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Pressure Calcination of VPO Catalyst Gregory S. Patience,*,† Richard E. Bockrath, John D. Sullivan,‡ and Harold S. Horowitz‡ E.I. du Pont de Nemours & Co., Wilmington, Delaware 19880
Calcination and activation are critical steps in the transformation of vanadyl hydrogen phosphate hemihydrate (VPO precursorsVOHPO4‚0.5H2O) to vanadyl pyrophosphate (VPO catalysts(VO)2P2O7). Whereas many published kinetics studies for butane oxidation over VPO demonstrate the relationship between catalytic activity and reaction conditions, the relationship between calcination conditions and activity has only received superficial attention in the open literature. This study delineates the operating boundaries in which to optimize catalytic performance. The control variables included temperature, pressure, time, and gas-phase composition (principally, oxygen and water vapor partial pressures), and their effects on surface area, oxidation state, and catalytic activity were measured. The experimental plan included several hundred calcination experiments. The commercial-design operating conditions were chosen as the base-case conditions: temperatures from 330 to 460 °C, pressures up to 6 atm, and oxygen and water concentrations between 0 and 20% and 0-3%, respectively. The optimal temperature was found to be about 390 °C. Catalyst performance declines with increasing pressure, but the pressure-time relationship allows a range of conditions: optimal catalytic performance was achieved within 20 h of calcination at atmospheric pressure, but it took less than 4 h to achieve this high level of performance when precursor was calcined at 3.5 atm (at longer times, the catalytic performance deteriorated). Co-feeding steam (in the presence of oxygen) generally had a deleterious affect on catalyst performance and, in particular, maleic anhydride yield. Moreover, oxidation states higher than 4.5 were achieved with steam/oxygen mixtures in as little as 20 h. Catalyst performance declined linearly with oxidation state for precursor calcined to vanadium oxidation states exceeding 4.5. Introduction Several technologies have been commercialized to produce maleic anhydride from n-butane and include fixed beds, fluid beds, and circulating fluidized beds (CFBs). DuPont’s CFB process involved cycling vanadium-phosphorus-oxide catalyst (VPO) between two vessels that were operated with a net reducing environment and an oxidizing environment, respectively.1 VPO catalyst was oxidized in an air-fed vessels regeneratorsoperating as a fluidized bed. It was transferred to a transport-bed reactor and reacted with n-butane to form maleic anhydride. The gas-solids suspension was carried upward though a riser at velocities approaching 6 m/s. The gas was separated from the catalyst in a stripper/rough-cut cyclone. The reaction gases passed through a second cyclone, filters, and, finally, an absorber where the anhydride was hydrolyzed to maleic acid. The VPO descended through the stripper at essentially minimum fluidization conditions and then through a standpipe, and finally the catalyst entered the top of the regenerator. The plant was operated at relatively low singlepass butane conversion (below 40%), and to achieve high conversion rates, unreacted butane was recycled. Catalyst manufacture represented a significant investment, and several programs were initiated to optimize the process. Two formulations are commonly used to synthesize VPO precursorsVOHPO4‚0.5H2O: one relies on an aqueous preparation route, and the second is based on an organic solvent route with isobutanol and benzyl alcohol. The former is generally considered to be more active and less selective. The precursor * To whom correspondence should be addressed. Tel.: (514) 3404711 ext. 3439. Fax: (514) 340-4159. E-mail: gregory-s.patience@ polymtl.ca. † Department of Chemical Engineering, Ecole Polytechnique de Montreal, C.P. 6079, Succ. “CV”, Montre´al, Que´bec, Canada H3C 3A7. ‡ INVISTA, Wilmington, Delaware.
