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Applied Catalysis A: General 221 (2001) 397­419

Dehydrogenation and oxydehydrogenation of paraffins to olefins

M.M. Bhasin a, , J.H. McCain a , B.V. Vora b , T. Imai b , P.R. Pujad´ b o


Union Carbide & UOP LLC, Subsidiary of Dow Chemical Co., P.O. Box 8361, 25303 South Charleston, WV, USA b UOP LLC, Subsidiary of Dow Chemical Co., P.O. Box 8361, 25303 South Charleston, WV, USA

Abstract Catalytic paraffin dehydrogenation for the production of olefins has been in commercial use since the late 1930s, while catalytic paraffin oxydehydrogenation for olefin production has not yet been commercialized. However, there are some interesting recent developments worthy of further research and development. During World War II, catalytic dehydrogenation of butanes over a chromia-alumina catalyst was practiced for the production of butenes that were then dimerized to octenes and hydrogenated to octanes to yield high-octane aviation fuel. Dehydrogenation employs chromia-alumina catalysts and, more recently, platinum or modified platinum catalysts. Important aspects in dehydrogenation entail approaching equilibrium or near-equilibrium conversions while minimizing side reactions and coke formation. Commercial processes for the catalytic dehydrogenation of propane and butanes attain per-pass conversions in the range of 30­60%, while the catalytic dehydrogenation of C10 ­C14 paraffins typically operates at conversion levels of 10­20%. In the year 2000, nearly 7 million metric tons of C3 ­C4 olefins and 2 million metric tons of C10 ­C14 range olefins were produced via catalytic dehydrogenation. Oxydehydrogenation employs catalysts containing vanadium and, more recently, platinum. Oxydehydrogenation at 1000 C and very short residence time over Pt and Pt-Sn catalysts can produce ethylene in higher yields than in steam cracking. However, there are a number of issues related to safety and process upsets that need to be addressed. Important objectives in oxydehydrogenation are attaining high selectivity to olefins with high conversion of paraffin and minimizing potentially dangerous mixtures of paraffin and oxidant. More recently, the use of carbon dioxide as an oxidant for ethane conversion to ethylene has been investigated as a potential way to reduce the negative impact of dangerous oxidant­paraffin mixtures and to achieve higher selectivity. While catalytic dehydrogenation reflects a relatively mature and well-established technology, oxydehydrogenation can in many respects be characterized as still being in its infancy. Oxydehydrogenation, however, offers substantial thermodynamic advantages and is an area of active research in many fronts. © 2001 Elsevier Science B.V. All rights reserved.

Keywords: Paraffin dehydrogenation; Olefin; Chromia-alumina catalyst; Paraffin oxydehydrogenation; Noble metal catalysts

1. Historical overview and chromia-alumina catalysts Paraffin dehydrogenation for the production of olefins has been in use since the late 1930s. During

Corresponding author. Tel.: +1-304-747-4910; fax: +1-304-747-5430. E-mail address: [email protected] (M.M. Bhasin).

World War II, catalytic dehydrogenation of butanes over a chromia-alumina catalyst was practiced for the production of butenes, which were then dimerized to octenes and hydrogenated to octanes to yield high-octane aviation fuel. Dehydrogenation of butanes over a chromia-alumina catalyst was first developed and commercialized at Leuna in Germany and was also independently developed by UOP (then Universal Oil Products) in the

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M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419

United States, together with ICI in England. The first UOP-designed plant came on stream in Billingham, England, in 1940 and was soon followed by two other units in Heysham, England, in 1941 [1]. The primary purpose of this butane dehydrogenation was to produce butenes, which were then dimerized to octenes using solid phosphoric acid catalysts discovered by Schaad and Ipatieff [2]. Other companies soon followed these pioneering efforts. For example, Phillips Petroleum built a multitubular dehydrogenation reactor near Borger, TX, in 1943 [1]. However, the most significant development was made by Houdry using dehydrogenation at less than atmospheric pressure for higher per-pass conversions. This process, which came on stream toward the end of World War II, was also used for the production of butenes. After the war, Houdry further developed and commercialized the chromia-alumina dehydrogenation system and extended it to the production of butadiene in what became known as the CatadieneTM process [3]. Other companies, including Shell, Gulf, and Dow, also practiced similar dehydrogenation technologies. In the dehydrogenation process using chromiaalumina catalysts, the catalyst is contained in a fixed shallow bed located inside a reactor that may be either a sphere, a squat vertical cylinder, or a horizontal cylinder. The actual design reflects a compromise between gas flow distribution across a large cross-sectional area and the need to maintain a low pressure drop. A significant amount of coke is deposited on the catalyst during the dehydrogenation step, therefore, a number of reactors are used in parallel--some for dehydrogenation while the rest are being purged or regenerated. The dehydrogenation reactions are strongly endothermic, and the heat is provided, at least in part, by the sensible heat stored in the catalyst bed during regeneration (carbon burn); additional heat is provided by direct fuel combustion and also by heat released in the chromium redox cycle. The length of the total reactor cycle is limited by the amount of heat available, and can be as short as 10­20 min. The Houdry Catadiene process was used extensively for the production of butadiene, either by itself (n-butane to butadiene) or in conjunction with catalytic oxydehydrogenation of n-butene to butadiene. The latter was commercialized by the Petro-Tex Chemical Corp. [3] and was called the Oxo-DTM

