6730
Ind. Eng. Chem. Res. 2003, 42, 6730-6742
Butane Oxidation to Maleic Anhydride: Kinetic Modeling and Byproducts Marı´a J. Lorences,†,‡ Gregory S. Patience,*,§,| Fernando V. Dı´ez,† and Jose´ Coca† Department of Chemical Engineering and Environmental Technology, University of Oviedo, Julian Claverı´a s/n, Oviedo, Asturias, Spain 33209, and DuPont Ibe´ rica SA, Tamo´ n, Avile´ s, Spain 33469
Reactor technology for maleic anhydride continues its evolution. New processes achieve higher yields with lower investment by operating in a net reducing environment where the oxygen concentration is lower than that required stoichiometrically to react all of the butane. In this paper, we examined a wide range of operating conditions to quantify the effect of a reducing environment on maleic anhydride selectivity, byproduct acid productivity, and reaction rates. The experiments were carried out with a vanadium phosphorus oxide catalyst in a fluidizedbed reactor and a novel feed gas manifold. Oxygen, carbon monoxide, butane, and acid concentrations were measured online at a frequency of about 1 Hz. Acetic and acrylic acids were the predominant byproduct acids, but fumaric, methacrylic, and phthalic acids were also detected. Under reducing conditions, carbon adsorbed on the catalyst surface, byproduct acid yields increased, and both the selectivity and reaction rates decreased. A redox kinetic model was developed to account for the experimental observations and included both V5+ and V4+ oxidation states and a “VC4” complex, which represented carbon adsorption. 1. Introduction Maleic anhydride is synthesized commercially by partially oxidizing n-butane over vanadium phosphorus oxide (VPO) catalysts. Over the last 10 years, its price on the merchant market has dropped significantly, and the decline has been due to catalyst improvements, process innovations, and economies of scale. Early technology was entirely based on fixed-bed reactors with benzene not n-butane as a feedstock. The fluidized-bed process was commercialized in the late 1980s and has several advantages including superior heat transfer, more concentrated product streams, and larger scale. In the mid-1990s, circulating fluidized-bed (CFB) technology was commercialized in which catalyst is shuttled between an oxidizing environment and a reducing (butane-rich) environment. This process has good heattransfer characteristics but has the advantage of even larger scale than conventional fluidized beds as well as an even more concentrated product stream.1 Concentrated product streams and high butane feed concentrations translate into reduced vessel sizes (catalyst inventories) and thus superior economics. Monsanto2 have operated a fixed-bed reactor in the flammability region (C4H10 > 1.8 vol % in air) and claimed that the hot spot can be overcome using 40% dilution in the first part of the reactor. Recently, Pantochim3 claims to have improved fixed-bed process economics by feeding pure O2 instead of air and recycling the non* To whom correspondence should be addressed. E-mail:
[email protected]. † University of Oviedo. Fax: +34-98-510-3434. E-mail: fds@ correo.uniovi.es. ‡ Current address: HALDOR TOPSØE A/S, Nymøllevej 55, 2800 Lyngby, Denmark. Tel.: +45-45272330. E-mail: mlp@ topsoe.dk. § DuPont Ibe´rica SA. | Current address: INVISTA (International) SA, European Technical Centre 146, route du Nant-d’Avril, CH-1217 Meyrin, Geneva, Switzerland. Tel.: +41-22-7176996.
condensable gases. The feed stream surpasses the 1.8 vol % butane limit and operates at concentrations near 4% with oxygen concentrations in the range of 10%. Although several processes under development operate with high inlet butane concentrations, most kinetic studies published in the literature are restricted to highly oxidizing reaction conditions typical of conventional fixed beds (see work by Hutchings4). However, under reducing conditions, neither is oxygen a limiting reagent nor is catalyst reoxidation a rate-limiting step. Therefore, the kinetic expressions have limited usefulness for either fluidized-bed processes, CFB technology, membrane reactors, or fixed-bed processes where oxygen is the limiting reactant. Many recent studies5-7 have been devoted to characterizing the catalyst performance under fuel-rich conditions and under a cyclic mode where the catalyst is exposed to successive pulses of a net reducing and oxidizing environment.8-11 The former work was devoted to evaluating the potential of membrane reactors, whereas the latter was pertinent to CFBs. The reducing conditions expected at the inlet of a membrane reactor lead to a rapid reduction of the catalyst surface and, therefore, formation of butenes and COx may be favored. Mallada et al.12 report that, in fixed-bed reactors operating at high butane concentrations, the catalyst near the effluent becomes less selective as the gas-phase oxygen is depleted: the COx selectivity increases, and the overall activity declines. In this paper, we present new experimental data under both butane-rich and -lean conditions. All experiments were run in a fluidized-bed reactor with a capacity of up to 0.5 kg. The gas manifold was specifically designed to operate either under steady-state feeding conditions or under unsteady-state cycling conditions to monitor transients and simulate typical CFB operation. Together with the wide range of butane and oxygen concentrations, we evaluated the catalytic performance at various temperatures, gas flow rates, and water vapor feed rates.
