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Transient n-butane partial oxidation kinetics over VPO
Applied Catalysis A: General 263 (2004) 193–202
Transient n-butane partial oxidation kinetics over VPO Mar´ıa J. Lorences a,1 , Gregory S. Patience b,∗ , Fernando V. D´ıez a , José Coca a a
Department of Chemical Engineering and Environmental Technology, University of Oviedo, Julián Claver´ıa s/n, Oviedo, Asturias 33209, Spain b DuPont Ibérica, S.A., Tamón, Avilés 33469, Spain Received 26 June 2003; received in revised form 5 December 2003; accepted 8 December 2003
Abstract A novel reactor configuration is proposed to study the kinetics of heterogeneous catalytic reactions. It is based on the bubbling fluidization regime; gas velocities are sufficiently low to minimize bypassing but high enough to ensure solids backmixing and a uniform temperature gradient. The reactor is equipped with a gas manifold designed to operate the reactor either under steady-state feed conditions or with cyclical feeds. By cycling VPO catalyst between a reducing environment with butane and an oxidizing environment, we quantify the lattice oxygen contribution to the reaction. Furthermore, we show that carbon can build up on the VPO surface and this phenomenon increases with an increase in the severity of the reducing environment. We propose a simple kinetic model involving three sites—V5+ , V4+ and VC4 . The model adequately characterizes the transient partial oxidation of butane but further research is required to create a generalized model that fits all reaction environments. © 2003 Elsevier B.V. All rights reserved. Keywords: Transient kinetics; VPO; Butane oxidation; Maleic anhydride; Fluid bed
1. Introduction Kinetic modeling of butane partial oxidation to maleic anhydride spans over 30 years [1,2]. There is a general consensus that the reaction is near first order in butane (under highly oxidizing conditions) and that most experimental evidence supports a Mars van Krevelen mechanism. For conventional fixed beds—fuel lean and oxygen rich—this knowledge may be sufficient for reactor design and process control. However, several alternative reactor types are commercial and new concepts under development that may either operate anaerobically or fuel rich, oxygen lean conditions. For example, butane concentrations in some fluid beds may reach as high as 5%. To avoid the flammability regime in the plenum, where there are no solids to quench free radicals, 100% butane may be sparged directly into the reactor. Therefore, on a microscopic time scale, catalyst would experience near anaerobic conditions at the sparger ∗ Corresponding author. Current addresses: INVISTA (International) S.A., European Technical Centre, 146, route Nant-d’Avril, CH-1217 Meyrin, Geneva, Switzerland. Tel.: +41-22-717-6996; fax: +41-22-717-6868. E-mail address: [email protected] (G.S. Patience). 1 HALDOR TOPSØE A/S, Nymøllevej 55, 2800 Lyngby, Denmark.
0926-860X/$ – see front matter © 2003 Elsevier B.V. All rights reserved. doi:10.1016/j.apcata.2003.12.023
tips but fuel lean conditions at the grid. DuPont commercialized a circulating fluidized bed concept in which catalyst is shuttled between a net oxidizing and a net reducing environment [3]. High butane feed concentrations are not restricted to fluid bed regimes. Parkinson and D’Aquino [4] report that Pantochim operates a fixed bed under a net reducing environment and sparges oxygen into their butane feed stream. Finally, membrane reactors are under development [5,6], in which high butane concentrations are fed to a fixed bed and oxygen diffuses across a porous “membrane” along the length. Clearly, for design purposes, a more fundamental understanding of the prevailing kinetics would be desirable to reduce the risk at the design stage and perhaps help optimize reactor performance when operational. Several novel experimental programs and analytical techniques have been developed to explore these revolutionary commercial designs. For example, Golbig and Werther [7] built a lab scale circulating fluidized bed with a 21 mm diameter riser. Schuurman and Gleaves [8] and Mills et al. [9] have adapted TAP reactors to measure reactor transients at elevated pressures. Mallada et al. [10] were among the first to explore the effect of high butane concentrations on reaction kinetics in lab scale fixed bed reactors. Recently, Wang and Barteau [11] developed a novel oscillating microbalance reactor to study oxidation and reduction steps independently.
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Nomenclature Eai ki k0i LOC MA ri Vi+ VPO
activation energy (kcal mol−1 ) reaction rate constant (s−1 ) pre-exponential factor (s−1 ) lattice oxygen contribution (mmol O2 /kg VPO) maleic anhydride concentration (mol m−3 ) reaction rate (mol s−1 m−3 ) oxidation state of the catalyst vanadium-phosphorous-oxide catalysts
Greek letters α moles of oxygen required to oxidize 1 mol of n-butane to MA β moles of oxygen required to oxidize 1 mol of n-butane to COx γ moles of oxygen required to oxidize 1 mol of MA to Cox λ conductivity (mS/cm)
Although the lattice oxygen and oxygen transfer from the catalyst to reagent are key elements in determining catalyst activity and selectivity, the nature of their role is far from clear. Most studies agree that (VO)2 P2 O7 is the predominant active phase but that three oxidation states have been identified under reaction conditions—V3+ , V4+ , and V5+ . Furthermore, Bej and Rao [12] identified two different active sites—selective to maleic anhydride (MA) and combustion—using selective poisoning studies and studied the kinetics using an integral, differential and recycle reactor configurations [13]. Mills et al. [9] propose that subsurface lattice diffusion contribute to activity, whereas Schuurman and Gleaves [8] suggest that the separate oxidation step decreases the activation energy resulting in increased activity. In conventional fixed bed technology, the local environment around the catalyst is oxygen rich; in the alternative reactor technologies, the local environment can vary between fuel rich and fuel lean. A further complication in the new reactor types is that the catalyst is periodically cycled between the two environments either in a controlled fashion as in the DuPont process or randomly like in a fluid bed equipped with butane spargers. Buchanan and Sundaresan [14] have shown increased reaction rates when the feed concentration is switched from a high oxygen-to-butane ratio to a low ratio. This would indicate that the behavior of the catalyst in transient conditions is different than under steady-state conditions. What is required to quantify these observations is an independent measure of both catalyst reduction and oxidation at industrial operating conditions. In this report, we describe a fluid bed reactor concept in which the gas feed may be cycled from a net reducing to a net oxidizing environment and not the catalyst. This configuration allows us to control the reaction conditions precisely. It is also equipped to run with a constant
feed gas composition and thus can simulate many commercial configurations. We show that oxidizing the catalyst surface indeed increases reaction rates but that the time scales are too long to be of commercial relevance. Short time scale cycling studies show that the catalyst surface may chemisorb carbonaceous species and the quantity depends on the exit oxygen concentration. This would imply the V3+ species may only exist as a carbon metal complex. Finally, we propose a simplified kinetic expression to describe the experimental observations. The model agrees well with the transient data but more development is required to include the trends that have been identified in the published literature.
2. Experimental All experiments were carried out in a 41 mm diameter and 790 mm height Hastelloy C-276 vessel, as shown in Fig. 1. It was designed to operate at pressures as high as 51 bar but tests were carried out below 5 bar and for safety reasons a 6.8 bar rupture disk was mounted at the top. A 10 point thermocouple monitored the temperature profile inside the reactor, which was loaded with 200–400 g of catalyst. The axial temperature gradient was nearly isothermal and this was largely due to the excellent heat transfer characteristics of an externally heated sand bath in which the reactor was immersed. Inlet gases passed through a tube coiled around the outer surface of the reactor and entered at the reaction temperature. A sintered metal frit separated the plenum and catalyst and ensured that the gas entered the fluid bed evenly across the diameter. Most tests were conducted in the bubbling fluid bed regime but low gas velocities were tested that simulated the fixed bed hydrodynamic regime. A flanged upper section minimized solids entrainment and a sintered metal filter retained all catalyst in the reactor.
Fig. 1. Reactor, feed and effluent.
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2.1. Gas manifold and analytical instrumentation Fig. 2 illustrates the feed reactor manifold as well as the effluent analysis configuration. Both nitrogen and air flow rates were controlled with Tylan FC-2900V-4S mass flow controllers. Butane was fed to the reactor through a heated Brooks 5850S mass flow controller. The feed manifold consisted of two multi-port valves; the four-way valve allowed either a mixture of butane/N2 to go to the reactor or only N2 ; the eight-port valve was downstream of the four-way valve and was connected to the air manifold. The eight-way valve had two effluent streams; one entered the reactor and the other vented to atmosphere. Both lines were maintained at the same pressure to avoid pressure spikes or flow transients when the feeds were cycled. To prevent product condensation at the reactor effluent, the line was maintained at 200 ◦ C with an electrical heat tape. Liquid from the quench was pumped to this line immediately after the heat tape to minimize solids build up in the line but also to improve the scrubbing efficiency of the quench. The quench vessel was made of quartz and the accumulated acids were sampled frequently and analyzed off-line by HPLC (Hewlett Packard 1050) equipped with a variable wavelength UV detector. In addition to the frequent sample analysis, the quench liquid conductivity (referred as λ) was measured continuously to monitor acid production rate. The
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acid concentration and conductivity are related by a second order polynomial function (MA = 37.308λ2 + 349.78λ, being maleic concentration in ppm and conductivity in mS/cm) determined experimentally. At the end of each experiment, the absorber was washed three times. These samples were also analyzed by HPLC and included in the mass balance calculations. A slip stream of the effluent gas passed through an ice trap and the non-condensables were measured with various on-line analytical instruments. An HP 5890 GC equipped with both a flame ionization detector and a thermal conductivity detector measured the concentrations of CO, CO2 , O2 , N2 and C4 H10 . Oxygen was also monitored on-line (in real time) with a Siemens Oxymat 5F paramagnetic analyzer that had a response time of 1 s for 80% of full scale. ABB model 501B IR analyzers monitored the both CO and C4 H10 gases (on-line in real time). 2.2. Catalyst VPO catalyst used in this study was prepared in an “organic medium” and encapsulated in a porous silica shell in order to make it resistant to forces and abrasion typical of fluid bed reactors. The catalyst was equilibrated during 400 h on stream before the experiments were run. Catalyst activity and inventory were constant throughout the several months required to complete the study.
