1 e Methane conversion and CO and H2 selectivities in thermodynamic equilibrium for several CH4/H2O/O2 ratios as a function of temperature.
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increase being greater for higher steam contents in the reacting gas. With regard to CO selectivity (not shown in Fig. 1), there was a similar increase with the temperature irrespective of the methane/steam/oxygen inlet composition. In the case of H2 selectivity, when the temperature was increased the selectivity slightly decreased for all the methane/steam/oxygen ratios. These results can be attributed to both the endothermic character of the main steam reforming reaction (Eq. (1)) that favours the conversion and the CO and H2 generation, and the exothermal character of the water gas shift reaction (Eq. (4)) that at high temperatures favours CO production at the expense of H2 (i.e., reverse water gas shift reaction RWGS). The SH2 =SCO ratio in equilibrium conditions had a value near to one (i.e., a gas composition H2/ CO ratio of around 3 according to Eq. (1)) for the range of temperatures considered. However, the lower the steam content in the reacting gas, the lower was the SH2 =SCO ratio at a given temperature. Moreover, a decrease with temperature was observed in the SH2 =SCO value at a given gas inlet composition. The effect of the steam/oxygen ratio at constant temperature and constant CH4/(H2O þ O2) ratio has been analysed. A sharp increase in CO selectivity was observed when the steam/oxygen ratio was increased, while methane conversion was almost unaffected by this variable. H2 selectivity greatly increased with the steam/oxygen ratio. This behaviour was due to the greater yield of the methane reforming reaction (Eq. (1)) at a high steam/oxygen ratio. The effect of the steam/methane ratio while maintaining constant both the CH4/O2 ratio and the temperature has also been analysed. A sharp decrease in the CO selectivity was observed, while the methane conversion was not affected by this variable. The H2 selectivity increased with the steam/ methane ratio. It seems clear that an excess of H2O favoured the WGS reaction (Eq. (4)).
1.2.
Reactors
The use of the two-zone fluidized-bed reactor (TZFBR) could serve to overcome the problems of oxidative steam methane reforming. Selectivity loss is not expected in this reactor because regeneration and reaction take place in the same vessel. The circulation of the solid avoids the formation of hot spots. At the same time, the catalyst regeneration will produce heat that can be used in steam reforming, mitigating in some degree the process endothermicity. Moreover, we aim to reduce the steam/methane ratio without having problems with coke deposition. This type of reactor consists of a fluidized bed where an oxygenesteam mixture is fed through the lower part of the reactor while methane is introduced at an intermediate point of the bed [18]. In this way, two zones are created in the reactor. In the lower zone most of the oxygen is consumed in the combustion of coke contained by the catalyst particles (regeneration zone), while in the upper part the desired chemical reaction takes place (reaction zone). The circulation of the solid between both zones leads to a steady state. As a consequence, a continuous regeneration of the catalyst takes place, avoiding the transference of big amounts of solid between reactors necessary in fluid catalytic cracking (FCC) type
reactors. There are two critical issues regarding the operation of a TZFBR. On the one hand, the gas phase oxygen must be consumed in the lower part of the reactor (although in some cases it may be desirable to maintain a small oxygen concentration in the hydrocarbon-rich zone). On the other hand, the backmixing of gases must be avoided, especially of the hydrocarbon [18]. The TZFBR was first proposed for oxidative coupling of methane [19], verifying the feasibility of the system. The applicability of this reactor to reactions where coke formation causes the deactivation of the catalyst was demonstrated in a work on the dehydrogenation of butane in a TZFBR by Callejas et al. [20]. In the last five years, the feasibility of using this type of reactor for different processes has been evaluated. For example, it has been tested in catalytic dehydrogenation processes: propane dehydrogenation using a PteSneK/gAl2O3 catalyst [21] and n-butane dehydrogenation using a PteSn/MgAl2O4 catalyst [22]. In both processes it was concluded that the TZFBR is useful for alkane dehydrogenation, the operational conditions being important for achieving maximum yield and steady state. For this reason, the operational variables were optimized in order to achieve equilibrium between the coke formation rate (in the reaction zone) and the coke combustion rate (in the regeneration zone). The reactor has recently been successfully tested for the oxidative steam reforming of ethanol [23] obtaining stable results and continuous operation without net catalyst deactivation. Moreover, a higher global yield to hydrogen has been observed in the TZFBR than in conventional fluidized-bed reactors or in fixed-bed reactors under similar operating conditions. The aim of this research is to analyse the improvements gained by the use of a TZFBR in the oxidative steam reforming of methane to obtain hydrogen and carbon monoxide. Specifically, the work focuses on studying the effects on the process performance of different operating variables such as temperature, relative gas velocity or inlet gas composition, working in the aforementioned TZFBR.
2.
Experimental
2.1.
Catalyst preparation
A commercially available g-Al2O3 (SigmaeAldrich; SBET ¼ 155 m2/g; 150 mesh) was calcined at 950 C in a muffle furnace for 1 h (heating rate ¼ 10 C/min); and the resulting solid was used as the support material (SBET ¼ 76 m2/g). The catalyst was prepared by incipient impregnation with an aqueous solution of the nickel precursor Ni(NO3)2$6H2O (99,999%; SigmaeAldrich). The resulting product was dried in an oven at 120 C for 24 h. Subsequently, the sample was calcined at 950 C in a muffle furnace for 1 h (SBET ¼ 59 m2/g). After calcination, the solid was sieved between 45 and 150 mm. This size fraction had a minimum fluidization velocity (umf) of 6.5 Ncm3/cm2.min, experimentally obtained with N2 at 800 C.
2.2.
Reaction
The experimental system consisted of a 2.8 cm i.d. fluidizedbed reactor made of quartz, with a distributing porous plate of
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quartz. A 0.4 cm o.d. quartz tube was inserted inside the reactor in order to feed methane at an intermediate point of the catalyst bed (Fig. 2). In the two-zone fluidized-bed reactor, methane was introduced through the top of the reactor and oxygen and steam through the bottom. Gaseous reagents were fed with mass flow controllers (model 5850 TR, Brooks Instruments). Water was fed with a HPLC pump (Shimadzu; model LC-10AT VP) to a vaporizer and heated up to 150 C. The resulting steam was mixed with oxygen. After the reaction, the water obtained from the outlet was collected in a cold trap and the gases were analysed by GC-TCD (Varian, CP-3800). The following parameters were defined to evaluate the performance of the process: methane conversion ðXCH4 Þ, selectivity from methane to product i (Si), and global yield to hydrogen from methane plus water ðYH2 =ðCH4 þH2 OÞ Þ. This last parameter was intended to evaluate the efficiency of the use of both raw materials (methane and water).
XCH4 ð%Þ ¼
out 100$ Fin CH4 FCH4
(9)
Fin CH4
100$ni Fout i Si ð%Þ ¼ out Fin CH4 FCH4
(10)
100$Fout H2 SH2 ð%Þ ¼ 3 Fin Fout CH4 CH4
(11)
YH2 =ðCH4 þH2 OÞ ¼
Fout H2 2Fin CH4
þ
Fin H2 O
¼
Fout H2 ð2 þ S=MÞ$Fin CH4
(12)
Table 1 e Operating variables and studied intervals in TZFBR. T ( C)
ur1 ()
ur2 ()
S/O (e)
S/M (e)
750e850
1.5e2.5
2.7e4.5
1e4
0.5e1.83
around 37.5 g. As it was already done in a previous work [23], the temporal evolution of methane conversion and distribution of products was analysed in long-lasting experiments of over 200 min. A steady performance was achieved in all the studied cases, which implies that there was no net coke formation in the TZFBR global process. In the following sections, all the results reported correspond to those obtained after 190 min on stream and thus correspond to the steady state.