10.1021/ie0302948 CCC: $25.00 © 2003 American Chemical Society Published on Web 11/20/2003
Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 6731
Figure 1. Fluidized-bed reactor in detail.
The objectives of the paper are to quantify the effect of a reducing inlet gas composition on the byproduct acid profile and evaluate the kinetic reaction rates. We compare several kinetic models with our experimental data. Many models predict butane conversion well but not selectivity. Most models are entirely empirical and therefore have limited use in predicting the reactor performance outside the scope of experimental conditions. To develop a more rigorous model, we consider both steady-state and transient experiments. Under cyclic operation, carbon may adsorb strongly to the surface and only desorb when oxygen is fed to the catalyst. The new kinetic expression characterizes this phenomenon and at the same time correctly characterizes the effect of butane and oxygen concentrations under steady-state conditions. Further work is required to incorporate the byproduct acids as well as the product split between CO and CO2. 2. Experimental Section All catalytic testing was carried out in a Hastelloy C-276 vessel of 0.04 m diameter and 0.79 m height. Its design pressure and temperature were 51 bar and -30/ 620 °C, respectively. The reactor was immersed in an
electrically heated sand bath (30 A) to maintain nearisothermal conditions. The axial temperature gradient was monitored with a 10-point thermocouple positioned equidistant along the height. The temperature was constant along the bed height but dropped substantially in the freeboard section, indicating that gas-phase product decomposition was minimal. Air, nitrogen, and butane feed streams are premixed outside of the bed. As illustrated in Figures 1 and 2, the feed gas is fed through a tube from the top of the heated sand bath. The tube is coiled around the circumference of the reactor in the sand bath to preheat the gas. It entered the bottom of the reactor into a plenum and then through a sintered metal frit that distributed it evenly across the diameter. We tested various superficial gas velocities, but most experiments were conducted in the bubbling fluidized-bed regime with a solids bed height between 100 and 500 mm. A flanged upper section minimized solids entrainment to the top of the reactor, and a sintered metal filter retained all catalyst in the reactor. Feed/Effluent Section. Figure 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 flow controllers. Accurate measurement of the butane flow was problematic, and we eventually succeeded in feeding butane at a constant rate with a Brooks 5850S mass flow controller heating the line and the flowmeter to 60 °C. An eight- and four-way valve system was developed either to cycle butane/N2 or oxygen/N2 (simulating CFB operation) or to feed them together. The reactor effluent was maintained at 200 °C to prevent product condensation in the line, and it was scrubbed with water recycled from the reactor quench. The accumulated acids in the quench/absorber were sampled frequently and analyzed offline by highperformance liquid chromatography (HPLC; HewlettPackard 1050) equipped with a variable-wavelength UV detector. An ION-300 column was used to separate the different acids, which included acetic, acrylic, succinic, fumaric, propionic, butyric, methacrylic, and phthalic
Figure 2. Experimental setup: (1) compressor; (2) electroneumatic valve; (3) mass flow controller; (4) check valve; (5) heater; (6) 10point thermocouple; (7) fluidized sand bath; (8) peristaltic pump; (9) absorber; (10) oxygen paramagnetic analyzer; (11) gas chromatograph; (12) integrator; (13) butane and CO IR analyzers; (14) data acquisition system; (15) computer.