Fig. 2. Feed gas configuration and analytical equipment.
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Particle characteristics have been described by Lorences [15]. The mean particle diameter was 64 m with 20% less than 44 m. The particle density was 1890 kg/m3 and the minimum fluidization velocity was 5.6 mm/s. 2.3. Experimental procedure The gas manifold design allows one to simulate many different configurations and examine transient behavior under industrially relevant reactor conditions. In a previous communication [16], we described steady-state experiments and explored the effect of butane and oxygen concentration on the production of by-product acids, of which acetic and acrylic predominate. Herein we concentrate on cycling feeds between net reducing and oxidizing environments; a mixture of butane/nitrogen/air is fed to the reactor for a specified period of time; the reactor is purged for 10 min with nitrogen and then an oxygen/nitrogen mixture is fed. The experimental conditions are shown in Table 1 and the procedure followed was: 1. Catalyst reduction during 40 min varying the butane/oxygen concentration in the feed from 2 to 9 and 0 to 10%, respectively, at temperatures of 380 and 410 ◦ C. 2. Nitrogen purge for 10 min. 3. Catalyst oxidation for 40 min with oxygen varying between 2 and 20% and pressures as high as 5 abs bar. 4. Nitrogen purge for 10 min. 5. Second oxygen cycle for 30 min at the same conditions as Step 3. This served as a reference to calculate oxygen uptake.
3. Results and discussion 3.1. Catalyst reduction by butane Several authors have demonstrated that VPO catalyst reacts butane under anaerobic conditions for periods of time on the order of minutes [17]. DuPont’s transport bed process [3] relies, in part, on lattice oxygen from the catalyst to react butane. In this process, a reducing gaseous environment rich in butane is fed to the bottom of a transport bed reactor. The gas carries the solids to the top of the reactor into a cyclone. The effluent gases exit through the top of the cyclone and the solids fall into a stripper. The catalyst is then carried down through a standpipe into a regenerator with a high partial pressure of oxygen. The reoxidized catalyst is returned to the bottom of the transport bed after the regenerator. In this study, the first set of experiments (#1–#3 in Table 1) were conducted analogously to DuPont process, in that we cycled the catalyst between oxidizing and reducing environments. Before the initial reduction step (Step 1), catalyst was oxidized for an extended period of time; the reduction was conducted with different butane/N2 mixtures and lasted 40 min. The reactor was purged by a nitrogen stream for 10 min and then air was fed to it for 40 min. (Oxidation and reduction steps are much shorter in the commercial process!) The nitrogen purge and air treatment was then repeated. The time delay between switching the valves and the gas phase analytical measurements was on the order of 1 min. It is lower for the conductivity measurement because the gas hold-up between the reactor and quench is much smaller than between the reactor and gas phase detectors.
Table 1 List of operating conditions under cycling mode Experiment
T (◦ C)
Step 1 (balance N2 )
Step 2
Step 3
Step 4
Step 5
#1 #2 #3 #4 #5 #6 #7 #8 #9 #10 #11 #12 #13 #14 #15 #16 #17 #18 #19 #20 #21 #22 #23
380 380 380 380 380 380 380 380 380 380 380 380 410 410 410 410 380 380 380 380 380 380 380
2% 5% 9% 2% 5% 9% 2% 5% 9% 2% 5% 9% 5% 5% 5% 5% 5% 5% 5% 5% 5% 5% 5%
N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2
Air Air Air Air Air Air Air Air Air Air Air Air Air Air Air Air 2% O2 /N2 5% O2 /N2 10% O2 /N2 5% O2 /N2 (3 bar) Air (3 bar) 5% O2 /N2 (5 bar) Air (5 bar)
N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2
Air Air Air Air Air Air Air Air Air Air Air Air Air Air Air Air 2% O2 /N2 5% O2 /N2 10% O2 /N2 5% O2 /N2 (3 bar) Air (3 bar) 5% O2 /N2 (5 bar) Air (5 bar)
C4 C4 C4 C4 /2% O2 C4 /2% O2 C4 /2% O2 C4 /4% O2 C4 /4% O2 C4 /4% O2 C4 /10% O2 C4 /10% O2 C4 /10% O2 C4 C4 /2% O2 C4 /4% O2 C4 /10% O2 C4 C4 C4 C4 C4 C4 C4
(3 bar) (3 bar) (5 bar) (5 bar)
(3 bar) (3 bar) (5 bar) (5 bar)
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Fig. 3. Butane oxidation by lattice oxygen. Evolution of the reaction products during the different reduction and oxidation steps at 380 ◦ C.
Three different concentrations of butane (nominally 2, 5 and 9%) using nitrogen as a diluent were tested. The temperature was set at 380 ◦ C. Fig. 3 shows the butane, carbon monoxide, maleic acid (measured as conductivity, λ) and oxygen concentrations as a function of time. During the initial reduction step (Step 1) the conductivity of the liquid from the quench, which relates to maleic anhydride production—increases sharply for the first several minutes then levels off. CO formation reaches a maximum of 1.2% at 4 min before tailing off to zero. The CO2 was analyzed by GC and its concentration was about the same as CO in the first trace at 2 min but was undetectable after the second trace. The evolution of these products was entirely due to the catalyst lattice oxygen since oxygen was not co-fed. The maximum CO concentration peaked at 1.5% with 5% butane and 2.5% with 9% butane. As shown in Fig. 3, CO is detected for up to 14 min while in the experiment with 9% butane, the peak dropped to zero after only 8 min. Peak height depended on butane concentration but the total CO detected for each of the experiments was nominally about the same, which implies that the available lattice oxygen is independent of the butane concentration, but the reaction rate increases with butane partial pressure. We were surprised to detect a significant quantity of CO during the first oxidation step (Step 3). Moreover, we found the concentration increased with the butane feed concentration and reached a maximum of almost 1%. The shape of the curve is similar to the first CO curve during the catalyst reduction step. However, CO evolution during the oxidation step increases with butane concentration during the reduction step; peaks were higher and the tails of the curves were longer. Emig et al. [18] also detected small amounts of CO and CO2 during an oxidation pulses whereas Wang and Barteau [17,19] claimed little carbon oxide evolves during an oxidation step after reducing the catalyst. Traces of butane are detected for as much as 20 min after the beginning of the nitrogen purge (Step 2). Based on blank experiments with oxygen, we expected a complete purge after 10 min and this expectation is supported by the second
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Fig. 4. Oxygen breakthrough during the first oxidation step depending on the butane concentration present in the reduction step. (T = 380 ◦ C; oxidation agent: air).
oxidation step: 10 min after the second nitrogen purge, the oxygen concentration reaches zero. Despite the low levels of butane detected during VPO oxidation, we concluded that the CO peak is due to oxidation of carbonaceous species that built up on the catalyst surface. There is a marginal rise in the quench conductivity shown above but the rise was barely detectable for the experiments run with 5% butane and 9% butane. Fig. 4 compares the oxygen breakthrough of each of the three experiments during the first oxidation step (Step 3). The curve shifts to the right with increasing butane feed concentration, indicating increasing oxygen conversion. Coincidentally, CO evolution is highest for the case of the highest butane feed concentration. Maleic anhydride production is negligible during Step 3 with 2% butane in the feed and zero for Experiments 2 and 3 with 5 and 9% butane, respectively. The data support the hypothesis that carbon deposition/adsorption depends on the partial pressure of butane in the feed. Poor stripping efficiency during the nitrogen purge could also explain the trend in the data. However, if the CO evolution and oxygen conversion were due to residual butane in the gas phase, we would also have expected to see maleic acid in the quench, which was not the case. In a second series of experiments, we examined the nature of the lattice oxygen by reacting butane with VPO that was previously oxidized for different periods of time—1 and 48 h. Immediately after the air treatment, the reactor was purged with nitrogen and then a stream containing 2% butane/98% nitrogen was fed to the reactor for 40 min. The reactor was purged again with nitrogen for 10 min followed by air for 40 min. The conductivity of the quench was monitored on-line and the data are shown in Fig. 5. After the 48 h exposure, maleic anhydride production (conductivity) is twice as high as it is after the 1 h exposure. Additionally, consistent with Fig. 3, there is a slight rise in conductivity during the air treatment after the 10 min nitrogen purge. A third test was conducted in which we oxidized the catalyst in air for 48 h then reduced the catalyst with 9% butane
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Fig. 5. Lattice oxygen availability after different oxidation pre-treatments.