3.1.
Effect of the temperature
The effect of the reactor temperature is shown in Fig. 3. An increase in methane conversion and in both carbon monoxide and hydrogen selectivities was observed when the temperature was raised. The trend of the thermodynamic equilibrium is shown in Fig. 3 in dotted lines. The increase in CO selectivity at equilibrium with the temperature is consistent with the experimental results. Although hydrogen selectivity at equilibrium decreases slightly with the increase of temperature, a small rise in hydrogen selectivity was experimentally observed. Moreover, this experimental hydrogen selectivity was significantly lower than the equilibrium selectivity, this difference being greater than that observed for carbon
XCH4
3.
Results and discussion
The operating variables and the intervals studied in the TZFBR are shown in Table 1. The total mass of catalyst in the bed was 75 g in all cases, so the mass of catalyst in each zone was
CH4
Conversion or selectivity (%)
100
SCO
SH2
XCH4 eq 90
S CO eq
80
SH2 eq
70 60 50
Exit gases
Reaction zone
H2/CO
3.6 3.2 H2/CO
2.8
eq
Regeneration zone
2.4
O2 + steam + N2
2.0
Solid flow
Quartz distributor plate
Fig. 2 e Diagram of the two-zone fluidized-bed reactor.
750
800
850
T (ºC) Fig. 3 e Effect of the temperature on the conversion and the product selectivity in a TZFBR with 2.5% Ni/Al2O3 catalyst (lines for visual help). t [ 190 min. Operating conditions: ur1 [ 2; ur2 [ 3.6; CH4/H2O/O2 [ 2/1.5/1; Wcat [ 37.5 g/ zone, W/F [ 6.3 g min/mmol (q [ 133 Nml/min).
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3.2.
Effect of the relative velocity
The effect of the relative gas velocity in the reaction zone (ur2) is shown in Fig. 4. It was observed that methane conversion decreased with the increase in the relative velocity. This is due to both the reduction of the gas contact time in the reaction zone, and the increase in the bubble size which involves an increase in the mass transfer resistance between the gas in the bubbles and the catalytic emulsion. In addition, carbon monoxide and hydrogen selectivities did not change significantly. Both were lower than the thermodynamic equilibrium values, particularly the hydrogen selectivity as already mentioned.
3.3.
Effect of the steam/oxygen ratio
The effect of the steam/oxygen ratio is shown in Fig. 5. A significant reduction in the experimental methane
XCH4
100
Conversion or selectivity (%)
monoxide selectivities. This can be explained by considering the particular contact mode of gases in the TZFBR. Steam and oxygen enter the reactor in the regeneration zone, carrying out the combustion of the coke deposited in the catalyst particles. The resulting CO2 enriched gas enters the reaction zone where CH4 is being fed. This high CO2 concentration favours methane dry reforming (Eq. (5)) more than predicted by equilibrium calculations. Consequently, the H2/CO ratios obtained experimentally with this particular contact mode were lower than in the equilibrium calculations. The differences were greater at high temperatures because of the endothermicity of the methane dry reforming reaction. In fact, such an influence of the temperature can be observed in Fig. 3 in both the experimental and the equilibrium H2/CO ratios. Other researchers have also reported the fall of this ratio with the temperature [24e26]. Indeed, De Souza et al. in 2010 [24] obtained a constant value of hydrogen selectivity with a drop in the H2/CO ratio when the temperature was increased. This was attributed to the reverse water gas shift (RWGS) reaction. Taking into account that the H2/CO ratio is equal to 2 in partial oxidation and equal to 3 in SRM, this means that if H2/ CO ratio is higher than 3, there is a significant WGS reaction (Eq. (4)) [25]. By contrast, low values of the H2/CO ratio indicate that there is a substantial RWGS reaction together with the SRM [26]. For the temperature range and conditions of this study, both WGS or RWGS reactions would occur giving H2/CO values higher or lower than 3, respectively (i.e., SH2 =SCO larger or lower than 1, as it can be seen in Fig. 1). In particular, a substantial RWGS reaction takes place for the conditions shown in Fig. 3 (CH4/H2O/O2 ¼ 2/1.5/1), rendering H2/CO values lower than 3. Under these conditions, H2/CO values higher than 3 (i.e., prevalence of the WGS reaction) would be achieved only at temperatures below 735 C, as can be seen in Fig. 1. Moreover, taking into account that in a TZFBR only a small quantity of oxygen, if any, is expected in the reaction zone, a partial oxidation reaction would not occur there. However, with the oxygen amount fed to the TZFBR reactor in these runs, coke combustion in the regeneration zone may generate enough heat to maintain the global process without requiring an external heating source.
SCO
SH 2
XCH4 eq 90 80 70
SCO eq SH2 eq
60 50
3.6
H2/CO
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H2/CO eq
3.2 2.8 2.4 2.0
3.0
3.5
4.0
4.5
ur2 Fig. 4 e Effect of the relative gas velocity in the reaction zone (ur2) on the conversion and the product selectivities in a TZFBR with 2.5% Ni/Al2O3 catalyst (lines for visual help). t [ 190 min. Operating conditions: T [ 800 C; CH4/H2O/ O2 [ 2/1.5/1; Wcat [ 37.5 g/zone, W/F [ 6.3 g min/mmol (q [ 133 Nml/min).
conversion, a rise in carbon monoxide selectivity and a significant increase in hydrogen selectivity were observed when the steam/oxygen ratio was raised. According to the experimental procedure, the increase in the steam/oxygen ratio (S/O) involves a decrease in the oxygen/methane ratio and an increase in the steam/methane ratio. When S/O ¼ 4 (CH4/H2O/O2 ¼ 2/2/0.5), the steam/ methane ratio is equal to 1, the stoichiometric value for methane steam reforming (Eq. (1)). If the flow of oxygen fed into the reactor is sufficient to burn the net coke flow (deposited on the solid catalyst particles) coming into the regeneration zone, almost all the steam can reach the upper reaction zone, maintaining the nominal steam/methane ratio. Therefore it makes sense that in these conditions the CO and H2 selectivities were high, with a H2/CO ratio near to 3 (related to SRM). The sharp fall in methane conversion with the increase in the steam/oxygen ratio could be explained by the fact that the lower the oxygen proportion in the feeding stream (high S/O values), the lower is the methane conversion. The increase in the S/O ratio also enlarges the time on stream required to achieve this stationary state, as it has been observed in a previous study (not shown). In fact, at experimental conditions of Fig. 5, for S/O ¼ 4 a time on stream around t ¼ 150 min was required to attain the stationary state. However the required time was only around t ¼ 20 min when S/O ¼ 1. This can be explained by considering that for high S/O values the
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3.4.