is typically about 60 µm after washing and drying and is comprised of small platelets. For fluid-bed catalysts, the second step in catalyst manufacture is to micronize the precursor to approximately 1-2 µm followed by spray-drying and calcination and activation-transformation of the vanadyl hydrogen phosphate hemihydrate to vanadyl pyrophosphates(VO)2P2O7. Although many investigations reported in the literature describe the calcination and activation procedures, a systematic study relating the operating conditions to catalytic performance is currently lacking. The transformation of the hemihydrate phase to the active (and selective) pyrophosphate can either be carried out in situ (under reaction conditions with butane) or ex situ in various gaseous environments with air, oxygen, nitrogen, and steam.2 Whereas only limited attention has been devoted to the calcination conditions in the open literature, several patents have tested the effect of gas-phase composition and temperature on catalyst performance.3,4 Most treatments specify air as the preferred gas for an average time of about 3 h and a temperature of 480 °C. The temperatures published in the literature range from 240 < T < 800 °C, and the time goes from between 30 min and 168 h. Besides air, oxygen, air plus water vapor, nitrogen, and mixtures of air and nitrogen have all been tested. O’Mahony et al.5 suggest that a thermal treatment with a mixture of n-butane in air (e.g., in situ) yields the best performance. Calcining and activating catalyst in large-scale commercial facilities presents some pragmatic design and engineering constraints. Calcining in situ has certain practical advantagess standard operating procedures may be followed and plant utility rates may be higher. However, some published studies recommend low butane concentrations2,5,6 (e1%), with a correspondingly lower production rate. Furthermore, it may take several hundred hours to achieve a stable, fully activated catalyst with the optimal morphology.7 Long calcination and activation procedures represent down-time for the process and, thus, loss
10.1021/ie0613016 CCC: $37.00 © 2007 American Chemical Society Published on Web 05/16/2007
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of productivity during start-up or when more precursor is required to make up for catalyst lost due to attrition (for the case of fluid-bed reactors). A continuous process is preferred compared to batch processing, and thus, the precursor should be injected into the fluid-bed vessel at reaction temperature, which results in an instantaneous hydrothermal shock for the hemihydrate. Direct-fired heaters are generally employed to heat up the reactors and maintain them at temperature (under standby conditions), and thus, water vapor partial pressures range between 1 and 3%. Most experimental studies rely on an ex situ procedure for calcination followed by a more or less prolonged “activation” with butane. The published range of conditions for activation is as broad as for calcination: temperatures range from 330 to 680 °C, time varies from 0.5 to 380 h, the partial pressure of oxygen varies between 0 and 100%, butane partial pressures go from 0 to 100%, and even 100% hydrogen has been reported. From an industrial perspective, the preferred process is one in which a stable, active, and selective VPO is injected directly into the plant. Therefore, ex situ calcination/activation is preferred. However, many of the conditions for activation (as well as calcination) discussed in the open literature would be economically unattractive for a catalyst manufacturer. Both gas temperature and residence times should be minimized. Additional operational costs (and potentially design costs) are incurred with elevated concentrations of oxygen or when low concentrations are specified such that significant quantities of nitrogen are required. Processing ex situ off-line would require permits, facilities to oxidize CO, an absorber to treat the water and acids produced during calcination (the precursor can lose up to 18% of its weight in water and organics), and an absorber sufficiently large to handle significant maleic anhydride production when butane is specified for activation. In this paper, we demonstrate the effect of calcination time, temperature, pressure, and gas composition on catalyst activity, vanadium oxidation state, and stability. The selected conditions for calcination were chosen to resemble those achievable in a commercial plant, and over 200 samples were generated (approximately 15 kg of VPO). Precursor was injected into a heated fluidized bed to simulate the expected thermal history expected for a continuous on-line process, and the reactor was co-fed up to 5% water vapor to simulate the resulting gas composition expected due to the direct-fire burner. Experimental Section Three laboratory-scale reactors were employed to examine calcination-activation protocol: catalyst was calcined and activated in a pressurized fluidized-bed vessel, and the performance was evaluated in a microfixed-bed reactor and a 1/4 in. diameter riser reactor. In addition to reactor testing, the samples were analyzed by X-ray diffraction, Brunauer-Emmett-Teller (BET) surface area, carbon content, and vanadium oxidation state. Pressurized Fluid-Bed Reactor. The pressurized fluid bed was a 41 mm i.d. by 844 mm stainless steel pipe with a 78 mm i.d. flanged section (to minimize entrainment to filters installed at the top of the vessel), as shown in Figure 1. It was rated for operation up to 600 °C and 10 atm and was heated by an external fluidized-sand bath. Gas entered through a porous plate distributor at the bottom of the reactor and exited through a sintered metal filtered at the top. Feed gases were controlled with rotameters; an Isco pump fed water to a steam generator that was heated to 170 °C. Feed gases passed through a 1/4 in. diameter tube coiled around the reactor in the sand bath to
Figure 1. Pressurized fluid-bed reactor.