instead of 2%. The maleic anhydride evolution with respect to time, as shown in Fig. 5, follows the same trend as when we fed only 2% butane. This would imply oxygen reaction is rate limiting for both cases and that butane is in excess with only a feed concentration of 2%. Several studies concur that the reaction kinetics are highly dependent upon the reoxidation cycle time. Mills et al. [9] suggested that oxygen participates in the reaction as a subsurface lattice via solid-state diffusion coupled with surface oxygen. This hypothesis is consistent with the data showing an increase in production by increasing the oxygen exposure. On the other hand, Schuurman and Gleaves [8] proposed that the activation energy depends on the surface oxidation state and the activation energy was lower for more highly oxidized VPO. The data reported in Fig. 5 would not seem to support this concept in that the rate of formation of maleic anhydride is about the same for both experiments; lattice oxygen contribution (LOC) is higher and not the reaction rate. Emig et al. [18] also observed that strongly oxidized catalyst showed high activity and selectivity. Their data showed that the amount of oxygen available for the reaction increased from 60 to 190 mmol O2 /kgcat when the regeneration time changed from 30 s to 18 h, while the amount of maleic anhydride formed increased from 0.3–0.7 to 2–4 gmac /kgcat . The activation energy was found to be lower for the more highly oxidized samples, in agreement with Schuurman and Gleaves [8] observation. Practical commercial application of the subsurface lattice or lowered activation energies would appear to be prohibitive considering that extensive residence times are required. Residence times much greater than 2 min would have significant negative economic impact. 3.2. Catalyst reduction by butane with oxygen co-feed Both Emig et al. [18] and Schuurman and Gleaves [8] reported that, together with the lower activation energies, formation of MA and desorption is faster on a more oxi-
Fig. 6. Conductivity and oxygen concentration vs. time after exposition to different environments. Operating conditions: 5% C4 , 10% O2 at 410 ◦ C.
dized catalyst surface. However, this effect was observed under anaerobic conditions. To examine the effect under butane/oxygen co-feed conditions, we first allowed the catalyst to reach steady state with 5% butane and 10% oxygen, treated the catalyst for different amounts of time with nitrogen or air and then resumed feeding 5% butane and 10% oxygen. We tested four different “stripping/regeneration” environments: (a) 1 h of air, (b) 1 h of N2 , (c) 10 min of air, and (d) 10 min of N2 . Fig. 6 shows the evolution of conductivity and oxygen for each of the four stripping/regeneration treatments. After 10 min or 1 h of nitrogen, the oxygen breakthrough (during the reduction step) increases rapidly during the first couple of minutes then levels off to over 2%. After the air treatment, the oxygen concentration overshoots the steady-state value by a considerable amount. The overshoot is greater after the 1 h treatment and reaches 4%, which is double the steady-state value. The maleic anhydride production rate is essentially the same for each of the four “stripping/regeneration” treatments: oxygen breakthrough depends on the pre-treatment conditions while maleic anhydride evolution is independent of the pre-treatment. This phenomenon would imply that VPO reduction by butane is the rate limiting step under these conditions. As long as the partial pressure of oxygen is sufficiently high, the butane concentration limits the maleic anhydride reaction rate.
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Fig. 7. Evolution of the different products: (a) C4 H10 , (b) CO, (c) maleic acid (conductivity), (d) O2 depending on the butane concentration during the reduction step. Operating conditions: 10% O2 at 380 ◦ C.
Under anaerobic reducing conditions, lattice oxygen is the limiting reagent, thus butane concentration has a limited impact on maleic yield. However, when oxygen is co-fed together with butane, reaction rates, maleic production and lattice oxygen concentrations may all depend on both the reoxidation and reduction step. Experiments #4–#12 explore the effect of butane and oxygen partial pressure during the reduction step on the catalyst reoxidation. Fig. 7 illustrates the product distribution as a function of the inlet butane concentration—2, 5 and 9%—with 10% oxygen (Experiments #10–#12). In this case, both the CO and the maleic acid concentration (conductivity) increase with increasing butane in the feed. (Note that the peak in the butane and oxygen concentration after 40 min corresponds to the feed gas analysis.) Whereas CO breakthrough during the oxidation step was evident at all anaerobic conditions, CO is only detected for the case with 9% butane and 10% oxygen. As a consequence, oxygen uptake during catalyst oxidation is about the same for all three cases. 3.3. Catalyst oxidation Several tests were conducted to evaluate the VPO oxidation rate. In the last series, the catalyst was reduced by a stream of 5% butane in nitrogen at 380 ◦ C then treated
with different oxygen concentrations—2, 5, 10 and 21% (#2, #17–#19 in Table 1)—and pressures of 1, 3 and 5 bar (#2, #18, #20–#23 in Table 1). Oxygen breakthrough was fastest at the higher pressures. The rate of CO evolution was highly dependent on the oxygen partial pressure; the maximum value was higher at higher oxygen partial pressures; at the lower oxygen partial pressures, the tail was longer and the peak time appeared several minutes later. Total CO evolution was calculated by integrating the area underneath the curve. Fig. 8(a) shows the total oxygen consumption—lattice oxygen contribution—is estimated by subtracting the area under the oxygen breakthrough curve in Step 5 from Step 3. The second part of the figure shows an estimate of the quantity of oxygen required to combust the adsorbed carbonaceous species. The calculation was based on the CO evolution during step and the ratio CO/CO2 given by the GC analysis (only CO is shown in the figure). We assumed that butane was adsorbed on the surface, although butene or butadiene is a more likely species. Fig. 8 is consistent with expectations and shows that carbon adsorption increases with the severity of the reducing conditions. More butane is deposited on the surface in an anaerobic atmosphere. Under highly oxidizing conditions the surface is free of carbon species. However, even when the feed contains 10% oxygen, carbon deposits on the surface at high butane partial pressures.
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the both butane conversion and MA selectivity. Although the model adequately characterized the data set, it was fundamentally flawed because it had no provision to allow for carbon–metal complexes that form under reducing conditions. The goodness-of-fit between an experimental data set and a theoretical model is a necessary consideration with respect to model discrimination; however, it is often not sufficient. The final step in the modeling exercise was to include the carbon–metal complex but again to attempt to minimize the degrees of freedom by limiting the number of possible reactions. The model assumed three types of sites: an active site, V5+ , that either reacted with butane to form maleic anhydride or it combusted the maleic anhydride, as shown by Eqs. (1) and (2); a reduced site, V4+ , resulting from the selective oxidation of butane or combustion of maleic anhydride; and finally, a carbon–metal complex, VC4 , formed when butane chemisorbs on to a V4+ site, Eq. (3). When the gas phase oxygen concentration is low, the number of V4+ sites increases and the surface coverage of the carbon–metal complex also increases. Eq. (3) assumes that this complex only reacts with gas phase oxygen. Therefore, a highly reduced catalyst will result in a large fraction of the so-called VC4 phase. Eventually, equilibrium is reached between the gas composition and the three types of sites. This model predicts that maleic selectivity depends to a large extent on maintaining the catalyst surface oxidized: k1
Fig. 8. Total oxygen uptake (a) and oxygen consumed to form CO (b) as a function of the butane and oxygen concentration in the reduction step: (䊊) 0% O2 , (䊐) 2% O2 , () 4% O2 , (䉫) 10% O2 .
C4 H10 + αV5+ − → MA + αV4+ + 4H2 O, r1 = k1 [C4 H10 ][V5+ ]
(1)
k2
As shown in Fig. 8, the lattice oxygen contribution follows the identical trend as for the carbon deposition: highly reducing anaerobic conditions extract more oxygen from the catalyst compared to oxidizing conditions. Under anaerobic conditions with 9% butane, the LOC number exceeds 140 mm O2 /kg VPO and it is almost an order of magnitude lower with a 10% O2 feed concentration.