SH 2
XCH4 eq 90 80
S CO eq
70 60
S H2
eq
50
H2/CO
3.6 3.2 2.8
CO H 2/
eq
2.4 2.0
1.0
1.5
2.0
2.5
3.0
3.5
4.0
S/O Fig. 5 e Effect of the steam/oxygen ratio on the conversion and the product selectivities in a TZFBR with 2.5% Ni/Al2O3 catalyst (lines for visual help). t [ 190 min. Operating conditions: T [ 850 C; ur1 [ 2; ur2 [ 3.6; Wcat [ 37.5 g/ zone; W/F [ 6.3 g min/mmol, (q [ 133 Nml/min).
amount of coke remaining over the catalyst surface in steady state (i.e., as a result of the dynamic equilibrium between coke formation and coke combustion processes) is higher than when a greater proportion of feed oxygen is used. Coke does not cause a selective deactivation (i.e., the loss of activity preferentially in some specific reaction), so that the changes in the CO and H2 selectivities are similar to those shown by the equilibrium curves (Fig. 5). However, the conversion values were clearly different from the equilibrium values at high steam/oxygen values. When the steam/oxygen ratio was increased, methane steam reforming (Eq. (1)) and partial oxidation (Eq. (3)) were more favoured than methane combustion (Eq. (2)). The products of these two reactions are CO and H2. This explains the evolution of the CO and H2 selectivities shown in Fig. 5. A considerable increase in the H2/CO ratio was observed with the increase in the steam/oxygen ratio. The changes in the experimental values matched those of the equilibrium values (dotted lines in Fig. 5), but the hydrogen selectivity was noticeably lower than the equilibrium selectivity in all cases. Dong et al. in 2002 [27] investigated the oxidative steam reforming of methane in a conventional fixed-bed microreactor, using Ni/CeeZrO2 catalysts at 750 C and S/M ¼ 3. These researchers observed a qualitatively similar behaviour to that shown in the present work when modifying the oxygen/methane feeding ratio. There was a sharp fall in carbon monoxide and hydrogen selectivities and an increase in methane conversion when the oxygen content in the feed gas was increased (i.e., lowering the S/O in Fig. 5).
Effect of the steam/methane ratio
The effect of the steam/methane ratio with constant oxygen/ methane is shown in Fig. 6. The methane conversion remained approximately constant, while a reduction in CO selectivity and an almost negligible change in hydrogen selectivity were observed when the steam/methane ratio was increased. This trend can be appreciated both in the experimental and in the equilibrium values (dotted lines in Fig. 6). It is important to emphasize that even with a very low steam/ methane ratio (S/M ¼ 0.5) the system was able to achieve a steady state and to maintain methane conversion close to the equilibrium value. The evolution of the CO and H2 selectivities was equivalent to that of the equilibrium values. The experimental values of hydrogen selectivity were, as in the previous cases, significantly lower than the equilibrium values. The decrease in carbon monoxide selectivity with the increase in the steam/ methane ratio was due to the major extension of the WGS reaction, where CO is consumed. This reaction is more significant at higher steam concentrations. Other authors [28,29] have observed similar trends. The influence of the steam/methane ratio on the H2/CO ratio was significant, as shown in Fig. 6. For example, with an increase from 1 to 1.83 in the steam/methane ratio, the experimental H2/CO ratio changed from 2.5 to 3. The desirable value of this ratio depends on the later application of the syngas (e.g., for FishereTropsch if H2/CO ¼ 3 or for methanol production if H2/CO ¼ 2).
XCH4
100
Conversion or selectivity (%)
Conversion or selectivity (%)
SCO
90
SCO
SH2
XCH4 eq
S CO eq
80 70
S H2 eq
60 50
3.6
H2/CO
XCH4
100
3.2 2.8
/CO H2
eq
2.4 2.0
0.5
1.0
1.5
2.0
S/M Fig. 6 e Effect of the steam/methane ratio on the conversion and the product selectivities in a TZFBR with 2.5% Ni/Al2O3 catalyst (lines for visual help). t [ 190 min. Operating conditions: T [ 850 C; ur1 [ 2; ur2 [ 3.6; CH4/ O2 [ 2; Wcat [ 37.5 g/zone; W/F [ 6.3 g min/mmol.
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1.0
0.8
0.6
0.4
YH
2
/ (CH4+H2O)
4.
Li et al., 2007 [30] Nurunnabi et al., 2007 [31] Mukainakano et al., 2007 [32] Mukainakano et al., 2008 [33] Yoshida et al., 2009 [15] Escritori et al., 2009 [34] Cai et al., 2006 [25] TZFBR, 2.5% Ni/Al O
0.2
0.0 0
20
40
60
80
100
XCH4 (%) Fig. 7 e Comparison between the best results in the TZFBR with a 2.5% Ni/Al2O3 catalyst obtained in this work and the results found in the literature using Ni-based catalysts. Temperature range [ 750e850 C.
According to the literature [5], Ni-based catalysts require high steam/methane ratios (higher than 3e3.5) in order to avoid coke formation and catalyst deactivation. However, in the TZFBR with a 2.5 %Ni/Al2O3 catalyst, it was possible to work with very low steam/methane ratios (0.5e1.83) without net coke formation and with methane conversion higher than 95%.
3.5.
Comparison with the literature
A comparison between the best results obtained in this work using a 2.5% Ni/Al2O3 catalyst in a TZFBR and the results found in the literature with Ni-based catalysts in fixed-bed reactors in a temperature range of 750e850 C is shown in Fig. 7. In this figure, the global yield to hydrogen from methane plus water (Eq. (12)) is compared in relation to the methane conversion achieved. The results obtained with the TZFBR are among the best results found in the literature. Li et al. (2007) used Ni/Al2O3 (0.9% and 2.6% Ni in weight) catalysts [30] while Nurunnabi et al. (2007) used Ni0.2Mg0.8Al2O4 catalysts [31]. Mukainakano et al. (2007, 2008) used Ni/Al2O3 (0.9% and 2.6% Ni in weight) catalysts [32,33]. Yoshida et al. (2009) used Ni-based catalysts supported on a-Al2O3 (10.6% Ni) [15]. Escritori et al. (2009) studied the autothermal reforming of methane with Ni catalysts supported on Al2O3, Ce0.5Zr0.5O2, CeO2/Al2O3 and Ce0.5Zr0.5O2/Al2O3 [34]. Cai et al. (2006) used Ni catalysts supported in mixed oxides of Al2O3 with oxides from rare earths and transient metals (Ni/ZrxCe30 xAl70Od) [25]. Only the results obtained by Yoshida et al. (2009) [15], using a catalyst with a higher amount of nickel in (10.6% wt.), are slightly better than the results of this work. It can be said that the results of the experiments in the TZFBR show high methane conversion and a high yield to hydrogen. These results represent an improvement over most of the results found in the literature, with the additional advantage that in many of them deactivation by coke was reported [31].
Conclusions
The 2.5% Ni/Al2O3 catalyst achieved steady state behaviour in the TZFBR, with methane conversion close to equilibrium. Neither deactivation nor sintering was observed in the range of conditions under study. In this work, very low steam/ methane ratios (0.5e1.83) were used in the TZFBR without net coke formation, obtaining values of methane conversion above 95%. Methane conversion was clearly affected by the temperature. In fact, a sharp increase in the methane conversion was observed with an increase of 100 C; obtaining an improvement in hydrogen and carbon monoxide selectivities. Methane conversion also changed with the steam/oxygen ratio. A sharp fall in the methane conversion resulted from an increase in this variable. There was also a considerable increase in the hydrogen and carbon monoxide selectivities. In fact, hydrogen selectivity is greatly influenced by the steam/ oxygen ratio, but not by the other variables studied. Carbon monoxide selectivity is heavily dependent on the steam/ methane ratio, showing a sharp decrease at high values because the WGS reaction occurs to a great extent. The experimental results in the TZFBR showed high methane conversion and high hydrogen yields. The global hydrogen yield was higher than those found in the literature for Ni-based catalysts. Using the TZFBR, global yields to hydrogen between 0.52 and 0.87, with methane conversions higher than 84%, were achieved in most experiments. These high values of global hydrogen yields correspond to hydrogen selectivities lower than the equilibrium values. The low steam/methane ratios used in the TZFBR in comparison to those used in traditional reactors explains these good hydrogen global yield results.