preheat the gases to reactor conditions. Acids and water were condensed in an absorber at the exit of the reactor and then passed through a demister before exiting to the exhaust. Two protocols were tested in the unit. During the initial development stages, 0.2-1 kg of catalyst was loaded into the reactor at ambient conditions and then the temperature of the unit was ramped up to the desired temperature. However, in order to better simulate “efficient” commercial calcination conditions, a solids injector was fabricated to allow the pulse injection of cold catalyst into the vessel that was maintained at the desired calcination conditions. In this procedure, batches of 10 g of catalyst were loaded into a hopper that was then pressurized with 7 atm house air. The differential pressure between the hopper and reactor vessel propelled the solids downward into the bed. A catalyst feed rate of 400 g/h was achievable with this apparatus. Several variables were selected in order to quantify their effect on maleic yield, including temperature, pressure, gas velocity, oxygen concentration, and steam partial pressure. Standard runs would last up to 5 days, and calcined catalyst was removed on-line through an in situ sampling port to monitor and establish the change in physicochemical properties with time. After several initial experiments to explore the variable space and develop standard operating procedures, a set of calcination conditions were defined as the base case for performance comparisons. The standard conditions selected were 390 °C, 3.5 atm, and a feed gas comprising 3% steam in air. A typical catalyst charge was 300 g that was pulse-injected. Up to 20 samples were withdrawn during the course of the 30-100 h calcination experiments. Steam was used during calcination to simulate the concentration exiting the direct-fired burner. Together with identifying the conditions for which maleic yield was maximum, the experiments were essential to define a safe operating window in which to maintain the catalyst in the case of a plant upset during the calcination and activation as well as standby conditions. The range of conditions selected corresponded to those that would be possible in the commercial CFB reactor are shown in Table 1. 1/4 in. Riser Reactor. The reactor used to evaluate catalyst samples was made of quartz, and the riser was 1/4 in. i.d. It
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Figure 3. Cumulative weight loss during calcination of 300 g of precursor subjected to a 2 °C/min temperature ramp. Table 2. Calculated Pseudo-First-Order Rate Constants of Pilot-Plant Catalyst
Figure 2. 1/4 in. quartz riser reactor. Table 1. Experimental Conditions (Variables and Levels) P, atm T, °C O2 inlet, vol % H2O inlet, vol % t,h
1, 3.5, 6 330, 360, 375, 390, 420, 460 0, 2, 5, 21 0, 2, 3 0.5-100
was developed to simulate contact times and gas compositions expected in the commercial reactor. It consisted of three basic vessels, as shown in Figure 2: transport bed/riser, stripper, and regenerator. Catalyst inventory in the reactor could be varied over a range from as low as 0.4 kg to about 2 kg. The catalyst was reduced in the riser with mixtures containing high concentrations of butane; the gas phase and adsorbed carbon were separated from the catalyst in the stripper; and finally, the catalyst was reoxidized by air in the regenerator. An in situ activation protocol was developed to evaluate long-term catalyst performance. A test would last up to 250 h, and it was fed 10% butane at a temperature of 380 °C. In order to compensate for catalyst losses during the experiment (due to attrition and cyclone inefficiency), up to 20% more catalyst was added versus a standard protocol test. To maintain consistent reactor conditions, the additional inventory was accumulated in the stripper that was fed inert gas so that the regenerator residence time was held constant. As the solids inventory level decreased, the catalyst level in the stripper fell. Multiple Automated Reactor System (MARS). Well over 200 samples of catalyst were generated in the pressurized fluidbed reactor. Testing all of these samples in the [1/4 in.] riser was not an option because of the large inventory requirements and time constraints. Therefore, all samples were evaluated in a microreactor systemsMARS. The system was equipped with a single feed manifold that delivered a uniform gas composition to as many as six microreactors simultaneously. Mass flow controllers regulated the gas compositionsbutane, oxygen, and nitrogensto the desired level. A gas mixture of 10% butane in nitrogen was employed instead of pure butane to avoid condensation and heating problems associated with a liquid feed. The reactors were controlled with a Siemens (TI545) process logic controller (PLC). The PLC was programmed with failsafe interlocks to allow unattended and overnight operation. Process variables were continually recorded with a data acquisition system (VANTAGE) for subsequent analysis. A series of tests were conducted in order to formulate a protocol in the microreactors that would correlate with
time-on-stream, days
k, 1/4 riser