6732 Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003
Figure 3. Evolution of the oxygen and acids concentration under experimental conditions.
acids. In addition to the frequent sample analysis, the quench liquid conductivity was measured continuously online to monitor the acid production rate. 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. Downstream of the absorber, a slip stream of the effluent gas passed through an ice trap and the noncondensables were measured with various online analytical instruments. The rest of the exit gas was bubbled in water, and both liquid samples were analyzed by HPLC. A relatively small amount of carbon was collected by the ice trap and in the purge gas. The efficiency of the absorber was greater than 99%. A Hewlett-Packard 5890 gas chromatograph equipped with both a flame ionization detector and a thermal conductivity detector measured the concentrations of CO, CO2, O2, N2, and C4H10. Oxygen was also monitored online (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 both CO and C4H10 gases (online and in real time). Figure 3 shows an example of the oxygen concentration of the effluent gas and the quench liquid conductivity of a typical experiment. During the first 5 min, oxygen rises sharply from 0% to over 7%. Thereafter, it reaches a constant value over 7.5% after 10 min. The conductivity increases rapidly during the first several minutes and then increases with time linearly. (Note that the acid concentration and conductivity are related by a second-order polynomial function.) To reduce error related to the transient, all mass balance calculations were based on acid and GC analysis measured after 18 min. A typical experiment lasted from 70 to 180 min and depended ultimately on the reproducibility of the GC traces; a minimum of three were collected for each experiment. To ensure safe operation, the reactor was controlled with a PLC and instrumented with interlocks for high reactor pressure and temperature. Power failure or low flow rates would also trigger interlocks, and the fail safe mode was to switch air/butane feeds to nitrogen. In the case of low flow and high temperature or pressure, power to the electrical heaters would shut off. Catalyst. A 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
Figure 4. SEM image of the VPO fluidized-bed catalyst. Table 1. Particle Properties of the VPO Catalyst dp < 10 10 < dp < 20 20 < dp < 25 25 < dp < 32 32 < dp < 44 44 < dp < 66 66 < dp < 88 88 < dp < 100
0 0.5 2.0 5.2 12.7 22.9 17.0 7.1
100 < dp < 125 125 < dp < 150 150 < dp < 200 200 < dp < 250 dp,sv Fp Fsk Umf
11.1 7.6 8.9 3.9 64 1890 2990 0.0056
forces and abrasion typical of fluidized-bed reactors.13 The scanning electron microscopy (SEM) image (Figure 4) shows that the particles are spherical and the surface is smooth with very few asperities. After several hundred hours of exposure to reaction conditions, we began the kinetic experiments. We frequently returned to a base case condition and found that the catalyst activity was constant throughout the several months required to complete the study. Table 1 shows the particle size distribution of the VPO catalyst analyzed with a Coulter instrument and some other particle properties relevant to fluidized-bed reactors. The mean particle diameter (dp,sv) was calculated according to the following expression:
dsv )
1 xi dpi
( )
∑
(1)
The particle and skeletal densities were measured by helium pycnometry, and the minimum fluidization velocity was measured experimentally.14 3. Results and Discussion 3.1. Experimental Results. Three different sets of steady-state experiments were carried out to assess the influence of operating conditions on the reactor performance and to develop a suitable model for the butane to maleic anhydride reaction kinetics. The first set of experiments was conducted at atmospheric pressure and different flows, temperatures, and feed concentrations. The second set was carried out with half of the amount of catalyst of that used in the first series. Finally, a statistical design of experiments was performed with the butane concentration, temperature, and reactor gas velocity as factors. Table 2 details the target
Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 6733 Table 2. Experimental Conditions (Variables and Levels) W P F T C4H10 inlet, vol % O2 inlet, vol %
W1, 2W1 1 0.76, 1.8, 3.6 623, 653, 693 2, 5, 9 4, 10, 20
value of the operating conditions, and Table 3 summarizes a selection of the raw data. A few experiments were conducted at 2 and 4 barg, but the results obtained were not fully satisfactory because of the unreliability of the feeding butane at these pressures. It would condense in the line, and thus its flow rate control was poor. Another set of experiments was also run, introducing different amounts of water in the feed, but the results were not conclusive. Experimental data where the error of either carbon or oxygen mass was greater than 5% were rejected. Figures 5-7 demonstrate the effect of both temperature and feed gas concentrations on butane conversion and maleic anhydride and main byproduct acid selectivity. At all temperature and feed gas concentrations, the maleic anhydride selectivity decreases with increasing butane conversion. Buchanan and Sundaresan15 reported a maximum in maleic yield at around 90% conversion, whereas Centi et al.16 observed a maximum at about 80% conversion operating at 300 °C on a high surface area fresh VPO catalyst. In this study, maleic yields (defined as conversion multiplied by selectivity) are highest at about 80% with 1-2% butane in the feed.