Acknowledgements This work has been partially funded by the Spanish Ministry of Education and Science (project CTQ 2007-63420). Financial aid for the maintenance of the consolidated research group CREG has been provided by the Fondo Social Europeo (FSE) through the Gobierno de Arago´n (Arago´n, Spain). L. Pe´rez-Moreno also thanks the Gobierno de Arago´n for a grant.
Nomenclature
Symbols FBR Fi ni OSRM q RWGS S/M S/O SBET
conventional fluidized-bed reactor molar flow of product i, mmol/min number of C atoms in a molecule of product i oxidative steam reforming of methane volumetric gas flowrate of reactants, Nml/min reverse water gas shift reaction steam to methane molar feed ratio steam to oxygen molar feed ratio specific surface BET, m2/g
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 7 8 3 0 e7 8 3 8
Si
selectivity from methane to product i as defined in Eqs. (10) and (11), % SRM steam reforming of methane t time, min T temperature, C TZFBR two-zone fluidized-bed reactor gas velocity in the bed, cm/min u0 minimum fluidization velocity, cm/min umf relative gas velocity in the regeneration zone defined ur1 as u1/umf ur2 relative gas velocity in the reaction zone defined as u2/umf W/F mass load catalyst/molar flow of reactants, g min/ mmol mass load of catalyst in the bed, g Wcat WGS water gas shift reaction methane conversion XCH4 global yield to H2 as defined in Eq. (12) YH 2 =ðCH4 þH2 OÞ standard enthalpy of reaction at 298 K, kJ/mol DH0298
Superscripts in at the reactor entrance out at the reactor exit
references
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[10] Freni S, Calogero G, Cavallaro S. Hydrogen production from methane through catalytic partial oxidation reactions. J Power Sourc 2000;87:28e38. [11] Li BT, Maruyama K, Nurunnabi M, Kunimori K, Tomishige K. Effect of Ni loading on catalyst bed temperature in oxidative steam reforming of methane over a-Al2O3 supported Ni catalysts. Ind Eng Chem Res 2005;44:485e94. [12] Tomishige K, Kanazawa S, Suzuki K, Asadullah M, Sato M, Ikushima K, et al. Effective heat supply from combustion to reforming in methane reforming with CO2 and O2: comparison between Ni and Pt catalysts. Appl Catal A 2002;233:35e44. [13] Tomishige K, Nurunnabi M, Maruyama K, Kunimori K. Effect of oxygen addition to steam and dry reforming of methane on bed temperature profile over Pt and Ni catalysts. Fuel Process Technol 2004;85:1103e20. [14] Li B, Maruyama K, Nurunnabi M, Kunimori K, Tomishige K. Temperature profiles of alumina-supported noble metal catalysts in autothermal reforming of methane. Appl Catal A 2004;275:157e72. [15] Yoshida K, Begum N, Ito S, Tomishige K. Oxidative steam reforming of methane over Ni/a-Al2O3 modified with trace noble metals. Appl Catal A 2009;358:186e92. [16] Pedernera MN, Pin˜a J, Borio DO. Kinetic evaluation of carbon formation in a membrane reactor for methane reforming. Chem Eng J 2007;134:138e44. [17] Bion N, Epron F, Duprez D. Bioethanol reforming for H2 production. A comparison with hydrocarbon reforming. Catalysis 2010;22:1e55. [18] Herguido J, Mene´ndez M, Santamarı´a J. On the use of fluidized bed catalytic reactors where reduction and oxidation zones are present simultaneously. Catal Today 2005;100:181e9. [19] Ramos R, Herguido J, Mene´ndez M, Santamarı´a J. Oxidation of hydrocarbons in an in situ redox fluidized bed reactor. J Catal 1996;163:218e21. [20] Callejas C, Soler J, Herguido J, Mene´ndez M, Santamarı´a J. Catalytic dehydrogenation of n-butane in a fluidized bed reactor with separate coking and regeneration zones. Stud Surf Sci Catal 2000;130:2717e22. [21] Lobera MP, Te´llez C, Herguido J, Mene´ndez M. Transient kinetic modelling of propane dehydrogenation over a Pt-SnK/Al2O3 catalyst. Appl Catal A 2008;349:156e64. [22] Lobera MP, Te´llez C, Herguido J, Mene´ndez M. Pt-Sn/MgAl2O4 as n-butane dehydrogenation catalyst in a two-zone fluidized bed reactor. Ind Eng Chem Res 2009;48:6573e8. [23] Pe´rez-Moreno L, Soler J, Herguido J, Mene´ndez M. Stable steam reforming of ethanol in a two-zone fluidized bed reactor. Ind Eng Chem Res 2012;51:8840e8. ´ M, Maciel LJL, Lima Filho NM, Abreu CAM. [24] De Souza AEA Catalytic activity evaluation for hydrogen production via autothermal reforming of methane. Catal Today 2010;149:413e7. [25] Cai X, Dong X, Lin W. Autothermal reforming of methane over Ni catalysts supported on CuOeZrO2eCeO2eA12O3. J Nat Gas Chem 2006;15:122e6. [26] Souza MMVM, Schmal M. Autothermal reforming of methane over Pt/ZrO2/Al2O3 catalysts. Appl Catal A 2005;281:19e24. [27] Dong WS, Roh HS, Jun KW, Park SE, Oh YS. Methane reforming over Ni/Ce-ZrO2 catalysts: effect of nickel content. Appl Catal A 2007;226:63e72. [28] Zeppieri M, Villa PL, Verdone N, Scarsella M, De Filippis P. Kinetic of methane steam reforming reaction over nickeland rhodium-based catalysts. Appl Catal A 2010;387:147e54. [29] Hou K, Hugues R. The kinetics of methane steam reforming over a Ni/a-Al2O catalyst. Chem Eng J 2001;82:311e28.
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[30] Li B, Kado S, Mukainakano Y, Miyazawa T, Miyao T, Naito S, et al. Surface modification of Ni catalysts with trace Pt for oxidative steam reforming of methane. J Catal 2007;245:144e55. [31] Nurunnabi M, Mukainakano Y, Kado S, Miyao T, Naito S, Okumura K, et al. Catalytic performance and characterization of Pd/Ni0.2Mg0.8Al2O4 in oxidative steam reforming of methane under atmospheric and pressurized conditions. Appl Catal A 2007;325:154e62. [32] Mukainakano Y, Li B, Kado S, Miyazawa T, Okumura K, Miyao T, et al. Surface modification of Ni catalysts with trace
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Stable hydrogen production by methane steam reforming in a two-zone fluidized-bed reactor: Effect of the operating variables L. Pe´rez-Moreno, J. Soler, J. Herguido*, M. Mene´ndez Catalysis, Molecular Separations and Reactor Engineering Group (CREG), Arago´n Institute of Engineering Research (I3A), University of Zaragoza, 50009 Zaragoza, Spain
article info
abstract
Article history:
The oxidative steam reforming of methane in a two-zone fluidized-bed reactor (TZFBR) was
Received 4 March 2013
investigated over a Ni/Al2O3 catalyst. The effects of the main operating variables (tem-
Received in revised form
perature, steam/oxygen ratio, steam/methane ratio and relative velocity with respect to
19 April 2013
the minimum fluidization velocity) have been studied. A comparison has been made with
Accepted 21 April 2013
results given in the literature in terms of hydrogen yield. Despite working with very low
Available online 18 May 2013
steam/methane ratios, high values of hydrogen yield at both high methane conversion and at steady state were obtained in the TZFBR.
Keywords:
Copyright ª 2013, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights
Oxidative steam reforming
reserved.