→ 4COx + βV4+ + H2 O, MA + βV5+ − r2 = k2 [MA][V5+ ] k3
(2)
→ 4VC4 , C4 H10 + 4V4+ − k4
O2 + V4+ − → V5+ ,
r3 = k3 [C4 H10 ][V4+ ]
r4 = k4 [O2 ][V4+ ]
(3) (4)
k5
γO2 + 4VC4 − → 4COx + 4V4+ + 5H2 O, 4. Kinetic modeling
r5 = k5 [O2 ][VC4 ]
In a previous publication, Lorences et al. [16] generated a series of over 120 steady-state experiments with a broad range of feed concentrations, temperatures and contact times. Several published models were compared to the experimental data and agreement between butane conversion and experimental observations were very good. However, the models poorly characterized maleic anhydride selectivity and a new kinetic expression was developed. The premise of the model was that the number of independent reactions should be minimized but that there were two active sites—one selective to maleic anhydride formation and the second non-selective. They successfully fit the model to
The reaction rate equations are re-written in terms of the normalized oxidation state by dividing by the total number of active sites —VT — and assuming that the total was constant (Eqs. (6)–(10)). The sum of the fraction of the sites in each “oxidation” state—V5+ , V4+ and VC4 —equaled 1. Some of the experimental evidence might suggest that the total number of sites varies with the oxygen treatment. For example, Fig. 5 shows that after 48 h of exposure to air, the maleic anhydride yield is higher compared to a 1 h treatment. On the other hand in the case of much shorter time scales and aerobic conditions, oxygen appears to react only with sites reduced by butane, as shown in Fig. 6: catalyst pre-treated
(5)
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with oxygen did not increase the maleic anhydride formation rate and the higher initial oxidation state resulted an overshoot in the oxygen concentration that corresponds to higher initial oxidation state: r1 =
k1 [C4 H10 ][VT − V4+ − VC4 ] VT
(6)
r2 =
k2 [MA][VT − V4+ − VC4 ] VT
(7)
r3 =
k3 [C4 H10 ][V4+ ] VT
(8)
r4 =
k4 [O2 ][V4+ ] VT
(9)
r5 =
k5 [O2 ][VC4 ] VT
(10)
4.1. Hydrodynamic modeling Most experimental treatments of butane oxidation are conducted in the fixed bed fluid regime: reactor modeling is simplified because the gas phase is in plug flow. On the other hand, these reactors can often have temperature and pressure gradients as well as radial or inter-particle gradients. Further complications arise from gas to particle diffusion rates that may cause localized hot spots. Gas phase modeling of fluid beds is generally more difficult because bubbles form at the grid and rise through the bed at velocities significantly higher than the superficial gas velocity. Consequently, both contact efficiency and contact time are reduced. However, the solids phase is backmixed to a high degree and the reactor is generally isothermal, which simplifies the modeling. In this work, we operated at relatively low gas velocities. At velocities in the vicinity of Umf , the gas is in plug flow and the reactor operates in the regime. Deviation from plug flow increases with increasing gas velocity. To evaluate the hydrodynamics, we conducted a series of inert gas step-up response experiments and then characterized the data with a CSTR in series model [16]. Agreement between the experimental RTD data and a model assuming five CSTRs in series was excellent. 4.2. Parameter estimation Fig. 9 illustrates the experimental data for the case where we fed 2% butane and 4% oxygen to catalyst that had been oxidized for a period of at least 50 min. The COx concentration is based on the IR response of the CO analyzer plus the CO/CO2 ratio measured by the GC. The liquid quench conductivity began to rise after 1 and 2 min the slope of the line was constant. This could imply that during the first minute of reaction, maleic anhydride was combusting on the oxidized catalyst surface but was constant thereafter. The butane concentration was zero for the first minute then began to increase modestly thereafter. It reached its steady-state value
Fig. 9. Product gas evolution in the initial reduction step (experimental and model predictions): Feed composition—2% butane, 4% oxygen, T = 380 ◦ C.
of 1.5% after about 10 min. Both oxygen and CO signals are recorded after 1 min, while the butane signal was still zero. Their concentrations significantly overshoot the steady-state value and reach a maximum of 2.7% for O2 and 4.5% for CO + CO2 . As with the butane, they reach the steady-state values at about 10 min. We combined the hydrodynamic model of the fluid bed reactor together with the kinetic model outlined in Eqs. (6)–(10) to simulate the catalytic performance of the VPO. In the first simulation, we adopted the kinetic rate constants derived from the steady-state experiments of Lorences et al. [16]. Agreement between the model predictions and experimental data was poor. The model did not predict the overshoot of the CO (and CO2 ) and O2 concentrations shown in Fig. 9. Rather, it predicted that the concentrations increased asymptotically from zero to the steady-state values. Clearly the pre-oxidation step has a substantial effect on catalytic performance. Instead of adopting the rate constants from the steady-state experiments, we optimized the rate constants that fit the experimental data. The most difficult part of this process was to adequately predict the overshoot, while at the same time match the steady-state value beyond 10 min. Fig. 9 compares the model predictions against the experimental data and the agreement over the entire range is excellent. However, beyond 14 min the model predictions of the oxygen appear to begin to deviate from the experimental data somewhat.
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Table 2 Kinetic rate constants and activation energies i
k0i (s−1 ) (transient)
k0i (s−1 ) (steady state)
Eai (kcal/mol)
1 2 3 4 5
0.91 2.5 0.13 0.28 0.34
0.33 0.37 0.05 0.57 0.06
15 14 17 9 28
Table 2 compares the best fit kinetic rate constants of the transient experiment with the values derived from the steady-state experiments. (Note that the same activation energies were used for both.) In the steady-state experiments, the rate constants for butane to maleic anhydride (k01 ) and maleic anhydride to COx (k02 ) differ by about 20%. The non-selective reaction rate is greater than 2.5× the selective reaction rate for the transient experiment. In general, the pre-exponential factors from the transient experiment are from three to five times the value of the best fit parameters from the steady-state experiments except for catalyst reoxidation from V4+ to V5+ (k04 ). In this case, the steady-state value is double the transient value. Presumably, one contribution to the large differences in the best fit parameters of the two data sets is due to the initial oxidation state of the catalyst. The data show that the available lattice (or bulk) oxygen for reaction may depend on the pre-treatment step. However, extended catalytic oxidation would appear to increase the rate of maleic oxidation but not the selective oxidation of butane to maleic anhydride. Further modeling efforts are required to take into account the relationship between oxygen pre-treatments (regeneration) and the available lattice oxygen for reaction with butane or maleic anhydride. This exercise clearly shows the value of employing alternative reactor types when modeling reaction kinetics. The steady-state experiments provide a data base with which to gauge the effect of temperature, composition and contact time on reaction rates and the transient experiments can help substantiate the reaction mechanism.
5. Conclusions We introduce a novel reactor configuration based on operating a fluid bed at low gas velocities such that gas bypassing is minimized but where solids backmixing is sufficiently high to ensure a uniform temperature distribution. The reactor is coupled with a means of quantitatively monitoring the primary and by-product acids on-line and in real time by measuring the conductivity of condensable gases in a quench and withdrawing frequent samples for subsequent HPLC analysis. The on-line gas analysis demonstrated that carbon is strongly adsorbed to the surface under net reducing conditions. Butane reacts with lattice oxygen and the kinetic model suggests that the same vanadium species also combusts
maleic anhydride. The series reaction to combustion products is the prime determinant of selectivity. Under fuel lean, oxygen rich conditions, reaction rates primarily depend on butane concentration. However, co-feed butane/oxygen conditions where the partial pressure of oxygen at the exit of the reactor approaches zero (in the vicinity of about 1%), carbon builds up on reduced sites—V4+ —to form an adsorbed species. This species reacts with molecular oxygen to form CO/CO2 . Although the same kinetic model can fit both transient and steady-state data, the differences in the optimized rate parameters are substantial. Reaction rates of a previously oxidized catalyst are significantly higher than a catalyst operating under steady-state conditions. This enhanced activity is most likely due to a higher density of oxygen on the surface or subsurface layers. Under extended oxygen exposure times, maleic anhydride production rates can be increased by a factor of 2. However, more detailed modeling is required to identify the nature of the non-selective reactions of a highly oxidized VPO catalyst.
Acknowledgements The authors gratefully acknowledge the financial and technical support of DuPont Ibérica S.A. and the assistance of Mónica Alonso Carreño during the experimental work.