Figure 5. Conversion of butane vs selectivity to maleic anhydride at different reaction temperatures and feed concentrations.
However, at these conditions, maleic acid production rates were 60% of those recorded with 4% butane and 9% oxygen! Maleic selectivity is greatest at low temperatures, indicating that its activation energy is lower than that for the reaction to combustion products. The maximum is in the range of 70% and is achieved at 350 °C regardless of the inlet gas composition. Figure 5 shows that there is a near-linear relationship between the maleic anhydride selectivity and conversion: For each inlet gas composition, the selectivity declines steadily with increasing conversion. The slope of the curve is almost vertical under the most reducing conditions and progressively becomes less steep as the environment
Table 3. Experimental Results T, K
Q inlet, cm3/min
C4H10 inlet, vol %
O2 inlet, vol %
X, %
SMA, %
Sacet, %
Sacry, %
Sfum, %
Smacr, %
Sphthal, %
SCO, %
SCO2, %
623 623 623 623 623 623 623 623 623 623 623 623 653 653 653 653 653 653 653 653 653 653 653 653 653 653 693 693 693 693 693 693 693 693 693 693 693 693 410
1800 1800 3580 3580 1790 1790 3590 3585 1790 1800 3590 3600 1800 1800 3580 3585 1790 1790 3590 3585 1790 1800 3590 3600 3615 1805 1805 1800 3580 3585 1790 1790 3590 3585 1790 1800 3595 1735 3450
1.30 1.30 1.61 1.61 4.67 4.68 4.81 4.82 9.05 9.00 9.02 9.00 1.50 1.50 1.61 1.61 4.67 4.68 4.81 4.82 9.05 9.00 9.02 9.00 2.00 1.84 1.50 1.50 1.61 1.61 4.67 4.68 4.81 4.82 9.05 9.00 9.02 2.49 1.80
4.06 10.03 3.96 9.94 3.95 9.95 3.94 9.93 3.95 9.89 3.94 9.89 4.05 10.01 3.96 9.94 3.95 9.95 3.94 9.93 3.95 9.89 3.94 9.89 19.81 19.84 4.05 10.01 3.96 9.94 3.95 9.95 3.94 9.93 3.95 9.89 3.94 19.71 19.85
35.61 45.34 22.35 27.84 12.75 20.46 8.14 11.65 7.33 13.09 4.62 8.04 43.08 56.71 29.66 42.39 16.25 31.11 12.30 23.84 8.73 18.95 6.34 13.73 35.53 60.30 50.44 68.48 35.39 47.04 19.08 41.52 14.65 29.53 8.90 23.82 8.98 64.06 48.86
66.32 64.42 64.68 67.17 64.15 66.16 68.59 66.10 54.60 64.92 60.79 64.61 56.65 52.09 60.21 51.88 48.74 55.59 54.72 54.46 41.69 47.69 49.85 52.42 62.87 49.43 51.87 44.60 54.96 52.98 35.01 49.85 45.03 53.59 28.91 42.75 37.28 47.00 51.10
4.74 5.94 8.87 5.85 5.04 7.21 0.92 4.74 10.71 6.12 4.77 6.03 5.09 8.56 9.60 6.81 6.13 8.12 8.76 5.86 9.52 5.34 5.74 6.91 2.77 3.18 3.39 7.01 4.42 4.09 3.92 4.46 3.74 4.50 4.23 5.35 5.91 3.27 3.62
0.86 0.86 1.15 1.11 1.45 1.31 1.53 1.31 1.52 0.03 1.46 1.54 0.84 0.77 1.06 1.17 1.59 1.41 1.64 1.37 1.92 2.14 1.84 1.56 1.23 0.76 0.68 1.62 0.82 0.85 1.54 1.08 0.02 1.26 2.32 1.66 2.00 0.86 1.06
0.08 0.06 0.04 0.04 0.05 0.06 0.05 0.05 0.06 0.08 0.04 0.05 0.10 0.06 0.06 0.04 0.05 0.06 0.06 0.06 0.11 0.07 0.07 0.07 0.06 0.09 0.04 0.05 0.03 0.36 0.48 0.08 0.05 0.06 0.13 0.06 0.13 0.08 0.05
0.03 0.04 0.08 0.07 0.11 0.09 0.15 0.12 0.19 0.12 0.25 0.17 0.06 0.08 0.09 0.12 0.08 0.08 0.10 0.20 0.16 0.12 0.18 0.23 0.07 0.04 0.00 0.03 0.04 0.04 0.11 0.14 0.12 0.06 0.19 0.09 0.15 0.05 0.08
0.00 0.00 0.00 0.00 0.53 0.19 0.54 0.22 1.33 0.43 1.01 0.41 0.10 0.16 0.15 0.27 0.82 0.13 0.74 0.12 0.01 0.00 0.01 0.01 0.04 0.06 0.00 0.00 0.00 0.00 1.19 0.36 1.01 0.18 1.76 0.00 0.11 0.03 0.02
12.69 13.06 13.88 12.88 12.76 11.30 14.17 13.43 13.34 13.64 14.99 13.23 18.94 20.15 14.19 23.07 21.48 17.91 16.20 21.31 20.49 23.82 21.27 22.00 17.30 25.37 22.01 23.83 21.77 21.55 29.34 21.75 26.03 21.24 27.18 25.11 27.18 24.53 22.93
15.27 15.63 11.30 12.87 15.90 13.68 14.06 14.03 18.24 14.66 16.69 13.96 18.22 18.14 14.64 16.64 21.10 16.69 17.78 16.61 26.11 20.81 21.04 16.80 15.66 21.08 22.01 22.86 17.96 20.12 28.41 22.27 23.99 19.09 35.28 24.98 27.24 24.18 21.14
6734 Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 Table 4. Product Distribution at 633 K in the Exit Gas Line t ) 20 min C4H10 CO2 CO maleic acid acetic acid fumaric acid acrylic acid methacrylic acid phthalic acid
Figure 6. Conversion of butane vs selectivity to acetic acid at different reaction temperatures and feed concentrations.
Figure 7. Conversion of butane vs selectivity to acrylic acid at different reaction temperatures and feed concentrations.