Methane Ni/Al2O3 catalyst TZFBR
1.
Introduction
Methane is the main component of natural gas (NG) and landfill gas as well as being a by-product of oil refining and chemical processing. It has enormous potential value as a source of clean fossil energy. Natural gas reservoirs are large and widespread throughout the world, mainly in the Middle East and Russia, although other areas also have their fair share. The estimated reserves at the end of 2006 amounted to approximately 6300 trillion cubic feet [1]. In recent years, hydrogen has attracted great interest as a clean fuel for combustion engines and fuel cells [2]. Among all the potential sources for obtaining hydrogen, natural gas, which has methane as its main component, is considered a good option because it is clean, abundant and can be easily converted into hydrogen [3].
Steam reforming of methane (SRM) is the main industrial route for obtaining hydrogen and synthesis gas [4]. About 50% of the worldwide hydrogen demand is satisfied by steam reforming of methane. However, in this process there are two main problems: a) the reaction is endothermic and requires external heating; b) nickel based catalysts require the use of high steam/methane ratios (3e3.5) in order to avoid the formation of coke that causes catalyst deactivation [5]. In order to solve these problems, several alternatives have been studied. The reduction of the steam/methane ratio decreases the heat requirements although it may increase coke deposition. The introduction of oxygen with the feed adds exothermic reactions such as combustion or partial oxidation. This process is called oxidative steam reforming of methane (OSRM). It avoids some of the problems found in conventional steam methane reforming but has some drawbacks.
* Corresponding author. Tel.: þ34 976762393; fax: þ34 976762142. E-mail address: [email protected] (J. Herguido). 0360-3199/$ e see front matter Copyright ª 2013, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.ijhydene.2013.04.122
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CH4 þ 2O2 /CO2 þ 2H2 O 1 CH4 þ O2 /CO þ 2H2 2
DH0298 ¼ 803 kJ=mol
DH0298 ¼ 36 kJ=mol
(1) (2) (3)
Other reactions that can also occur are the water gas shift reaction (WGS) (Eq. (4)) and methane dry reforming (Eq. (5)) CO þ H2 O4CO2 þ H2 CH4 þ CO2 /2CO þ 2H2
DH0298 ¼ 41 kJ=mol DH0298 ¼ 247 kJ=mol
(4) (5)
Although oxidative steam methane reforming avoids the need for the large heat supply that characterizes conventional steam methane reforming, this process presents several drawbacks: 1) the formation of hot spots in the initial part of the bed [6]; 2) low activity of the nickel based catalysts due to the oxidation of the nickel metallic species; 3) selectivity loss; and 4) coke formation in the final part of the bed due to the lack of oxygen [7]. The use of Ni-supported catalysts in oxidative steam methane reforming has been extensively studied [3,8e11]. It has been reported that Rh and Pd based catalysts are effective for eliminating the formation of hot spots [12,13] whereas Nibased catalysts tend to form them [11,12,14]. However, noble metals such as Pt, Pd and Rh are not suitable components due to their limited availability and their high price. For this reason, two strategies can be proposed: a) the development of Ni-based catalysts modified with low quantities of noble metals [15] or b) the study of processes with Ni-based catalysts but using different types of reactor in order to minimize the problems. The high temperatures associated with methane steam reforming also favour coke formation. The deactivation of nickel catalysts by carbon formation is a significant problem in methane reforming because of Ni surface fouling, catalyst pore blockage and disintegration of the support material [16]. Thermodynamically, the most probable reactions for carbon formation are the following: CH4 4C þ 2H2 ðCH4 crackingÞ 2CO4C þ CO2 ðBoudouardÞ
DH0298 DH0298
¼ 122:3 kJ=mol
(6)
¼ 125:2 kJ=mol
(7)
CO þ H2 4C þ H2 OðCO reductionÞ DH0298 ¼ 84:0 kJ=mol
1.1.
100 98
Methane conversion (%)
DH0298 ¼ 206 kJ=mol
96 94
C H 4 /H 2 O /O 2
92
2/1.25/1.2 5 2/1.5/1 2/1.67/0.8 3 2/1.875 /0 .625 2/2/0.5
90 88
90
80
SH2 (%)
CH4 þ H2 O/CO þ 3H2
increased because of the preferential formation of CO at high temperatures. This fact is consistent with the WGS equilibrium (Eq. (4)) because CO2 generation is favoured at low temperatures due to the exothermicity of the reaction, while the reverse reaction (RWGS) yielding CO is favoured at high temperatures. At 900 C, total conversion of methane can be achieved, yielding syngas with a composition close to that given by Eq. (1) (75% H2 þ 25% CO). Methane and steam conversions are very high but not total, even at 700 C. The hydrogen expected in dry gases amounts to 70 vol.%. In our study, a simulation of the thermodynamic equilibrium was done using the Hysys v7.0 Aspen Tech program with the Gibbs reactor option. Methane conversion and selectivity to the main products of the process have been studied in relation to the operating variables: temperature, steam/oxygen ratio and steam/methane ratio. In Fig. 1 the variation of methane conversion and selectivities ðSH2 and SH2 =SCO Þ with the temperature are shown. The methane conversion increased with the temperature, this
70
60
1.1
SH2/SCO
In oxidative steam methane reforming, oxygen must be provided to the reactor together with the steam in order to combine steam methane reforming (endothermic, Eq. (1)) with two exothermic reactions: methane combustion (Eq. (2)) and the partial oxidation of methane (Eq. (3)). In this way the global process may be autothermal.
1.0
0.9
(8)
Thermodynamic equilibrium
Bion et al. (2010) studied the effect of temperature on the gas composition at thermodynamic equilibrium for methane steam reforming working with H2O/CH4 ¼ 1 [17]. They observed maximum hydrogen production at around 700 C. Above this temperature, the H2 concentration no longer
0.8 700
750
800
850
900
T (ºC)
Fig. 1 e Methane conversion and CO and H2 selectivities in thermodynamic equilibrium for several CH4/H2O/O2 ratios as a function of temperature.
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increase being greater for higher steam contents in the reacting gas. With regard to CO selectivity (not shown in Fig. 1), there was a similar increase with the temperature irrespective of the methane/steam/oxygen inlet composition. In the case of H2 selectivity, when the temperature was increased the selectivity slightly decreased for all the methane/steam/oxygen ratios. These results can be attributed to both the endothermic character of the main steam reforming reaction (Eq. (1)) that favours the conversion and the CO and H2 generation, and the exothermal character of the water gas shift reaction (Eq. (4)) that at high temperatures favours CO production at the expense of H2 (i.e., reverse water gas shift reaction RWGS). The SH2 =SCO ratio in equilibrium conditions had a value near to one (i.e., a gas composition H2/ CO ratio of around 3 according to Eq. (1)) for the range of temperatures considered. However, the lower the steam content in the reacting gas, the lower was the SH2 =SCO ratio at a given temperature. Moreover, a decrease with temperature was observed in the SH2 =SCO value at a given gas inlet composition. The effect of the steam/oxygen ratio at constant temperature and constant CH4/(H2O þ O2) ratio has been analysed. A sharp increase in CO selectivity was observed when the steam/oxygen ratio was increased, while methane conversion was almost unaffected by this variable. H2 selectivity greatly increased with the steam/oxygen ratio. This behaviour was due to the greater yield of the methane reforming reaction (Eq. (1)) at a high steam/oxygen ratio. The effect of the steam/methane ratio while maintaining constant both the CH4/O2 ratio and the temperature has also been analysed. A sharp decrease in the CO selectivity was observed, while the methane conversion was not affected by this variable. The H2 selectivity increased with the steam/ methane ratio. It seems clear that an excess of H2O favoured the WGS reaction (Eq. (4)).