process. A similar oxydehydrogenation approach for the production of butadiene was also practiced by Phillips Petroleum [3]. Large quantities of butadiene have become available over the past 30 years, mostly as a by-product from the thermal cracking of naphtha and other heavy hydrocarbons. This market shift has resulted in the shutdown of all on-purpose catalytic dehydrogenation units for butadiene production in North America, western Europe, and the far East. In the late 1980s, the application of chromia-alumina catalysts was extended by Houdry to the dehydrogenation of propane to propylene and isobutane to isobutylene. The new process application called CatofinTM [4,5] operates on the same cyclic principle as in the former Catadiene process. As of late 2000, a total of eight Catofin units exist for the production of isobutylene (including two converted older Catadiene units) with an aggregate capacity of about 2.8 million metric tons per annum (MTA) isobutylene. In addition, two Catofin units were built for the production of propylene, but it is understood that only one is operational with a nameplate dehydrogenation capacity of about 250,000 MTA propylene, but usually operating only on a seasonal basis. Plans for another 450,000 MTA Catofin propane dehydrogenation unit in Saudi Arabia have also been announced. The Catofin process technology is currently owned by S¨ d-Chemie and is u offered for license by ABB Lummus. In 1959, an alternative chromia-alumina catalytic dehydrogenation process was developed in the former Soviet Union. This process avoided the use of the cyclic operation by using a fluidized bed reactor configuration similar to the fluidized catalytic cracking (FCC) process used in refineries [6]. However, backmixing common to dense fluidized bed operations results in poor selectivity and increases the formation of heavies, sometimes called "green oils". Circulating regenerated catalyst is used to provide the heat of reaction in the riser and spent catalyst is reheated by carbon burn in the regenerator. During the 1990s, a large scale fluid bed isobutane dehydrogenation unit for about 450,000 MTA isobutylene was commercialized by Snamprogetti in Saudi Arabia based on technology from Yarsintez in Russia [6], but it is understood that this unit has only operated at lower than design capacity. Recent literature articles report further improvements by Snamprogetti [7,8].

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2. Noble metal dehydrogenation catalysts A different approach to catalytic dehydrogenation was first introduced in the mid-1960s for the supply of long-chain linear olefins for the production of biodegradable detergents. Synthetic detergents, based on the use of branched alkylbenzene sulfonates derived from propylene tetramer and benzene, had been introduced in the 1940s. By the early 1960s, however, it became apparent that branched dodecylbenzene-based detergents, though very active and offering excellent performance characteristics, did not biodegrade readily and were accumulating in the environment. The need for biodegradable detergents prompted the development of catalytic dehydrogenation of long-chain linear paraffins to linear olefins. The work on catalytic reforming with noble metal (Pt) catalysts done in the 1940s by Haensel clearly demonstrated that Pt-based catalysts had high activity for the dehydrogenation of paraffins to the corresponding olefins [9]. In the 1960s, Bloch [10] further extended this thinking by developing Pt-based catalysts that could selectively dehydrogenate long-chain linear paraffins to the corresponding internal mono-olefins with high activity and stability and with minimum cracking. This was the basis for the UOP PacolTM process for the production of linear olefins for the manufacture of biodegradable detergents [11]. In 1999, there were more than 30 commercial Pt-catalyzed dehydrogenation units in operation for the manufacture of detergent alkylate. Long-chain paraffins are both valuable and highly prone to cracking. Therefore, in order to maintain high selectivity and yield, it is necessary to operate at relatively mild conditions, typically below 500 C, and at relatively low per-pass conversions. While this is economical for the production of heavy linear olefins, it is not for the production of light olefins. Paraffin dehydrogenation is an endothermic reaction that is limited by chemical equilibrium and, according to Le Chatelier's principle, higher conversion will require either higher temperatures or lower pressures. In a somewhat abbreviated form for the production of mono-olefins, this can be expressed as follows:

2 xe =

Fig. 1. Propane dehydrogenation equilibrium at 1.00 atm abs. pressure.

where xe is the equilibrium conversion, P the total absolute pressure and Kp is the equilibrium constant for the dehydrogenation reaction. The equilibrium constant can be easily calculated from Gibbs free energies as tabulated in the API 44 report or in similar sources of thermodynamic data. Figs. 1 and 2 illustrate the equilibrium conversion levels that can be obtained for propane at 1 and 0.23 atm abs. (175 Torr), respectively.

Kp Kp + P

Fig. 2. Propane dehydrogenation equilibrium at 0.23 atm abs. pressure.


M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419

Fig. 3. Equilibrium constants for n-paraffin dehydrogenation at 500 C.

The equilibrium constant for paraffin dehydrogenation increases significantly as the carbon number increases. Fig. 3 shows the equilibrium constant for the dehydrogenation of n-paraffins ranging from ethane to pentadecane [12]. Fig. 4 shows the temperatures required to achieve 10­40% equilibrium conversion based on these equilibrium constants. Fig. 4 indicates that the temperature required for the dehydrogenation of light paraffins is much higher than for heavy paraffins. For 40% conversion, for example, the dehydrogenation of propane requires a temperature of at least about 580 C, while dodecane can be theoretically dehydrogenated to the same extent at only 450 C. The equilibrium conversion increases at higher

Fig. 4. Temperatures required to achieve 10 and 40% conversion of C2 ­C15 n-paraffins at 1 atm.

temperatures, but side reactions, coke formation, and catalyst deactivation are also accelerated. Thus, extrapolation directly from heavy olefins to light olefins cannot be done without taking other factors into consideration. Production of light olefins by the catalytic dehydrogenation of light paraffins must be able to maintain reasonable per-pass conversion levels and high olefin selectivity. Very importantly, it must be able to produce olefins in high yields over long periods of time without shutdowns. In the early 1970s, UOP introduced continuous catalyst regeneration (CCR) technology that enabled noble metal catalysts to remain at their most desirable stable activity for several years without having to shut down the reactor for catalyst regeneration. The combination of noble metal catalysts operating at high severity in conjunction with CCR technology made it possible to design, build, and economically operate large catalytic dehydrogenation units that can produce light olefins, in particular, propylene and isobutylene, at high selectivities while still operating at superatmospheric pressures. This technology is known as the UOP OleflexTM process. As of late 2000, there were four propane dehydrogenation units, five isobutane dehydrogenation units, and one combined propane/isobutane dehydrogenation unit of this type in commercial operation, with an aggregate operating capacity of 900,000 MTA polymer grade propylene and 2.3 million MTA isobutylene. In addition, another propane dehydrogenation unit for 350,000 MTA polymer grade propylene was under design and construction. The world propylene production capacity, based on the use of catalytic dehydrogenation of propane has increased steadily over the past 10 years [13] and is expected to grow even further under the right economic conditions relative to the availability of propane; on the other hand, environmental concerns on the use of MTBE are expected to adversely impact the future expansion of isobutane dehydrogenation applications. Although production of ethylene via catalytic dehydrogenation over Pt catalysts is very selective (about 95%), extension of this dehydrogenation technology to ethane has not taken place due to the need for even more severe operating conditions; higher temperatures and lower pressures. Such conditions cause excessive coking of the catalyst or require costlier operation under vacuum.

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Fig. 5. Reactions by platinum and acid sites in light paraffin dehydrogenation with unmodified catalyst.

Practically, all existing catalytic dehydrogenation capacity based on Pt catalysts is based on the Oleflex process with CCR; however, there are also two smaller units for isobutane dehydrogenation for 118,000 and 13,000 MTA isobutylene, respectively, both based on the STAR technology developed by Phillips Petroleum and derived from their earlier multitubular reactor design experience. This reactor design resembles a typical steam reformer that is operated until the catalyst deactivates as a result of coke deposition; banks of tubes are sequentially taken out of service for catalyst regeneration. The STAR technology is currently owned and licensed by Krupp­Uhde.