References [1] A. Escardino, C. Sola, F. Ruiz, An. Quim. 69 (1973) 385. [2] A. Escardino, C. Sola, F. Ruiz, An. Quim. 69 (1973) 1157. [3] R.M. Contractor, D.I. Garnett, H.S. Horowitz, H.E. Bergna, G.S. Patience, J.T. Schwartz, G.M. Sisler, in: V.C. Corberán, S.V. Bellón (Eds.), New Developments in Selective Oxidation, Elsevier, Amsterdam, 1994, p. 233. [4] G. Parkinson, R. D’Aquino, Chem. Eng. (1999) 28. [5] M. Alonso, M.J. Lorences, M.P. Pina, G.S. Patience, Catal. Today 67 (2001) 151. [6] R. Mallada, S. Sajip, C.J. Kiely, M. Menéndez, J. Santamar´ıa, J. Catal. 1–7 (2000) 196. [7] K.G. Golbig, J. Werther, Chem. Eng. Sci. 52 (4) (1997) 583. [8] Y. Schuurman, J.T. Gleaves, Catal. Today 33 (1997) 25. [9] P.L. Mills, H.T. Randall, J.S. McCracken, Chem. Eng. Sci. 54 (1999) 3709. [10] R. Mallada, M. Menéndez, J. Santamar´ıa, Catal. Today 56 (2000) 191. [11] D. Wang, M.A. Barteau, Appl. Catal. A 223 (2002) 205. [12] S.K. Bej, M.S. Rao, Ind. Eng. Chem. Res. 30 (1992) 1824. [13] S.K. Bej, M.S. Rao, Ind. Eng. Chem. Res. 30 (1992) 1829. [14] J.S. Buchanan, S. Sundaresan, Appl. Catal. 26 (1986) 211. [15] M.J. Lorences, Ph.D. Thesis, University of Oviedo, Oviedo, 2000. [16] M.J. Lorences, G.S. Patience, F.V. D´ıez, J. Coca, I&EC Research 42 (26) (2003) 6730. [17] D. Wang, M.A. Barteau, J. Catal. 197 (2001) 17. [18] G. Emig, K. Uihlein, C.J. Häcker, in: V.C. Corberán, S.V. Bellón (Eds.), New Developments in Selective Oxidation II, Elsevier, Amsterdam, 1994, p. 243. [19] D. Wang, H.H. Kung, M.A. Barteau, Appl. Catal. A 201 (2000) 203.
Transient n-butane partial oxidation kinetics over VPO Mar´ıa J. Lorences a,1 , Gregory S. Patience b,∗ , Fernando V. D´ıez a , José Coca a a
Department of Chemical Engineering and Environmental Technology, University of Oviedo, Julián Claver´ıa s/n, Oviedo, Asturias 33209, Spain b DuPont Ibérica, S.A., Tamón, Avilés 33469, Spain Received 26 June 2003; received in revised form 5 December 2003; accepted 8 December 2003
Abstract A novel reactor configuration is proposed to study the kinetics of heterogeneous catalytic reactions. It is based on the bubbling fluidization regime; gas velocities are sufficiently low to minimize bypassing but high enough to ensure solids backmixing and a uniform temperature gradient. The reactor is equipped with a gas manifold designed to operate the reactor either under steady-state feed conditions or with cyclical feeds. By cycling VPO catalyst between a reducing environment with butane and an oxidizing environment, we quantify the lattice oxygen contribution to the reaction. Furthermore, we show that carbon can build up on the VPO surface and this phenomenon increases with an increase in the severity of the reducing environment. We propose a simple kinetic model involving three sites—V5+ , V4+ and VC4 . The model adequately characterizes the transient partial oxidation of butane but further research is required to create a generalized model that fits all reaction environments. © 2003 Elsevier B.V. All rights reserved. Keywords: Transient kinetics; VPO; Butane oxidation; Maleic anhydride; Fluid bed
1. Introduction Kinetic modeling of butane partial oxidation to maleic anhydride spans over 30 years [1,2]. There is a general consensus that the reaction is near first order in butane (under highly oxidizing conditions) and that most experimental evidence supports a Mars van Krevelen mechanism. For conventional fixed beds—fuel lean and oxygen rich—this knowledge may be sufficient for reactor design and process control. However, several alternative reactor types are commercial and new concepts under development that may either operate anaerobically or fuel rich, oxygen lean conditions. For example, butane concentrations in some fluid beds may reach as high as 5%. To avoid the flammability regime in the plenum, where there are no solids to quench free radicals, 100% butane may be sparged directly into the reactor. Therefore, on a microscopic time scale, catalyst would experience near anaerobic conditions at the sparger ∗ Corresponding author. Current addresses: INVISTA (International) S.A., European Technical Centre, 146, route Nant-d’Avril, CH-1217 Meyrin, Geneva, Switzerland. Tel.: +41-22-717-6996; fax: +41-22-717-6868. E-mail address: [email protected] (G.S. Patience). 1 HALDOR TOPSØE A/S, Nymøllevej 55, 2800 Lyngby, Denmark.
0926-860X/$ – see front matter © 2003 Elsevier B.V. All rights reserved. doi:10.1016/j.apcata.2003.12.023
tips but fuel lean conditions at the grid. DuPont commercialized a circulating fluidized bed concept in which catalyst is shuttled between a net oxidizing and a net reducing environment [3]. High butane feed concentrations are not restricted to fluid bed regimes. Parkinson and D’Aquino [4] report that Pantochim operates a fixed bed under a net reducing environment and sparges oxygen into their butane feed stream. Finally, membrane reactors are under development [5,6], in which high butane concentrations are fed to a fixed bed and oxygen diffuses across a porous “membrane” along the length. Clearly, for design purposes, a more fundamental understanding of the prevailing kinetics would be desirable to reduce the risk at the design stage and perhaps help optimize reactor performance when operational. Several novel experimental programs and analytical techniques have been developed to explore these revolutionary commercial designs. For example, Golbig and Werther [7] built a lab scale circulating fluidized bed with a 21 mm diameter riser. Schuurman and Gleaves [8] and Mills et al. [9] have adapted TAP reactors to measure reactor transients at elevated pressures. Mallada et al. [10] were among the first to explore the effect of high butane concentrations on reaction kinetics in lab scale fixed bed reactors. Recently, Wang and Barteau [11] developed a novel oscillating microbalance reactor to study oxidation and reduction steps independently.
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Nomenclature Eai ki k0i LOC MA ri Vi+ VPO
activation energy (kcal mol−1 ) reaction rate constant (s−1 ) pre-exponential factor (s−1 ) lattice oxygen contribution (mmol O2 /kg VPO) maleic anhydride concentration (mol m−3 ) reaction rate (mol s−1 m−3 ) oxidation state of the catalyst vanadium-phosphorous-oxide catalysts
Greek letters α moles of oxygen required to oxidize 1 mol of n-butane to MA β moles of oxygen required to oxidize 1 mol of n-butane to COx γ moles of oxygen required to oxidize 1 mol of MA to Cox λ conductivity (mS/cm)
Although the lattice oxygen and oxygen transfer from the catalyst to reagent are key elements in determining catalyst activity and selectivity, the nature of their role is far from clear. Most studies agree that (VO)2 P2 O7 is the predominant active phase but that three oxidation states have been identified under reaction conditions—V3+ , V4+ , and V5+ . Furthermore, Bej and Rao [12] identified two different active sites—selective to maleic anhydride (MA) and combustion—using selective poisoning studies and studied the kinetics using an integral, differential and recycle reactor configurations [13]. Mills et al. [9] propose that subsurface lattice diffusion contribute to activity, whereas Schuurman and Gleaves [8] suggest that the separate oxidation step decreases the activation energy resulting in increased activity. In conventional fixed bed technology, the local environment around the catalyst is oxygen rich; in the alternative reactor technologies, the local environment can vary between fuel rich and fuel lean. A further complication in the new reactor types is that the catalyst is periodically cycled between the two environments either in a controlled fashion as in the DuPont process or randomly like in a fluid bed equipped with butane spargers. Buchanan and Sundaresan [14] have shown increased reaction rates when the feed concentration is switched from a high oxygen-to-butane ratio to a low ratio. This would indicate that the behavior of the catalyst in transient conditions is different than under steady-state conditions. What is required to quantify these observations is an independent measure of both catalyst reduction and oxidation at industrial operating conditions. In this report, we describe a fluid bed reactor concept in which the gas feed may be cycled from a net reducing to a net oxidizing environment and not the catalyst. This configuration allows us to control the reaction conditions precisely. It is also equipped to run with a constant
feed gas composition and thus can simulate many commercial configurations. We show that oxidizing the catalyst surface indeed increases reaction rates but that the time scales are too long to be of commercial relevance. Short time scale cycling studies show that the catalyst surface may chemisorb carbonaceous species and the quantity depends on the exit oxygen concentration. This would imply the V3+ species may only exist as a carbon metal complex. Finally, we propose a simplified kinetic expression to describe the experimental observations. The model agrees well with the transient data but more development is required to include the trends that have been identified in the published literature.
2. Experimental All experiments were carried out in a 41 mm diameter and 790 mm height Hastelloy C-276 vessel, as shown in Fig. 1. It was designed to operate at pressures as high as 51 bar but tests were carried out below 5 bar and for safety reasons a 6.8 bar rupture disk was mounted at the top. A 10 point thermocouple monitored the temperature profile inside the reactor, which was loaded with 200–400 g of catalyst. The axial temperature gradient was nearly isothermal and this was largely due to the excellent heat transfer characteristics of an externally heated sand bath in which the reactor was immersed. Inlet gases passed through a tube coiled around the outer surface of the reactor and entered at the reaction temperature. A sintered metal frit separated the plenum and catalyst and ensured that the gas entered the fluid bed evenly across the diameter. Most tests were conducted in the bubbling fluid bed regime but low gas velocities were tested that simulated the fixed bed hydrodynamic regime. A flanged upper section minimized solids entrainment and a sintered metal filter retained all catalyst in the reactor.
Fig. 1. Reactor, feed and effluent.