becomes more oxidizing. On average, a 30 °C temperature rise drops the selectivity by 10%; at low butane concentrations (2%), it drops by 8% for a 30 °C rise, and it drops by 12% at the highest butane concentration tested (Figure 5). Centi et al.16 reported a decrease in the selectivity related to the ratio of residual concentration of oxygen to n-butane (operating at 0.32, 0.7, and 1.1% butane in air). This ratio is proportional to the oxidizing power of the mixture. It is suggested that at the end of the catalytic bed the maleic anhydride produced is not stable as a result of the overoxidation of the catalyst surface. Under these low feed butane concentrations, an increase of the V5+ concentration at the end of the catalyst bed was confirmed by chemical and spectroscopic analyses of the spent catalyst.17 The relationship between the reaction conditions and conversion is more complex. Under highly oxidizing conditions (2% butane and air), butane is the limiting reagent and, therefore, conversion is limited by the contact time and temperature. Under reducing conditions (9% butane/4% oxygen), oxygen is the limiting reagent. The temperature or contact time has little impact on the absolute value of the butane conversion but affects the selectivity considerably. For example, by increasing the temperature by 60 °C, the butane conversion increased from 7% to 9% but the selectivity dropped by over 25%. Buchanan and Sundaresan15 showed that acetic and acrylic acids are the principle byproduct acids, and they also detected trace amounts of ethylene. Together with these acids, we also detected low concentrations of phthalic, methacrylic, and fumaric acids. The average selectivities were 0.3%, 0.1%, and 0.08%, respectively, and the data are summarized in Table 3. Selectivity of phthalic acid reached as high as 1.7% at high temperatures and a reducing atmosphere. Under oxidizing
3.14% 2.83% 1.90% 1075 ppm 86 ppm 1.95 ppm 276 ppm 5.4 ppm 0 ppm
mg/min
t )75 min
mg/min
26.9 2.2 0.05 6.9 0.27 0
3.19% 1.69% 1.70% 12900 ppm 592 ppm 11 ppm 475 ppm 5.5 ppm 25 ppm
108 4.6 0.08 1.81 0.0009 0.23
conditions, it was generally lower than 0.04%. The origin of phthalic anhydride is unclear, but one possibility would be impurities in the n-butane feed: n-pentane is known to react over VPO to produce phthalic anhydride. However, the butane purity was 99.5% with less than 0.4% i-butane, less than 0.1% propane, and only 0.05% other hydrocarbons; therefore, it would seem unlikely that n-pentane would be the origin of this acid. Several other unidentified peaks were present on the chromatogram, but because the response is so highly dependent upon the carbon-carbon bonds, estimating their relative abundance was not attempted. Figure 6 plots butane conversion versus acetic acid selectivity at different temperatures and inlet gas compositions. The tendency is similar to that for maleic acid: selectivity is generally higher at lower temperatures and decreases with increasing conversion. Under reducing conditions, the selectivity drops very rapidly with small increases in the conversion. Under oxidizing conditions, acetic acid selectivity is somewhat lower but is less sensitive to butane conversion; it is an order of magnitude lower than maleic acid and about 5 times higher than acrylic acid. Under reducing conditions, the acrylic acid selectivity can reach as high as 50% of the acetic acid selectivity but it is generally on the order of 20%. At high oxygen partial pressures, the acrylic acid selectivity is generally below 1% and it reaches 2.5% at high butane partial pressures. It increases with an increase in both the temperature and conversion, as shown in Figure 7, which is opposite to the trend for both maleic and acetic acids. This might indicate that its formation follows a completely independent pathway or that it may be a decomposition product of maleic acid. Some experimental data suggest that acrylic acid is indeed a maleic anhydride decomposition product. The effluent line was generally maintained at a temperature of 200 °C in order to avoid product condensation. In