3. Process chemistry The main reaction in catalytic dehydrogenation is the formation of mono-olefins from the corresponding feed paraffin. Other reactions include consecutive and side reactions. The reaction pathways involved in heavy paraffin dehydrogenation (e.g. detergent-range C10 ­C14 n-paraffins) are more complicated than those in light paraffin dehydrogenation (e.g. propane and isobutane). The main difference in reaction pathways is that a significant amount of cyclic compounds can form via dehydrocyclization from heavy paraffins; this is not the case for light paraffins. Figs. 5 and 6

Fig. 6. Reactions by platinum and acid sites in heavy paraffin dehydrogenation with unmodified catalyst.


M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419

illustrate possible reactions that take place on Pt and acid (A) sites, respectively, in the dehydrogenation of light and heavy paraffins when the catalyst is not selective, e.g. unmodified platinum catalysts supported on alumina. The consecutive reactions, the dehydrogenation of mono-olefins to diolefins and triolefins, are catalyzed on the same active sites as the dehydrogenation of paraffins to mono-olefins. The consecutive reactions that form triolefins, aromatics, dimers, and polymers must be suppressed kinetically or by catalyst modifications.

4. Role of catalysts and supports The discussion in this section pertains to aluminasupported platinum catalysts. The work by Poole and coworkers [14,15] provides an extensive review of chromia-alumina catalysts. The key role of dehydrogenation catalysts is to accelerate the main reaction while controlling other reactions. Unmodified alumina-supported platinum catalysts are highly active but are not selective to dehydrogenation. Various by-products, as indicated in Figs. 5 and 6, can also form. In addition, the catalyst rapidly deactivates due to fouling by heavy carbonaceous materials. Therefore, the properties of platinum and the alumina support need to be modified to suppress the formation of by-products and to increase catalytic stability.

The reaction of olefins on platinum is faster than that of paraffins, because olefins interact with platinum more strongly than do paraffins. The role of platinum modifiers is to weaken the platinum­olefin interaction selectively without affecting the platinum­paraffin interaction. Arsenic, tin, germanium, lead, bismuth are among metals reported as platinum activity modifiers. The consecutive dehydrogenation rate of mono-olefins and diolefins is decreased by this modification without lowering the rate of paraffin dehydrogenation significantly. The modifier also improves the stability against fouling by heavy carbonaceous materials. Platinum is a highly active catalytic element and is not required in large quantities to catalyze the reaction when it is dispersed on a high surface-area support. The high dispersion is also necessary to achieve high selectivity to dehydrogenation relative to undesirable side reactions, such as cracking. The typical high surface area alumina supports employed have acidic sites that accelerate skeletal isomerization, cracking, oligomerization, and polymerization of olefinic materials, and enhance "coke" formation. Alkali or alkaline earth metals assist in the control of the acidity. Also, -alumina supports that have essentially no acidity can be utilized; however, the challenge is to obtain high dispersion of platinum on such very low surface area supports. Therefore, acidity must be eliminated by using suitable modifiers. Modified catalysts possess high activity and high selectivity to mono-olefins. The major by-products are diolefins that can be controlled kinetically. Coke

Fig. 7. Paraffin dehydrogenation on modified Pt catalyst.

M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419


formation is also suppressed and, therefore, stability is greatly improved. Over modified catalysts, the major reaction pathways for both light and heavy paraffin dehydrogenation systems are simpler (Fig. 7). Alumina has excellent thermal stability and mechanical strength under processing, transport, and catalyst regeneration conditions. However, the most important reason alumina is used as support material is its superior capability to maintain a high degree of platinum dispersion, which is essential for achieving high dehydrogenation activity and selectivity. The catalytic reaction rate is limited by the intraparticle mass transfer rate. If the mass transfer rate is relatively slow, both activity and selectivity are lowered. As a result, the support must have a low pore diffusional resistance (high effectiveness factor). For a given pore volume, the surface area and the strength of the support increase as the pore diameter decreases, and the pore diffusional resistance decreases as the pore diameter increases. Thus, an appropriate pore structure must be determined for the support to achieve optimal catalytic performance.

ratio (R). Therefore, Eq. (3) can be rewritten as f2 [ki (T ), Ki (T ), x, P , R] d(sx) = dx f1 [ki (T ), Ki (T ), x, P , R] (4)