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2.1. Gas manifold and analytical instrumentation Fig. 2 illustrates the feed reactor manifold as well as the effluent analysis configuration. Both nitrogen and air flow rates were controlled with Tylan FC-2900V-4S mass flow controllers. Butane was fed to the reactor through a heated Brooks 5850S mass flow controller. The feed manifold consisted of two multi-port valves; the four-way valve allowed either a mixture of butane/N2 to go to the reactor or only N2 ; the eight-port valve was downstream of the four-way valve and was connected to the air manifold. The eight-way valve had two effluent streams; one entered the reactor and the other vented to atmosphere. Both lines were maintained at the same pressure to avoid pressure spikes or flow transients when the feeds were cycled. To prevent product condensation at the reactor effluent, the line was maintained at 200 ◦ C with an electrical heat tape. Liquid from the quench was pumped to this line immediately after the heat tape to minimize solids build up in the line but also to improve the scrubbing efficiency of the quench. The quench vessel was made of quartz and the accumulated acids were sampled frequently and analyzed off-line by HPLC (Hewlett Packard 1050) equipped with a variable wavelength UV detector. In addition to the frequent sample analysis, the quench liquid conductivity (referred as λ) was measured continuously to monitor acid production rate. The
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acid concentration and conductivity are related by a second order polynomial function (MA = 37.308λ2 + 349.78λ, being maleic concentration in ppm and conductivity in mS/cm) determined experimentally. At the end of each experiment, the absorber was washed three times. These samples were also analyzed by HPLC and included in the mass balance calculations. A slip stream of the effluent gas passed through an ice trap and the non-condensables were measured with various on-line analytical instruments. An HP 5890 GC equipped with both a flame ionization detector and a thermal conductivity detector measured the concentrations of CO, CO2 , O2 , N2 and C4 H10 . Oxygen was also monitored on-line (in real time) with a Siemens Oxymat 5F paramagnetic analyzer that had a response time of 1 s for 80% of full scale. ABB model 501B IR analyzers monitored the both CO and C4 H10 gases (on-line in real time). 2.2. Catalyst VPO catalyst used in this study was prepared in an “organic medium” and encapsulated in a porous silica shell in order to make it resistant to forces and abrasion typical of fluid bed reactors. The catalyst was equilibrated during 400 h on stream before the experiments were run. Catalyst activity and inventory were constant throughout the several months required to complete the study.
Fig. 2. Feed gas configuration and analytical equipment.
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Particle characteristics have been described by Lorences [15]. The mean particle diameter was 64 m with 20% less than 44 m. The particle density was 1890 kg/m3 and the minimum fluidization velocity was 5.6 mm/s. 2.3. Experimental procedure The gas manifold design allows one to simulate many different configurations and examine transient behavior under industrially relevant reactor conditions. In a previous communication [16], we described steady-state experiments and explored the effect of butane and oxygen concentration on the production of by-product acids, of which acetic and acrylic predominate. Herein we concentrate on cycling feeds between net reducing and oxidizing environments; a mixture of butane/nitrogen/air is fed to the reactor for a specified period of time; the reactor is purged for 10 min with nitrogen and then an oxygen/nitrogen mixture is fed. The experimental conditions are shown in Table 1 and the procedure followed was: 1. Catalyst reduction during 40 min varying the butane/oxygen concentration in the feed from 2 to 9 and 0 to 10%, respectively, at temperatures of 380 and 410 ◦ C. 2. Nitrogen purge for 10 min. 3. Catalyst oxidation for 40 min with oxygen varying between 2 and 20% and pressures as high as 5 abs bar. 4. Nitrogen purge for 10 min. 5. Second oxygen cycle for 30 min at the same conditions as Step 3. This served as a reference to calculate oxygen uptake.
3. Results and discussion 3.1. Catalyst reduction by butane Several authors have demonstrated that VPO catalyst reacts butane under anaerobic conditions for periods of time on the order of minutes [17]. DuPont’s transport bed process [3] relies, in part, on lattice oxygen from the catalyst to react butane. In this process, a reducing gaseous environment rich in butane is fed to the bottom of a transport bed reactor. The gas carries the solids to the top of the reactor into a cyclone. The effluent gases exit through the top of the cyclone and the solids fall into a stripper. The catalyst is then carried down through a standpipe into a regenerator with a high partial pressure of oxygen. The reoxidized catalyst is returned to the bottom of the transport bed after the regenerator. In this study, the first set of experiments (#1–#3 in Table 1) were conducted analogously to DuPont process, in that we cycled the catalyst between oxidizing and reducing environments. Before the initial reduction step (Step 1), catalyst was oxidized for an extended period of time; the reduction was conducted with different butane/N2 mixtures and lasted 40 min. The reactor was purged by a nitrogen stream for 10 min and then air was fed to it for 40 min. (Oxidation and reduction steps are much shorter in the commercial process!) The nitrogen purge and air treatment was then repeated. The time delay between switching the valves and the gas phase analytical measurements was on the order of 1 min. It is lower for the conductivity measurement because the gas hold-up between the reactor and quench is much smaller than between the reactor and gas phase detectors.
Table 1 List of operating conditions under cycling mode Experiment
T (◦ C)
Step 1 (balance N2 )
Step 2
Step 3
Step 4
Step 5
#1 #2 #3 #4 #5 #6 #7 #8 #9 #10 #11 #12 #13 #14 #15 #16 #17 #18 #19 #20 #21 #22 #23
380 380 380 380 380 380 380 380 380 380 380 380 410 410 410 410 380 380 380 380 380 380 380
2% 5% 9% 2% 5% 9% 2% 5% 9% 2% 5% 9% 5% 5% 5% 5% 5% 5% 5% 5% 5% 5% 5%
N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2
Air Air Air Air Air Air Air Air Air Air Air Air Air Air Air Air 2% O2 /N2 5% O2 /N2 10% O2 /N2 5% O2 /N2 (3 bar) Air (3 bar) 5% O2 /N2 (5 bar) Air (5 bar)
N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2 N2
Air Air Air Air Air Air Air Air Air Air Air Air Air Air Air Air 2% O2 /N2 5% O2 /N2 10% O2 /N2 5% O2 /N2 (3 bar) Air (3 bar) 5% O2 /N2 (5 bar) Air (5 bar)
C4 C4 C4 C4 /2% O2 C4 /2% O2 C4 /2% O2 C4 /4% O2 C4 /4% O2 C4 /4% O2 C4 /10% O2 C4 /10% O2 C4 /10% O2 C4 C4 /2% O2 C4 /4% O2 C4 /10% O2 C4 C4 C4 C4 C4 C4 C4
(3 bar) (3 bar) (5 bar) (5 bar)
(3 bar) (3 bar) (5 bar) (5 bar)
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Fig. 3. Butane oxidation by lattice oxygen. Evolution of the reaction products during the different reduction and oxidation steps at 380 ◦ C.
Three different concentrations of butane (nominally 2, 5 and 9%) using nitrogen as a diluent were tested. The temperature was set at 380 ◦ C. Fig. 3 shows the butane, carbon monoxide, maleic acid (measured as conductivity, λ) and oxygen concentrations as a function of time. During the initial reduction step (Step 1) the conductivity of the liquid from the quench, which relates to maleic anhydride production—increases sharply for the first several minutes then levels off. CO formation reaches a maximum of 1.2% at 4 min before tailing off to zero. The CO2 was analyzed by GC and its concentration was about the same as CO in the first trace at 2 min but was undetectable after the second trace. The evolution of these products was entirely due to the catalyst lattice oxygen since oxygen was not co-fed. The maximum CO concentration peaked at 1.5% with 5% butane and 2.5% with 9% butane. As shown in Fig. 3, CO is detected for up to 14 min while in the experiment with 9% butane, the peak dropped to zero after only 8 min. Peak height depended on butane concentration but the total CO detected for each of the experiments was nominally about the same, which implies that the available lattice oxygen is independent of the butane concentration, but the reaction rate increases with butane partial pressure. We were surprised to detect a significant quantity of CO during the first oxidation step (Step 3). Moreover, we found the concentration increased with the butane feed concentration and reached a maximum of almost 1%. The shape of the curve is similar to the first CO curve during the catalyst reduction step. However, CO evolution during the oxidation step increases with butane concentration during the reduction step; peaks were higher and the tails of the curves were longer. Emig et al. [18] also detected small amounts of CO and CO2 during an oxidation pulses whereas Wang and Barteau [17,19] claimed little carbon oxide evolves during an oxidation step after reducing the catalyst. Traces of butane are detected for as much as 20 min after the beginning of the nitrogen purge (Step 2). Based on blank experiments with oxygen, we expected a complete purge after 10 min and this expectation is supported by the second
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Fig. 4. Oxygen breakthrough during the first oxidation step depending on the butane concentration present in the reduction step. (T = 380 ◦ C; oxidation agent: air).