1.2.
Reactors
The use of the two-zone fluidized-bed reactor (TZFBR) could serve to overcome the problems of oxidative steam methane reforming. Selectivity loss is not expected in this reactor because regeneration and reaction take place in the same vessel. The circulation of the solid avoids the formation of hot spots. At the same time, the catalyst regeneration will produce heat that can be used in steam reforming, mitigating in some degree the process endothermicity. Moreover, we aim to reduce the steam/methane ratio without having problems with coke deposition. This type of reactor consists of a fluidized bed where an oxygenesteam mixture is fed through the lower part of the reactor while methane is introduced at an intermediate point of the bed [18]. In this way, two zones are created in the reactor. In the lower zone most of the oxygen is consumed in the combustion of coke contained by the catalyst particles (regeneration zone), while in the upper part the desired chemical reaction takes place (reaction zone). The circulation of the solid between both zones leads to a steady state. As a consequence, a continuous regeneration of the catalyst takes place, avoiding the transference of big amounts of solid between reactors necessary in fluid catalytic cracking (FCC) type
reactors. There are two critical issues regarding the operation of a TZFBR. On the one hand, the gas phase oxygen must be consumed in the lower part of the reactor (although in some cases it may be desirable to maintain a small oxygen concentration in the hydrocarbon-rich zone). On the other hand, the backmixing of gases must be avoided, especially of the hydrocarbon [18]. The TZFBR was first proposed for oxidative coupling of methane [19], verifying the feasibility of the system. The applicability of this reactor to reactions where coke formation causes the deactivation of the catalyst was demonstrated in a work on the dehydrogenation of butane in a TZFBR by Callejas et al. [20]. In the last five years, the feasibility of using this type of reactor for different processes has been evaluated. For example, it has been tested in catalytic dehydrogenation processes: propane dehydrogenation using a PteSneK/gAl2O3 catalyst [21] and n-butane dehydrogenation using a PteSn/MgAl2O4 catalyst [22]. In both processes it was concluded that the TZFBR is useful for alkane dehydrogenation, the operational conditions being important for achieving maximum yield and steady state. For this reason, the operational variables were optimized in order to achieve equilibrium between the coke formation rate (in the reaction zone) and the coke combustion rate (in the regeneration zone). The reactor has recently been successfully tested for the oxidative steam reforming of ethanol [23] obtaining stable results and continuous operation without net catalyst deactivation. Moreover, a higher global yield to hydrogen has been observed in the TZFBR than in conventional fluidized-bed reactors or in fixed-bed reactors under similar operating conditions. The aim of this research is to analyse the improvements gained by the use of a TZFBR in the oxidative steam reforming of methane to obtain hydrogen and carbon monoxide. Specifically, the work focuses on studying the effects on the process performance of different operating variables such as temperature, relative gas velocity or inlet gas composition, working in the aforementioned TZFBR.
2.
Experimental
2.1.
Catalyst preparation
A commercially available g-Al2O3 (SigmaeAldrich; SBET ¼ 155 m2/g; 150 mesh) was calcined at 950 C in a muffle furnace for 1 h (heating rate ¼ 10 C/min); and the resulting solid was used as the support material (SBET ¼ 76 m2/g). The catalyst was prepared by incipient impregnation with an aqueous solution of the nickel precursor Ni(NO3)2$6H2O (99,999%; SigmaeAldrich). The resulting product was dried in an oven at 120 C for 24 h. Subsequently, the sample was calcined at 950 C in a muffle furnace for 1 h (SBET ¼ 59 m2/g). After calcination, the solid was sieved between 45 and 150 mm. This size fraction had a minimum fluidization velocity (umf) of 6.5 Ncm3/cm2.min, experimentally obtained with N2 at 800 C.
2.2.
Reaction
The experimental system consisted of a 2.8 cm i.d. fluidizedbed reactor made of quartz, with a distributing porous plate of
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quartz. A 0.4 cm o.d. quartz tube was inserted inside the reactor in order to feed methane at an intermediate point of the catalyst bed (Fig. 2). In the two-zone fluidized-bed reactor, methane was introduced through the top of the reactor and oxygen and steam through the bottom. Gaseous reagents were fed with mass flow controllers (model 5850 TR, Brooks Instruments). Water was fed with a HPLC pump (Shimadzu; model LC-10AT VP) to a vaporizer and heated up to 150 C. The resulting steam was mixed with oxygen. After the reaction, the water obtained from the outlet was collected in a cold trap and the gases were analysed by GC-TCD (Varian, CP-3800). The following parameters were defined to evaluate the performance of the process: methane conversion ðXCH4 Þ, selectivity from methane to product i (Si), and global yield to hydrogen from methane plus water ðYH2 =ðCH4 þH2 OÞ Þ. This last parameter was intended to evaluate the efficiency of the use of both raw materials (methane and water).
XCH4 ð%Þ ¼
out 100$ Fin CH4 FCH4
(9)
Fin CH4
100$ni Fout i Si ð%Þ ¼ out Fin CH4 FCH4
(10)
100$Fout H2 SH2 ð%Þ ¼ 3 Fin Fout CH4 CH4
(11)
YH2 =ðCH4 þH2 OÞ ¼
Fout H2 2Fin CH4
þ
Fin H2 O
¼
Fout H2 ð2 þ S=MÞ$Fin CH4
(12)
Table 1 e Operating variables and studied intervals in TZFBR. T ( C)
ur1 ()
ur2 ()
S/O (e)
S/M (e)
750e850
1.5e2.5
2.7e4.5
1e4
0.5e1.83
around 37.5 g. As it was already done in a previous work [23], the temporal evolution of methane conversion and distribution of products was analysed in long-lasting experiments of over 200 min. A steady performance was achieved in all the studied cases, which implies that there was no net coke formation in the TZFBR global process. In the following sections, all the results reported correspond to those obtained after 190 min on stream and thus correspond to the steady state.
3.1.
Effect of the temperature
The effect of the reactor temperature is shown in Fig. 3. An increase in methane conversion and in both carbon monoxide and hydrogen selectivities was observed when the temperature was raised. The trend of the thermodynamic equilibrium is shown in Fig. 3 in dotted lines. The increase in CO selectivity at equilibrium with the temperature is consistent with the experimental results. Although hydrogen selectivity at equilibrium decreases slightly with the increase of temperature, a small rise in hydrogen selectivity was experimentally observed. Moreover, this experimental hydrogen selectivity was significantly lower than the equilibrium selectivity, this difference being greater than that observed for carbon
XCH4
3.
Results and discussion
The operating variables and the intervals studied in the TZFBR are shown in Table 1. The total mass of catalyst in the bed was 75 g in all cases, so the mass of catalyst in each zone was
CH4
Conversion or selectivity (%)
100
SCO
SH2
XCH4 eq 90
S CO eq
80
SH2 eq
70 60 50
Exit gases
Reaction zone
H2/CO
3.6 3.2 H2/CO
2.8
eq
Regeneration zone
2.4
O2 + steam + N2
2.0
Solid flow
Quartz distributor plate
Fig. 2 e Diagram of the two-zone fluidized-bed reactor.
750
800
850
T (ºC) Fig. 3 e Effect of the temperature on the conversion and the product selectivity in a TZFBR with 2.5% Ni/Al2O3 catalyst (lines for visual help). t [ 190 min. Operating conditions: ur1 [ 2; ur2 [ 3.6; CH4/H2O/O2 [ 2/1.5/1; Wcat [ 37.5 g/ zone, W/F [ 6.3 g min/mmol (q [ 133 Nml/min).