one test, it was heated to 360 °C for 30 min and brought back down to the standard temperature afterward. Acids were collected for an additional 55 min under identical operating conditions (5% C4, 10% O2, 410 °C, 3600 mL/min). Two different liquid samples were taken after 20 and 75 min and were analyzed by HPLC. The data are summarized in Table 4 and show some remarkable differences in product distribution. Exit butane concentrations were equivalent at both times, which would indicate that it was essentially operating at steady state. However, both the CO and CO2 concentrations were higher when the heating tape was at 360 °C; the CO2 concentration was twice as high, which indicates that some of the acids were combusting. On the basis of the change in acid production rates, clearly both the maleic anhydride and acetic acid combust: their formation rates are 4 and 2 times lower during
Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 6735
Figure 8. Simplified butane oxidation reaction network.
the time at which the heating tape was maintained at the higher temperature. 3.2. Kinetic Modeling. In general, the experimental evidence in the published literature suggests that vanadium pyrophosphate [(VO)2P2O7 (VPO)] is the active phase for selectively reacting butane to maleic anhydride. The reaction is very complex and involves abstraction of eight hydrogen atoms and insertion of three oxygen atoms into the four-carbon skeleton: 14 electron transfer. Although this phase has been well characterized, its performance may depend on the preparation method, and many patents have been granted declaring an optimal manufacturing method. Together with the precursor recipe (reducing agent and solvent), other factors that influence the VPO catalyst performance include the P/V ratio, promoters, vanadium oxidation state, and calcination/activation procedure. The relation between these parameters and the catalytic performance was reviewed by Hodnett,18 and it can be concluded by the different studies that the strength of the reducing agent determines the average oxidation state of the precursor and therefore determines the activation temperature. The P/V ratio influences the redox properties of the catalyst, the phase composition, and the distribution of vanadium oxidation states in the catalyst. The VPO precursor is prepared in either an aqueous or organic medium. The latter usually results in a lower oxidation potential of V4+ to V5+, higher surface area, and higher maleic anhydride yields.19 The P/V ratio strongly influences the oxidation state as well as the selectivity to maleic anhydride. P/V ratios on the order of 1.1:1 and bulk catalyst oxidation states near 4+ are considered optimal. Phosphorus appears to modulate the catalyst activity, and as it is depleted with time, the activity generally increases and the selectivity decreases. A simple triangular reaction scheme is frequently adopted to describe the reaction mechanism, as shown in Figure 8. Butane may react to form either maleic anhydride or CO/CO2 (COx). Maleic anhydride may also combust to form COx. This scheme clearly neglects byproduct acid evolution, which may account for as much as 10% of the total under certain conditions. Moreover, it oversimplifies the variation of the carbon oxide selectivity with reaction conditions. For example, under oxidizing conditions, ratios of CO/CO2 may run as high as 1.6:1.8. A similar CO/CO2 ratio was also observed by Buchanan and Sundaresan15 using a VPO catalyst with P/V equal to 1. They reported a decreasing of the CO/CO2 ratio with increasing conversion and decreasing temperature. A CO/CO2 higher than 1 was also observed by Wohlfahrt and Hofmann.20 In our case, the ratio drops to below 1.2 as the conditions become more reducing and tend to 0.9 at a butane concentration of 5-10%. Lumping CO and CO2 together when modeling the kinetics is not strictly correct, but because the carbon-surface oxide reaction mechanism is so complicated, it is convenient to ignore side reactions and lump the carbon oxides. It is generally accepted that the surface activation of butane is the rate-limiting step and that a redox model best represents the reaction mechanism: Butane reacts