k, MARS
0 10 29 43
0.13 0.10 0.084 0.083
0.13 0.10 0.093 0.087
performance in the 1/4 in. riser. A basis of comparison was developed by calculating a pseudo-first-order rate constant in the MARS and comparing the value against that calculated in the riser. Table 2 compares the pseudo-first-order rate constants for four samples withdrawn from a pilot plant for the MARS reactor and the riser. In both reactor systems, maleic anhydride yield declined with time-on-stream. The MARS was operated at 380 °C and a gaseous feed composition of 9% butane and 10% oxygen. Catalyst. The VPO precursor was synthesized on a commercial scale in an organic medium with isobutanol and benzyl alcohol. It was micronized to 1-2 µm, then spray-dried with polysilicic acid to form a porous silica shell containing up to 10% silica.1 Spray-dried catalyst is somewhat less attritionresistant than calcined-activated catalyst but sufficiently resistant to withstand 100 h in the fluid bed. A narrow fraction of the catalyst particles were analyzed, between about 80 and 150 µm as measured by a Coulter counter. The particle density was on the order of 1900 kg/m3, and the skeletal density, measured by helium pycnometry, was almost 3000 kg/m3. Results and Discussion Injecting precursor into a hot vessel has a clear advantage with respect to minimizing cycle time versus applying a temperature ramp. However, most experimental work published in the open literature and patent literature is conducted at atmospheric pressure with a temperature ramp. In this study, a few trials were made with a temperature ramp of 2 °C/min, and the weight loss of the precursor was monitored by condensing the produced gases in a quench. The liquid samples were subsequently analyzed by high-performance liquid chromatography (HPLC). In one experiment, of the original 300 g sample loaded into the reactor, 41 g of liquid was recovered in the quench. The highest weight loss was recorded at between 340 and 370 °Cs75% of the total was collected during this interval, as demonstrated in Figure 3. O’Mahony et al.5 subjected precursor to a 1 °C/min temperature ramp and found that the hemihydrate fully decomposed at 300 °C into an amorphous solid. They showed that, at 350 °C and based on X-ray diffraction (XRD) analysis, a new phase started to develop, and that it was fully developed at 400 °C. In pure oxygen, the hemihydrate decomposed at 270 °C. These temperatures are
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Figure 4. Precursor oxidation rate as a function of time and feed gas concentration at a pressure of 3.5 atm and 390 °C. Table 3. Chemical Analyses of Condensed Off-Gases during Calcination acid
ppm
maleic acetic phthalic acrylic methacrylic fumaric
2800 552 399 81 56 3
somewhat lower than those found herein and perhaps are related to the temperature ramp. Maleic acid and phthalic acid were detected in all liquid quench samples. Acetic acid was the second most abundant acid after maleic acid. Chemical analyses of the acids produced during injection into a hot fluidized bed are given in Table 3. All of these acids have been detected under standard operating conditions with butane.8 In this work, the principal calcination variables studied included temperature, pressure, and oxygen and water vapor partial pressures. For each experiment, the calcination conditions were studied as a function of time by withdrawing a 10 g sample of catalyst on-line. After initial testing and with some consideration toward the constraints of an industrial process, a base case set of conditions was defined: temperature ) 390 °C, pressure ) 3.5 atm, and 3% water vapor in air. The use of steam is mandated by commercial heating constraints. A direct-fired burner is the most economical means of heating industrial vessels, which can result in non-negligible concentrations of water. Effect of Steam. The presence of steam significantly reduces the acceptable calcination operating window and is one of the primary causes of poor catalyst performance. At all temperatures and pressures investigated, long-term exposure to steam in air leads to a highly oxidized catalyst with low maleic anhydride productivity. As has been reported previously, VOHPO4‚0.5H2O prepared in alcohol media shows evidence of residual organic matter trapped between its basal planes; the presence of this entrapped organic impedes conversion to the (VO)2P2O7, both by disruption of the solid-state transformation mechanism and its reducing ability.2,9-11 Oxygen is, therefore, necessary to aid the removal of the organic for the transformation of the hemihydrate phase to the pyrophosphate, but overoxidation only occurs when steam makes up part of the calcination feed gas composition. Overoxidation correlates strongly with poor catalytic performance. Steam appears to enhance and accelerate the formation of pentavalent vanadium phases. In the absence of steam, the maximum vanadium oxidation state achieved during calcination
Figure 5. Decline of surface area and maleic anhydride productivity (gC/ h/kgcat) as a function of oxidation state of precursor calcined at 390 °C, 3.5 atm, and 3% steam in air.