As Eq. (4) is the ratio of two functions, the rate constants become relative values and can be expressed as ki (T)/k0 (T), where k0 (T) is the rate constant for the forward reaction of paraffin dehydrogenation to mono-olefins. Eq. (4) can be written in a functional form in F as follows: ki (T ) d(sx) =F , Ki (T ), x, P , R dx k0 (T ) (5)

5. Dehydrogenation catalyst evaluation In paraffin dehydrogenation, the rate of paraffin conversion (x) and mono-olefin production (sx) are given by Eqs. (1) and (2) respectively: dx = f1 (ki , Ki pj ) dt d(sx) = f2 (ki , Ki , pj ) dt (1)


where s is the selectivity to n-mono-olefins, t the contact time, f the rate function, ki the rate constant for reaction step i, Ki the equilibrium constant for reaction step i, and pj is the partial pressure of the j compound. The following relationship between selectivity and conversion can be derived from Eqs. (1) and (2): f2 (ki , Ki , pj ) d(sx) = dx f1 (ki , Ki , pj ) (3)

The ki and Ki are a function of temperature and pj is a function of conversion, total pressure (P), and feed

Eq. (5) indicates that selectivity is a function of conversion for the catalyst used (relative rate constants) and the given reaction conditions (temperature, pressure, feed ratio). Selectivity decreases as the conversion increases because n-mono-olefins are consecutively converted into by-products. Selectivity decreases sharply as conversion approaches equilibrium because the main dehydrogenation process is limited by equilibrium, but other reactions continue to occur. Therefore, if side reactions are controlled, the selectivity is improved as the equilibrium conversion becomes higher by increasing the temperature and by decreasing the pressure and the feed ratio of hydrogen to paraffin. The relationship between selectivity and conversion can be simulated according to Eq. (5), if rate functions, relative rate constants, and equilibrium constants are known. Fig. 8 shows simulated selectivities to n-heptene and n-heptadiene for the dehydrogenation of n-heptane. In this simulation, the relative rate constants used are unity, which represents that the catalyst possesses perfect selectivity regarding consecutive dehydrogenation; the dehydrogenation rate of paraffin is equal to that of mono-olefin and diolefin. Experimental selectivities obtained over a UOP dehydrogenation catalyst show good agreement with the predicted values. The rate of light paraffin conversion (Eq. (1)) over a Pt catalyst (Oleflex type process) can be expressed as a modified first order equation according to a Langmuir­Hinshelwood mechanism. The rest of the equations may be derived accordingly.


M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419

Snamprogetti's dehydrogenation process consists of a fluidized bed reactor and regeneration system. Here too the coke build-up is very low and the "regeneration" loop is actually a means of supplying heat to the reactor.

7. Heat of reaction The heat of reaction for paraffin dehydrogenation is about 30 kcal/mol (125 kJ/mol). In a cyclic adiabatic operation (e.g. Houdry), heat is provided by reheating the catalyst to a high temperature during the regeneration step, so that the catalyst cools down and conversion decreases during the reaction step; because several reactors are used in parallel, an average conversion is obtained. In an isothermal process (e.g. STAR), the catalyst is loaded inside vertical tubes inside a furnace and the heat is introduced through the tube walls. In a fluidized reactor, the temperature profile can be maintained uniformly in the backmixed zone of the bed, while heat is provided by introducing hot regenerated catalyst. In Oleflex adiabatic reactors, a significant temperature drop occurs across the catalyst bed which lowers the equilibrium conversion level; a multistage reactor system with interstage reheating is used for higher paraffin conversions. Fig. 9 illustrates conversion, equilibrium conversion, and temperature along the catalyst bed in a

Fig. 8. Simulation of selectivity for dehydrogenation of n-heptane.

6. Catalyst stability and regeneration The dehydrogenation of long-chain paraffins is performed under relatively mild temperature conditions of 400­500 C. Thus, the catalyst can maintain a long life even at high space velocity, and high catalyst productivity. Therefore, it is not economical to build facilities for catalyst regeneration. Because of equilibrium limitations, the dehydrogenation of light paraffins requires significantly higher temperatures above 600 C to achieve economically attractive conversions. The catalyst deactivation is accelerated under high-temperature conditions, and frequent catalyst regeneration is necessary for light paraffin dehydrogenation. For the dehydrogenation of light paraffins, a number of different types of reactorregeneration systems are commercially utilized. Houdry's Catofin and similar processes employ a cyclic sequence of steps--process, purge, air regeneration, purge, hydrogen reduction, and back to process. The Phillips STAR process also regenerates the catalyst on a cyclic basis, but while the Houdry regeneration is actually a mechanism to provide the heat for the reaction even when coke build-up is still very low, the catalyst in the isothermal STAR process is only regenerated after coke has accumulated to appreciable levels that result in low catalyst activity. UOP's Oleflex process uses multi-stage adiabatic reactors with CCR.