with a surface oxygen and the subsequently reduced catalyst is reoxidized by molecular oxygen. Such redox mechanisms can be formally described by the kinetic equations of Mars-van Krevelen or Eley-Rideal. Several authors18,21-23 report the presence of different vanadium oxidation states under reaction conditions, and several reaction mechanisms have been proposed. Some of these assume both the selective oxidation of butane to maleic anhydride and maleic anhydride decomposition taking place on the same active site,18,22 while poisoning studies revealed24 that these reactions occur on two different sites corresponding to V4+and V5+ oxidation states. Although it is now generally accepted that the presence of both V4+and V5+ is required during the reaction, the role of the different oxidation states of vanadium is still unclear.25 Hutchings et al.26 demonstrate a distinct parallel between improving the catalytic performance and a decrease in the amount of V5+ present. Many intrinsic rate expressions have been developed to describe both the selective and nonselective reaction kinetics. Some of the expressions are entirely empirical, while others are based upon Langmuir-Hinshelwood and redox-type mechanisms. Table 5 summarizes a number of different kinetic models proposed in the literature, as well as the reactor and catalyst type and the experimental conditions. Hydrodynamic Modeling. To accurately assess the reaction kinetics, an adequate transport model for the gas and solid phases is essential. Both steady-state and transient techniques are commonly employed to quantify the gas-phase hydrodynamics. In this work, we employed a transient step response method. A constant flow of nitrogen was fed to the reactor, and then it was switched to air. The oxygen concentration was monitored online at a frequencey of 1 Hz. This step change to the input function is referred to as the heavyside unit step function and is shown in Figure 9. Time “0” is the moment at which the valve is turned. As the air sweeps across the catalyst, the oxygen concentration increases gradually to its steady-state value of approximately 21%. At low gas velocities, the reactor is operating in the fixed-bed fluid regime and axial dispersion is quite low. At 6 mm/s, the bed operates between the minimum fluidization and minimum bubbling fluidization regimes. The bed is in the bubbling fluidized regime at 12 and 18 mm/s, and these were the velocities at which most kinetic experiments were conducted. Cold flow experiments (atmospheric pressure and room temperature) were made in a quartz reactor (4.1 cm i.d.) at different gas velocities in the fixed-bed and bubbling regimes, a second set of experiments were made in the Hastelloy reactor at temperatures of 350, 380, and 410 °C, pressures of 1, 3, and 5 bar abs, and different amounts of catalyst, and a third set was conducted to evaluate the transport effects due to the connecting tubing and empty reactor volume. All tests were repeated three times at each gas velocity, and the plots are superimposed in Figure 9. Reproducibility was very good. Figure 10 shows the first derivative of the step response experiments and represents the Dirac δ input function. In the case of ideal gas flow (fixed beds), the response curve is symmetrical around the average residence time. At low gas velocities, 0.3 cm/s, the response curve is nearly symmetrical. At higher gas velocities, the curves are less symmetrical and have a
6736 Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 Table 5. Kinetic Models for the Selective Oxidation of Butane to Maleic Anhydride over VPO Catalyst catalyst
T, K
O2
C4H10
kinetic expressions r1,2 )
organic mediuma,15
663-713
21%
0.62-1.55%
r3 )
r1 ) organic mediuma,16
573-613
O2/C4H10 ) 15
0.32-1.1%
k1,2CB K1CB K2CMA 1+ + CO2 CO 2
k3CMA K1CB K2CMA 1+ + CO 2 CO 2 k1KBCBCO2R 1 + KBCB
r2 ) k2CO2β r3 ) k3CMA
r1 ) industrial catalyst?a,27
573-653
21%
e3%
693-713
13%
0.56-2.38%
V2O5/P2O3a,31,32
673-753
21%