oxidation step: 10 min after the second nitrogen purge, the oxygen concentration reaches zero. Despite the low levels of butane detected during VPO oxidation, we concluded that the CO peak is due to oxidation of carbonaceous species that built up on the catalyst surface. There is a marginal rise in the quench conductivity shown above but the rise was barely detectable for the experiments run with 5% butane and 9% butane. Fig. 4 compares the oxygen breakthrough of each of the three experiments during the first oxidation step (Step 3). The curve shifts to the right with increasing butane feed concentration, indicating increasing oxygen conversion. Coincidentally, CO evolution is highest for the case of the highest butane feed concentration. Maleic anhydride production is negligible during Step 3 with 2% butane in the feed and zero for Experiments 2 and 3 with 5 and 9% butane, respectively. The data support the hypothesis that carbon deposition/adsorption depends on the partial pressure of butane in the feed. Poor stripping efficiency during the nitrogen purge could also explain the trend in the data. However, if the CO evolution and oxygen conversion were due to residual butane in the gas phase, we would also have expected to see maleic acid in the quench, which was not the case. In a second series of experiments, we examined the nature of the lattice oxygen by reacting butane with VPO that was previously oxidized for different periods of time—1 and 48 h. Immediately after the air treatment, the reactor was purged with nitrogen and then a stream containing 2% butane/98% nitrogen was fed to the reactor for 40 min. The reactor was purged again with nitrogen for 10 min followed by air for 40 min. The conductivity of the quench was monitored on-line and the data are shown in Fig. 5. After the 48 h exposure, maleic anhydride production (conductivity) is twice as high as it is after the 1 h exposure. Additionally, consistent with Fig. 3, there is a slight rise in conductivity during the air treatment after the 10 min nitrogen purge. A third test was conducted in which we oxidized the catalyst in air for 48 h then reduced the catalyst with 9% butane
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Fig. 5. Lattice oxygen availability after different oxidation pre-treatments.
instead of 2%. The maleic anhydride evolution with respect to time, as shown in Fig. 5, follows the same trend as when we fed only 2% butane. This would imply oxygen reaction is rate limiting for both cases and that butane is in excess with only a feed concentration of 2%. Several studies concur that the reaction kinetics are highly dependent upon the reoxidation cycle time. Mills et al. [9] suggested that oxygen participates in the reaction as a subsurface lattice via solid-state diffusion coupled with surface oxygen. This hypothesis is consistent with the data showing an increase in production by increasing the oxygen exposure. On the other hand, Schuurman and Gleaves [8] proposed that the activation energy depends on the surface oxidation state and the activation energy was lower for more highly oxidized VPO. The data reported in Fig. 5 would not seem to support this concept in that the rate of formation of maleic anhydride is about the same for both experiments; lattice oxygen contribution (LOC) is higher and not the reaction rate. Emig et al. [18] also observed that strongly oxidized catalyst showed high activity and selectivity. Their data showed that the amount of oxygen available for the reaction increased from 60 to 190 mmol O2 /kgcat when the regeneration time changed from 30 s to 18 h, while the amount of maleic anhydride formed increased from 0.3–0.7 to 2–4 gmac /kgcat . The activation energy was found to be lower for the more highly oxidized samples, in agreement with Schuurman and Gleaves [8] observation. Practical commercial application of the subsurface lattice or lowered activation energies would appear to be prohibitive considering that extensive residence times are required. Residence times much greater than 2 min would have significant negative economic impact. 3.2. Catalyst reduction by butane with oxygen co-feed Both Emig et al. [18] and Schuurman and Gleaves [8] reported that, together with the lower activation energies, formation of MA and desorption is faster on a more oxi-
Fig. 6. Conductivity and oxygen concentration vs. time after exposition to different environments. Operating conditions: 5% C4 , 10% O2 at 410 ◦ C.
dized catalyst surface. However, this effect was observed under anaerobic conditions. To examine the effect under butane/oxygen co-feed conditions, we first allowed the catalyst to reach steady state with 5% butane and 10% oxygen, treated the catalyst for different amounts of time with nitrogen or air and then resumed feeding 5% butane and 10% oxygen. We tested four different “stripping/regeneration” environments: (a) 1 h of air, (b) 1 h of N2 , (c) 10 min of air, and (d) 10 min of N2 . Fig. 6 shows the evolution of conductivity and oxygen for each of the four stripping/regeneration treatments. After 10 min or 1 h of nitrogen, the oxygen breakthrough (during the reduction step) increases rapidly during the first couple of minutes then levels off to over 2%. After the air treatment, the oxygen concentration overshoots the steady-state value by a considerable amount. The overshoot is greater after the 1 h treatment and reaches 4%, which is double the steady-state value. The maleic anhydride production rate is essentially the same for each of the four “stripping/regeneration” treatments: oxygen breakthrough depends on the pre-treatment conditions while maleic anhydride evolution is independent of the pre-treatment. This phenomenon would imply that VPO reduction by butane is the rate limiting step under these conditions. As long as the partial pressure of oxygen is sufficiently high, the butane concentration limits the maleic anhydride reaction rate.
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Fig. 7. Evolution of the different products: (a) C4 H10 , (b) CO, (c) maleic acid (conductivity), (d) O2 depending on the butane concentration during the reduction step. Operating conditions: 10% O2 at 380 ◦ C.
Under anaerobic reducing conditions, lattice oxygen is the limiting reagent, thus butane concentration has a limited impact on maleic yield. However, when oxygen is co-fed together with butane, reaction rates, maleic production and lattice oxygen concentrations may all depend on both the reoxidation and reduction step. Experiments #4–#12 explore the effect of butane and oxygen partial pressure during the reduction step on the catalyst reoxidation. Fig. 7 illustrates the product distribution as a function of the inlet butane concentration—2, 5 and 9%—with 10% oxygen (Experiments #10–#12). In this case, both the CO and the maleic acid concentration (conductivity) increase with increasing butane in the feed. (Note that the peak in the butane and oxygen concentration after 40 min corresponds to the feed gas analysis.) Whereas CO breakthrough during the oxidation step was evident at all anaerobic conditions, CO is only detected for the case with 9% butane and 10% oxygen. As a consequence, oxygen uptake during catalyst oxidation is about the same for all three cases. 3.3. Catalyst oxidation Several tests were conducted to evaluate the VPO oxidation rate. In the last series, the catalyst was reduced by a stream of 5% butane in nitrogen at 380 ◦ C then treated
with different oxygen concentrations—2, 5, 10 and 21% (#2, #17–#19 in Table 1)—and pressures of 1, 3 and 5 bar (#2, #18, #20–#23 in Table 1). Oxygen breakthrough was fastest at the higher pressures. The rate of CO evolution was highly dependent on the oxygen partial pressure; the maximum value was higher at higher oxygen partial pressures; at the lower oxygen partial pressures, the tail was longer and the peak time appeared several minutes later. Total CO evolution was calculated by integrating the area underneath the curve. Fig. 8(a) shows the total oxygen consumption—lattice oxygen contribution—is estimated by subtracting the area under the oxygen breakthrough curve in Step 5 from Step 3. The second part of the figure shows an estimate of the quantity of oxygen required to combust the adsorbed carbonaceous species. The calculation was based on the CO evolution during step and the ratio CO/CO2 given by the GC analysis (only CO is shown in the figure). We assumed that butane was adsorbed on the surface, although butene or butadiene is a more likely species. Fig. 8 is consistent with expectations and shows that carbon adsorption increases with the severity of the reducing conditions. More butane is deposited on the surface in an anaerobic atmosphere. Under highly oxidizing conditions the surface is free of carbon species. However, even when the feed contains 10% oxygen, carbon deposits on the surface at high butane partial pressures.
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the both butane conversion and MA selectivity. Although the model adequately characterized the data set, it was fundamentally flawed because it had no provision to allow for carbon–metal complexes that form under reducing conditions. The goodness-of-fit between an experimental data set and a theoretical model is a necessary consideration with respect to model discrimination; however, it is often not sufficient. The final step in the modeling exercise was to include the carbon–metal complex but again to attempt to minimize the degrees of freedom by limiting the number of possible reactions. The model assumed three types of sites: an active site, V5+ , that either reacted with butane to form maleic anhydride or it combusted the maleic anhydride, as shown by Eqs. (1) and (2); a reduced site, V4+ , resulting from the selective oxidation of butane or combustion of maleic anhydride; and finally, a carbon–metal complex, VC4 , formed when butane chemisorbs on to a V4+ site, Eq. (3). When the gas phase oxygen concentration is low, the number of V4+ sites increases and the surface coverage of the carbon–metal complex also increases. Eq. (3) assumes that this complex only reacts with gas phase oxygen. Therefore, a highly reduced catalyst will result in a large fraction of the so-called VC4 phase. Eventually, equilibrium is reached between the gas composition and the three types of sites. This model predicts that maleic selectivity depends to a large extent on maintaining the catalyst surface oxidized: k1
Fig. 8. Total oxygen uptake (a) and oxygen consumed to form CO (b) as a function of the butane and oxygen concentration in the reduction step: (䊊) 0% O2 , (䊐) 2% O2 , () 4% O2 , (䉫) 10% O2 .