is about 4.5, whereas oxidation states as high as 4.9 were reached with steam/air mixtures. Figure 4 illustrates the oxidation state as a function of time for two samples calcined at 390 °C and 3.5 atm. Within 2 h, the oxidation state of the catalyst increased from about 4.0 to 4.35 under conditions with 3% steam and in the absence of steam in the feed gases. Thereafter, the oxidation state of the sample with steam co-feed increased steadily to 4.8 at 70 h. The oxidation state of the sample with no steam in the feed rose to approximately 4.5 after 20-30 h. Additional experiments at various distinct conditions demonstrated that the limiting oxidation state with no steam in the feed gases was 4.5. In one case, precursor was calcined in air for 150 h, at which point the oxidation state was 4.51. After the 150 h period, steam was introduced, and after an additional 60 h, the oxidation state rose to 4.81. There is a strong correlation between overoxidation and the steady-state maleic anhydride productivity. We prepared several dozen samples at standard conditions of 390 °C, 3.5 atm, and air with water vapor concentrations varying between 1.5 and 5%. The oxidation states at the end of the calcination varied between 4.1 and 4.9. Below an oxidation state of about 4.5, there was no evident correlation between maleic anhydride productivity (as measured in the MARS after 80 h) and the initial oxidation state. However, irrespective of the calcination conditions, the yield decreased linearly with increasing initial oxidation state. Combined with a decline in the performance of steam calcined precursor is a loss in the surface area. At an oxidation state of 4.5, surface area was on the order of 14-18 g/m2 and it dropped linearly to 5-8 m2/g at an oxidation state of 4.8, as shown in Figure 5. X-ray diffraction data on samples calcined for short periods of time ( 60 h) calcination conditions
oxidation states, Vox
T, °C
P, atm
initial
after MARS evaluation
390 390 390 390 420 420 460 460 360 360 360 390 390 390 390 390 390 390
3.5 3.5 3.5 3.5 3.5 3.5 3.5 3.5 3.5 3.5 3.5 6.0 6.0 6.0 1.0 1.0 1.0 1.0
4.53 4.61 4.71 4.76 4.45 4.66 4.53 4.81 4.19 4.71 4.84 4.36 4.50 4.83 4.37 4.53 4.54 4.62
4.12 4.05 4.43 4.63 4.10 4.38 4.36 4.36 4.08 4.44 4.62 4.30 4.43 4.74 4.00 4.02 4.03 4.13
first 50-100 h before stabilizing. (Samples were withdrawn online, so the source material and calcination history were identical). In addition to maleic yield, pressure has a significant effect on oxidation rate and maximum oxidation state, Vox,max. A plot of the vanadium oxidation state as a function of time at 1, 3.5, and 6 atm, is shown in Figure 9. This plot indicates that lowering pressures could be an effective means of expanding the operating window during calcination or for start-up, shutdown, or standby. Figure 10 is an expanded view of the oxidation state as a function of time. There is some spread in the data, but it would appear that, below an oxidation state of 4.5, oxidation rate was independent of pressure. The data were characterized with a first-order reaction model, and the best-fit rate constant was equal to 0.37 h.
dVox ) kVox dt Effect of Oxygen Partial Pressure. The optimal partial pressure of oxygen remains to be evaluated thoroughly. The preceding discussion around pressure, in fact, confounds two effects: total pressure and oxygen partial pressure. We have shown that the presence of steam in air may have a detrimental affect on catalytic performance for extended calcination periods. In this section, we explore the effect of low oxygen concentrations (in nitrogen)s0%, 2%, 5%, and 20%son catalytic performance with 3% steam and at 390 °C and 3.5 atm. In the absence of oxygen, 150 h of treatment was insufficient to transform the precursor into the active pyrophosphate phase. X-ray diffraction at the end of the test showed that the sample was completely amorphous and there was little to no maleic anhydride detected during the MARS evaluations. The oxidation state remained unchanged above 3.9. (Presumably, some of the organic solvents were oxidized, thereby reducing the precursor). The limited data collected to this point are inconclusive. As shown in Figure 4, a feed gas composition of 5% oxygen/3% steam appears to provide the same oxidation rate as air/3% steam. However, the oxidation rate for one sample was slowed by reducing the feed composition to 2% oxygen and 2% steam after 10 h at 3% steam in air. We showed that oxidation rates may be slowed by reducing pressure, which is not optimal for plant operations for standby conditions. However, reducing the partial pressure of oxygen is often possible by diluting the gas stream with nitrogen; therefore, standby conditions should be
specified to be