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 7 8 3 0 e7 8 3 8
3.2.
Effect of the relative velocity
The effect of the relative gas velocity in the reaction zone (ur2) is shown in Fig. 4. It was observed that methane conversion decreased with the increase in the relative velocity. This is due to both the reduction of the gas contact time in the reaction zone, and the increase in the bubble size which involves an increase in the mass transfer resistance between the gas in the bubbles and the catalytic emulsion. In addition, carbon monoxide and hydrogen selectivities did not change significantly. Both were lower than the thermodynamic equilibrium values, particularly the hydrogen selectivity as already mentioned.
3.3.
Effect of the steam/oxygen ratio
The effect of the steam/oxygen ratio is shown in Fig. 5. A significant reduction in the experimental methane
XCH4
100
Conversion or selectivity (%)
monoxide selectivities. This can be explained by considering the particular contact mode of gases in the TZFBR. Steam and oxygen enter the reactor in the regeneration zone, carrying out the combustion of the coke deposited in the catalyst particles. The resulting CO2 enriched gas enters the reaction zone where CH4 is being fed. This high CO2 concentration favours methane dry reforming (Eq. (5)) more than predicted by equilibrium calculations. Consequently, the H2/CO ratios obtained experimentally with this particular contact mode were lower than in the equilibrium calculations. The differences were greater at high temperatures because of the endothermicity of the methane dry reforming reaction. In fact, such an influence of the temperature can be observed in Fig. 3 in both the experimental and the equilibrium H2/CO ratios. Other researchers have also reported the fall of this ratio with the temperature [24e26]. Indeed, De Souza et al. in 2010 [24] obtained a constant value of hydrogen selectivity with a drop in the H2/CO ratio when the temperature was increased. This was attributed to the reverse water gas shift (RWGS) reaction. Taking into account that the H2/CO ratio is equal to 2 in partial oxidation and equal to 3 in SRM, this means that if H2/ CO ratio is higher than 3, there is a significant WGS reaction (Eq. (4)) [25]. By contrast, low values of the H2/CO ratio indicate that there is a substantial RWGS reaction together with the SRM [26]. For the temperature range and conditions of this study, both WGS or RWGS reactions would occur giving H2/CO values higher or lower than 3, respectively (i.e., SH2 =SCO larger or lower than 1, as it can be seen in Fig. 1). In particular, a substantial RWGS reaction takes place for the conditions shown in Fig. 3 (CH4/H2O/O2 ¼ 2/1.5/1), rendering H2/CO values lower than 3. Under these conditions, H2/CO values higher than 3 (i.e., prevalence of the WGS reaction) would be achieved only at temperatures below 735 C, as can be seen in Fig. 1. Moreover, taking into account that in a TZFBR only a small quantity of oxygen, if any, is expected in the reaction zone, a partial oxidation reaction would not occur there. However, with the oxygen amount fed to the TZFBR reactor in these runs, coke combustion in the regeneration zone may generate enough heat to maintain the global process without requiring an external heating source.
SCO
SH 2
XCH4 eq 90 80 70
SCO eq SH2 eq
60 50
3.6
H2/CO
7834
H2/CO eq
3.2 2.8 2.4 2.0
3.0
3.5
4.0
4.5
ur2 Fig. 4 e Effect of the relative gas velocity in the reaction zone (ur2) on the conversion and the product selectivities in a TZFBR with 2.5% Ni/Al2O3 catalyst (lines for visual help). t [ 190 min. Operating conditions: T [ 800 C; CH4/H2O/ O2 [ 2/1.5/1; Wcat [ 37.5 g/zone, W/F [ 6.3 g min/mmol (q [ 133 Nml/min).
conversion, a rise in carbon monoxide selectivity and a significant increase in hydrogen selectivity were observed when the steam/oxygen ratio was raised. According to the experimental procedure, the increase in the steam/oxygen ratio (S/O) involves a decrease in the oxygen/methane ratio and an increase in the steam/methane ratio. When S/O ¼ 4 (CH4/H2O/O2 ¼ 2/2/0.5), the steam/ methane ratio is equal to 1, the stoichiometric value for methane steam reforming (Eq. (1)). If the flow of oxygen fed into the reactor is sufficient to burn the net coke flow (deposited on the solid catalyst particles) coming into the regeneration zone, almost all the steam can reach the upper reaction zone, maintaining the nominal steam/methane ratio. Therefore it makes sense that in these conditions the CO and H2 selectivities were high, with a H2/CO ratio near to 3 (related to SRM). The sharp fall in methane conversion with the increase in the steam/oxygen ratio could be explained by the fact that the lower the oxygen proportion in the feeding stream (high S/O values), the lower is the methane conversion. The increase in the S/O ratio also enlarges the time on stream required to achieve this stationary state, as it has been observed in a previous study (not shown). In fact, at experimental conditions of Fig. 5, for S/O ¼ 4 a time on stream around t ¼ 150 min was required to attain the stationary state. However the required time was only around t ¼ 20 min when S/O ¼ 1. This can be explained by considering that for high S/O values the
7835
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 7 8 3 0 e7 8 3 8
3.4.
SH 2
XCH4 eq 90 80
S CO eq
70 60
S H2
eq
50
H2/CO
3.6 3.2 2.8
CO H 2/
eq
2.4 2.0
1.0
1.5
2.0
2.5
3.0
3.5
4.0
S/O Fig. 5 e Effect of the steam/oxygen ratio on the conversion and the product selectivities in a TZFBR with 2.5% Ni/Al2O3 catalyst (lines for visual help). t [ 190 min. Operating conditions: T [ 850 C; ur1 [ 2; ur2 [ 3.6; Wcat [ 37.5 g/ zone; W/F [ 6.3 g min/mmol, (q [ 133 Nml/min).
amount of coke remaining over the catalyst surface in steady state (i.e., as a result of the dynamic equilibrium between coke formation and coke combustion processes) is higher than when a greater proportion of feed oxygen is used. Coke does not cause a selective deactivation (i.e., the loss of activity preferentially in some specific reaction), so that the changes in the CO and H2 selectivities are similar to those shown by the equilibrium curves (Fig. 5). However, the conversion values were clearly different from the equilibrium values at high steam/oxygen values. When the steam/oxygen ratio was increased, methane steam reforming (Eq. (1)) and partial oxidation (Eq. (3)) were more favoured than methane combustion (Eq. (2)). The products of these two reactions are CO and H2. This explains the evolution of the CO and H2 selectivities shown in Fig. 5. A considerable increase in the H2/CO ratio was observed with the increase in the steam/oxygen ratio. The changes in the experimental values matched those of the equilibrium values (dotted lines in Fig. 5), but the hydrogen selectivity was noticeably lower than the equilibrium selectivity in all cases. Dong et al. in 2002 [27] investigated the oxidative steam reforming of methane in a conventional fixed-bed microreactor, using Ni/CeeZrO2 catalysts at 750 C and S/M ¼ 3. These researchers observed a qualitatively similar behaviour to that shown in the present work when modifying the oxygen/methane feeding ratio. There was a sharp fall in carbon monoxide and hydrogen selectivities and an increase in methane conversion when the oxygen content in the feed gas was increased (i.e., lowering the S/O in Fig. 5).