Fig. 9. Temperature profile and conversions of three-stage isobutane dehydrogenation process.

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(Fig. 10). Thus, thermal cracking and catalyst deactivation, which are accelerated at higher temperatures, can be controlled to low levels.

8. Process flow and reactor characteristics 8.1. Cyclical processes As described earlier, the Houdry Catadiene process, the Houdry Catofin process, and other similar cyclical processes make use of parallel reactors that contain a shallow bed of chromia-alumina catalyst. Fig. 11 illustrates a schematic of such a process. This technology has been used extensively for the production of butadiene and, in more recent years, for the production of isobutylene and propylene [16,17]. The feed is preheated through a fired heater before being passed over the catalyst in the reactors. The hot reactor effluent is cooled, compressed, and sent to the product fractionation and recovery section. The dehydrogenation reactors are refractory-lined carbon steel vessels

Fig. 10. Isobutane dehydrogenation.

three-stage adiabatic reactor system for the dehydrogenation of isobutane. For propane dehydrogenation, a four-stage reactor system becomes more economical because higher average temperatures are needed. A multi-stage reactor system also affords lower inlet temperatures, relative to a single stage reactor system

Fig. 11. Catofin process flow diagram.


M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419

Fig. 12. Typical timing cycle for a five reactor system.

(i.e. cold wall design). In order to accommodate continuous flow of the main streams (hydrocarbons and regeneration air), the reactors are operated on a timing cycle that satisfies the following requirement [1]: on-stream time + regeneration time + purge time = total cycle time. The number of reactors in each cycle is the prorated time fraction of the total cycle time. Thus, with five reactors, two reactors can be on stream simultaneously, two on regeneration, and one on purge, evacuation, and valve changes. Fig. 12 provides a typical timing cycle for a five-reactor unit [1], but as many as eight reactors in parallel have been provided in some units. The total cycle time is usually in the range of 15­30 min. The on-stream period at sub-atmospheric pressure is followed by a purge. Next comes regeneration at essentially atmospheric pressure, followed by purge, hydrogen reduction, and evacuation to reaction pressure, after which the reactor is ready for another on-stream period. Process streams enter and leave the reactors through fast-acting gate valves. The gate valves can

range up to 40 in. (1 m) in diameter, are designed for high-temperature service, and are equipped with a pressurized inert seal in the bonnet to prevent leakage of air into the process gas when the valve is closed. Overall, this mechanical design has proven to be very reliable over many years of operation. The regeneration is done with air that has been preheated through a direct fired burner or, alternatively, with the exhaust of a gas turbine. The regeneration step is intended to preheat the catalyst to the on-stream temperature necessary to initiate the next process cycle and to remove coke deposits on the catalyst. Flue gas sensible heat may be recovered in a waste heat boiler. The hydrogenation step prepares the catalyst for the dehydrogenation phase and also contributes additional heat from the reduction of Cr6+ ­Cr3+ . Another cyclical process is the Phillips STAR (steam active reforming) process [18]. This process uses a fixed-bed fired-tube reactor operating at a positive superatmospheric pressure. In many respects, it is similar in design to a steam reforming furnace with the heat of reaction provided by firing outside the tubes, thus operating at near isothermal conditions.

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M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419

Steam is used as a diluent to lower the partial pressure of the reactants and, thus, to achieve reasonable conversion levels of about 30­40% for propane and 45­55% for butanes. Steam also helps slow down the deposition of carbon (coke) on the catalyst, thereby, extending cycle time from minutes to hours. Periodic catalyst regeneration or carbon burnoff is required to maintain the activity of the catalyst. Typical cycle time is reported to be at least 8 h, with 7 h of process time and 1 h of regeneration time. For continuous operation, various furnace modules can be operated such that, for example, seven operate in the process mode while one is in the regeneration mode. Fig. 13 shows a schematic diagram of a STAR process unit [18].