C4 H10 + αV5+ − → MA + αV4+ + 4H2 O, r1 = k1 [C4 H10 ][V5+ ]
(1)
k2
As shown in Fig. 8, the lattice oxygen contribution follows the identical trend as for the carbon deposition: highly reducing anaerobic conditions extract more oxygen from the catalyst compared to oxidizing conditions. Under anaerobic conditions with 9% butane, the LOC number exceeds 140 mm O2 /kg VPO and it is almost an order of magnitude lower with a 10% O2 feed concentration.
→ 4COx + βV4+ + H2 O, MA + βV5+ − r2 = k2 [MA][V5+ ] k3
(2)
→ 4VC4 , C4 H10 + 4V4+ − k4
O2 + V4+ − → V5+ ,
r3 = k3 [C4 H10 ][V4+ ]
r4 = k4 [O2 ][V4+ ]
(3) (4)
k5
γO2 + 4VC4 − → 4COx + 4V4+ + 5H2 O, 4. Kinetic modeling
r5 = k5 [O2 ][VC4 ]
In a previous publication, Lorences et al. [16] generated a series of over 120 steady-state experiments with a broad range of feed concentrations, temperatures and contact times. Several published models were compared to the experimental data and agreement between butane conversion and experimental observations were very good. However, the models poorly characterized maleic anhydride selectivity and a new kinetic expression was developed. The premise of the model was that the number of independent reactions should be minimized but that there were two active sites—one selective to maleic anhydride formation and the second non-selective. They successfully fit the model to
The reaction rate equations are re-written in terms of the normalized oxidation state by dividing by the total number of active sites —VT — and assuming that the total was constant (Eqs. (6)–(10)). The sum of the fraction of the sites in each “oxidation” state—V5+ , V4+ and VC4 —equaled 1. Some of the experimental evidence might suggest that the total number of sites varies with the oxygen treatment. For example, Fig. 5 shows that after 48 h of exposure to air, the maleic anhydride yield is higher compared to a 1 h treatment. On the other hand in the case of much shorter time scales and aerobic conditions, oxygen appears to react only with sites reduced by butane, as shown in Fig. 6: catalyst pre-treated
(5)
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with oxygen did not increase the maleic anhydride formation rate and the higher initial oxidation state resulted an overshoot in the oxygen concentration that corresponds to higher initial oxidation state: r1 =
k1 [C4 H10 ][VT − V4+ − VC4 ] VT
(6)
r2 =
k2 [MA][VT − V4+ − VC4 ] VT
(7)
r3 =
k3 [C4 H10 ][V4+ ] VT
(8)
r4 =
k4 [O2 ][V4+ ] VT
(9)
r5 =
k5 [O2 ][VC4 ] VT
(10)
4.1. Hydrodynamic modeling Most experimental treatments of butane oxidation are conducted in the fixed bed fluid regime: reactor modeling is simplified because the gas phase is in plug flow. On the other hand, these reactors can often have temperature and pressure gradients as well as radial or inter-particle gradients. Further complications arise from gas to particle diffusion rates that may cause localized hot spots. Gas phase modeling of fluid beds is generally more difficult because bubbles form at the grid and rise through the bed at velocities significantly higher than the superficial gas velocity. Consequently, both contact efficiency and contact time are reduced. However, the solids phase is backmixed to a high degree and the reactor is generally isothermal, which simplifies the modeling. In this work, we operated at relatively low gas velocities. At velocities in the vicinity of Umf , the gas is in plug flow and the reactor operates in the regime. Deviation from plug flow increases with increasing gas velocity. To evaluate the hydrodynamics, we conducted a series of inert gas step-up response experiments and then characterized the data with a CSTR in series model [16]. Agreement between the experimental RTD data and a model assuming five CSTRs in series was excellent. 4.2. Parameter estimation Fig. 9 illustrates the experimental data for the case where we fed 2% butane and 4% oxygen to catalyst that had been oxidized for a period of at least 50 min. The COx concentration is based on the IR response of the CO analyzer plus the CO/CO2 ratio measured by the GC. The liquid quench conductivity began to rise after 1 and 2 min the slope of the line was constant. This could imply that during the first minute of reaction, maleic anhydride was combusting on the oxidized catalyst surface but was constant thereafter. The butane concentration was zero for the first minute then began to increase modestly thereafter. It reached its steady-state value
Fig. 9. Product gas evolution in the initial reduction step (experimental and model predictions): Feed composition—2% butane, 4% oxygen, T = 380 ◦ C.
of 1.5% after about 10 min. Both oxygen and CO signals are recorded after 1 min, while the butane signal was still zero. Their concentrations significantly overshoot the steady-state value and reach a maximum of 2.7% for O2 and 4.5% for CO + CO2 . As with the butane, they reach the steady-state values at about 10 min. We combined the hydrodynamic model of the fluid bed reactor together with the kinetic model outlined in Eqs. (6)–(10) to simulate the catalytic performance of the VPO. In the first simulation, we adopted the kinetic rate constants derived from the steady-state experiments of Lorences et al. [16]. Agreement between the model predictions and experimental data was poor. The model did not predict the overshoot of the CO (and CO2 ) and O2 concentrations shown in Fig. 9. Rather, it predicted that the concentrations increased asymptotically from zero to the steady-state values. Clearly the pre-oxidation step has a substantial effect on catalytic performance. Instead of adopting the rate constants from the steady-state experiments, we optimized the rate constants that fit the experimental data. The most difficult part of this process was to adequately predict the overshoot, while at the same time match the steady-state value beyond 10 min. Fig. 9 compares the model predictions against the experimental data and the agreement over the entire range is excellent. However, beyond 14 min the model predictions of the oxygen appear to begin to deviate from the experimental data somewhat.
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Table 2 Kinetic rate constants and activation energies i
k0i (s−1 ) (transient)
k0i (s−1 ) (steady state)
Eai (kcal/mol)
1 2 3 4 5
0.91 2.5 0.13 0.28 0.34
0.33 0.37 0.05 0.57 0.06
15 14 17 9 28
Table 2 compares the best fit kinetic rate constants of the transient experiment with the values derived from the steady-state experiments. (Note that the same activation energies were used for both.) In the steady-state experiments, the rate constants for butane to maleic anhydride (k01 ) and maleic anhydride to COx (k02 ) differ by about 20%. The non-selective reaction rate is greater than 2.5× the selective reaction rate for the transient experiment. In general, the pre-exponential factors from the transient experiment are from three to five times the value of the best fit parameters from the steady-state experiments except for catalyst reoxidation from V4+ to V5+ (k04 ). In this case, the steady-state value is double the transient value. Presumably, one contribution to the large differences in the best fit parameters of the two data sets is due to the initial oxidation state of the catalyst. The data show that the available lattice (or bulk) oxygen for reaction may depend on the pre-treatment step. However, extended catalytic oxidation would appear to increase the rate of maleic oxidation but not the selective oxidation of butane to maleic anhydride. Further modeling efforts are required to take into account the relationship between oxygen pre-treatments (regeneration) and the available lattice oxygen for reaction with butane or maleic anhydride. This exercise clearly shows the value of employing alternative reactor types when modeling reaction kinetics. The steady-state experiments provide a data base with which to gauge the effect of temperature, composition and contact time on reaction rates and the transient experiments can help substantiate the reaction mechanism.
5. Conclusions We introduce a novel reactor configuration based on operating a fluid bed at low gas velocities such that gas bypassing is minimized but where solids backmixing is sufficiently high to ensure a uniform temperature distribution. The reactor is coupled with a means of quantitatively monitoring the primary and by-product acids on-line and in real time by measuring the conductivity of condensable gases in a quench and withdrawing frequent samples for subsequent HPLC analysis. The on-line gas analysis demonstrated that carbon is strongly adsorbed to the surface under net reducing conditions. Butane reacts with lattice oxygen and the kinetic model suggests that the same vanadium species also combusts
maleic anhydride. The series reaction to combustion products is the prime determinant of selectivity. Under fuel lean, oxygen rich conditions, reaction rates primarily depend on butane concentration. However, co-feed butane/oxygen conditions where the partial pressure of oxygen at the exit of the reactor approaches zero (in the vicinity of about 1%), carbon builds up on reduced sites—V4+ —to form an adsorbed species. This species reacts with molecular oxygen to form CO/CO2 . Although the same kinetic model can fit both transient and steady-state data, the differences in the optimized rate parameters are substantial. Reaction rates of a previously oxidized catalyst are significantly higher than a catalyst operating under steady-state conditions. This enhanced activity is most likely due to a higher density of oxygen on the surface or subsurface layers. Under extended oxygen exposure times, maleic anhydride production rates can be increased by a factor of 2. However, more detailed modeling is required to identify the nature of the non-selective reactions of a highly oxidized VPO catalyst.
Acknowledgements The authors gratefully acknowledge the financial and technical support of DuPont Ibérica S.A. and the assistance of Mónica Alonso Carreño during the experimental work.
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