Effect of the steam/methane ratio
The effect of the steam/methane ratio with constant oxygen/ methane is shown in Fig. 6. The methane conversion remained approximately constant, while a reduction in CO selectivity and an almost negligible change in hydrogen selectivity were observed when the steam/methane ratio was increased. This trend can be appreciated both in the experimental and in the equilibrium values (dotted lines in Fig. 6). It is important to emphasize that even with a very low steam/ methane ratio (S/M ¼ 0.5) the system was able to achieve a steady state and to maintain methane conversion close to the equilibrium value. The evolution of the CO and H2 selectivities was equivalent to that of the equilibrium values. The experimental values of hydrogen selectivity were, as in the previous cases, significantly lower than the equilibrium values. The decrease in carbon monoxide selectivity with the increase in the steam/ methane ratio was due to the major extension of the WGS reaction, where CO is consumed. This reaction is more significant at higher steam concentrations. Other authors [28,29] have observed similar trends. The influence of the steam/methane ratio on the H2/CO ratio was significant, as shown in Fig. 6. For example, with an increase from 1 to 1.83 in the steam/methane ratio, the experimental H2/CO ratio changed from 2.5 to 3. The desirable value of this ratio depends on the later application of the syngas (e.g., for FishereTropsch if H2/CO ¼ 3 or for methanol production if H2/CO ¼ 2).
XCH4
100
Conversion or selectivity (%)
Conversion or selectivity (%)
SCO
90
SCO
SH2
XCH4 eq
S CO eq
80 70
S H2 eq
60 50
3.6
H2/CO
XCH4
100
3.2 2.8
/CO H2
eq
2.4 2.0
0.5
1.0
1.5
2.0
S/M Fig. 6 e Effect of the steam/methane ratio on the conversion and the product selectivities in a TZFBR with 2.5% Ni/Al2O3 catalyst (lines for visual help). t [ 190 min. Operating conditions: T [ 850 C; ur1 [ 2; ur2 [ 3.6; CH4/ O2 [ 2; Wcat [ 37.5 g/zone; W/F [ 6.3 g min/mmol.
7836
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 7 8 3 0 e7 8 3 8
1.0
0.8
0.6
0.4
YH
2
/ (CH4+H2O)
4.
Li et al., 2007 [30] Nurunnabi et al., 2007 [31] Mukainakano et al., 2007 [32] Mukainakano et al., 2008 [33] Yoshida et al., 2009 [15] Escritori et al., 2009 [34] Cai et al., 2006 [25] TZFBR, 2.5% Ni/Al O
0.2
0.0 0
20
40
60
80
100
XCH4 (%) Fig. 7 e Comparison between the best results in the TZFBR with a 2.5% Ni/Al2O3 catalyst obtained in this work and the results found in the literature using Ni-based catalysts. Temperature range [ 750e850 C.
According to the literature [5], Ni-based catalysts require high steam/methane ratios (higher than 3e3.5) in order to avoid coke formation and catalyst deactivation. However, in the TZFBR with a 2.5 %Ni/Al2O3 catalyst, it was possible to work with very low steam/methane ratios (0.5e1.83) without net coke formation and with methane conversion higher than 95%.
3.5.
Comparison with the literature
A comparison between the best results obtained in this work using a 2.5% Ni/Al2O3 catalyst in a TZFBR and the results found in the literature with Ni-based catalysts in fixed-bed reactors in a temperature range of 750e850 C is shown in Fig. 7. In this figure, the global yield to hydrogen from methane plus water (Eq. (12)) is compared in relation to the methane conversion achieved. The results obtained with the TZFBR are among the best results found in the literature. Li et al. (2007) used Ni/Al2O3 (0.9% and 2.6% Ni in weight) catalysts [30] while Nurunnabi et al. (2007) used Ni0.2Mg0.8Al2O4 catalysts [31]. Mukainakano et al. (2007, 2008) used Ni/Al2O3 (0.9% and 2.6% Ni in weight) catalysts [32,33]. Yoshida et al. (2009) used Ni-based catalysts supported on a-Al2O3 (10.6% Ni) [15]. Escritori et al. (2009) studied the autothermal reforming of methane with Ni catalysts supported on Al2O3, Ce0.5Zr0.5O2, CeO2/Al2O3 and Ce0.5Zr0.5O2/Al2O3 [34]. Cai et al. (2006) used Ni catalysts supported in mixed oxides of Al2O3 with oxides from rare earths and transient metals (Ni/ZrxCe30 xAl70Od) [25]. Only the results obtained by Yoshida et al. (2009) [15], using a catalyst with a higher amount of nickel in (10.6% wt.), are slightly better than the results of this work. It can be said that the results of the experiments in the TZFBR show high methane conversion and a high yield to hydrogen. These results represent an improvement over most of the results found in the literature, with the additional advantage that in many of them deactivation by coke was reported [31].
Conclusions
The 2.5% Ni/Al2O3 catalyst achieved steady state behaviour in the TZFBR, with methane conversion close to equilibrium. Neither deactivation nor sintering was observed in the range of conditions under study. In this work, very low steam/ methane ratios (0.5e1.83) were used in the TZFBR without net coke formation, obtaining values of methane conversion above 95%. Methane conversion was clearly affected by the temperature. In fact, a sharp increase in the methane conversion was observed with an increase of 100 C; obtaining an improvement in hydrogen and carbon monoxide selectivities. Methane conversion also changed with the steam/oxygen ratio. A sharp fall in the methane conversion resulted from an increase in this variable. There was also a considerable increase in the hydrogen and carbon monoxide selectivities. In fact, hydrogen selectivity is greatly influenced by the steam/ oxygen ratio, but not by the other variables studied. Carbon monoxide selectivity is heavily dependent on the steam/ methane ratio, showing a sharp decrease at high values because the WGS reaction occurs to a great extent. The experimental results in the TZFBR showed high methane conversion and high hydrogen yields. The global hydrogen yield was higher than those found in the literature for Ni-based catalysts. Using the TZFBR, global yields to hydrogen between 0.52 and 0.87, with methane conversions higher than 84%, were achieved in most experiments. These high values of global hydrogen yields correspond to hydrogen selectivities lower than the equilibrium values. The low steam/methane ratios used in the TZFBR in comparison to those used in traditional reactors explains these good hydrogen global yield results.
Acknowledgements This work has been partially funded by the Spanish Ministry of Education and Science (project CTQ 2007-63420). Financial aid for the maintenance of the consolidated research group CREG has been provided by the Fondo Social Europeo (FSE) through the Gobierno de Arago´n (Arago´n, Spain). L. Pe´rez-Moreno also thanks the Gobierno de Arago´n for a grant.
Nomenclature
Symbols FBR Fi ni OSRM q RWGS S/M S/O SBET
conventional fluidized-bed reactor molar flow of product i, mmol/min number of C atoms in a molecule of product i oxidative steam reforming of methane volumetric gas flowrate of reactants, Nml/min reverse water gas shift reaction steam to methane molar feed ratio steam to oxygen molar feed ratio specific surface BET, m2/g
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 7 8 3 0 e7 8 3 8
Si
selectivity from methane to product i as defined in Eqs. (10) and (11), % SRM steam reforming of methane t time, min T temperature, C TZFBR two-zone fluidized-bed reactor gas velocity in the bed, cm/min u0 minimum fluidization velocity, cm/min umf relative gas velocity in the regeneration zone defined ur1 as u1/umf ur2 relative gas velocity in the reaction zone defined as u2/umf W/F mass load catalyst/molar flow of reactants, g min/ mmol mass load of catalyst in the bed, g Wcat WGS water gas shift reaction methane conversion XCH4 global yield to H2 as defined in Eq. (12) YH 2 =ðCH4 þH2 OÞ standard enthalpy of reaction at 298 K, kJ/mol DH0298
Superscripts in at the reactor entrance out at the reactor exit
references
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