9. Continuous processes Snamprogetti, an Italian company, has commercialized fluidized bed dehydrogenation (FBD) for the catalytic dehydrogenation of light paraffins using a chromia-alumina catalyst with an alkaline promoter [6­8], which is used primarily for the dehydrogenation of isobutane to isobutylene during the manufacture of MTBE. The catalyst is microspheroidal with

an average diameter <100 m and an apparent bulk density <2000 kg/m3 [19]. The heat of reaction is provided by circulating hot regenerated catalyst back to the reactor. In all concepts, the FBD process is very similar to the FCC process units commonly used in petroleum refineries. However, because backmixing has a negative effect on the yields, horizontal baffles with suitable openings are inserted within the fluidized bed to limit the back-flow of solids, such that the fluidized bed is split into a series of stages, each comparable to a CSTR [19]. A typical process scheme is shown in Fig. 14. Fresh feed is vaporized, mixed with the recycle of unconverted paraffins, and fed to the fluidized reactor through a distributor for optimal even distribution. Entrained catalyst is removed from the product off-gas by means of cyclones. Catalyst circulates continuously from the reactor to the regenerator and vice-versa by means of transfer lines. Coke deposited on the catalyst is burnt off in the regenerator; however, because the amount of coke is relatively small, additional fuel must be burnt in the regenerator in order to satisfy the thermal requirements of the endothermic dehydrogenation reaction. However, while this approach is similar to that in the Houdry process, FBD does not have a catalyst reduction step with hydrogen before proceeding

Fig. 14. Snamprogetti's FBD process scheme.

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Fig. 15. UOP Pacol dehydrogenation process.

to the dehydrogenation cycle; lack of this step is believed to be detrimental to the overall performance of the process. UOP's catalytic dehydrogenation processes typically make use of radial flow adiabatic fixed-bed (or slowly moving bed) reactors with modified Pt-alumina catalysts. The UOP Pacol process for selective long-chain paraffin dehydrogenation to produce linear monoolefins is shown in Fig. 15 in combination with the UOP detergent alkylation process. The Pacol process consists of a radial-flow reactor and a product recovery section. Worldwide, more than 2 million MTA of linear alkyl benzene (LAB) is produced employing this process [20]. The flow diagram of the UOP Oleflex process is shown in Fig. 16. The process consists of a reactor section and a product recovery section. The reactor section consists of three or four stages of radial-flow reactors, charge and interstage heaters, reactor feed-effluent exchangers, and the CCR section (Fig. 17). As noted earlier, today more than 1 million metric tons propylene and 2 million metric tons isobutylene are produced via this route [13]. The

performances of the Pacol and Oleflex processes are summarized in Table 1. Use of the Oleflex process for the dehydrogenation of ethane to ethylene has also been investigated but, to date, the economics do not appear to be favorable because of the low equilibrium conversion and the need to operate at a pressure lower than atmospheric if a reasonable ethane conversion is to be expected. The cost of fractionating ethylene in an ethane­ethylene splitter is otherwise too high. Dow Chemical has recently been awarded a patent [21] for the dehydrogenation of ethane over a metal-mordenite catalyst complex at relatively low conversions in which the product

Table 1 Performance of Pacol and Oleflex processes Process Oleflex Feed Propane n-Butane Isobutane n-Heptane n-C10 ­C13 n-C11 ­C14 Conversion (%) 40 50 50 20 13 13 Selectivity (%) 90 85 92 90 90 90



M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419

Fig. 16. UOP Oleflex process.

Fig. 17. Oleflex regeneration section.

M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397­419 Table 2 Characteristics of five reactor systemsa Downflow Low pressure drop Plug flow Catalyst addition or removal High heat transfer, near isothermal Variable heat transfer coefficient ­ Radial flow ­ ­ ­ Tubular ­ ­ ­ ­


Fluidized bed ­

a The choice of the right reactor depends on the catalyst and the selection of operating conditions. The `dash' represents beneficial characteristics of each reactor type.

ethylene is selectively recovered from the dilute ethylene­ethane stream by alkylating it with benzene.

10. Reactor design options The choice of reactor design plays a very important role in the success of catalytic processes. The following types of reactor design are commercial today for endothermic catalytic dehydrogenation processes: · · · · downflow adiabatic fixed-bed; radial flow fixed-bed or moving bed adiabatic; tubular isothermal and fluidized bed

Table 2 summarizes the main characteristics of the four reactor systems.

11. Other dehydrogenation technologies The processes discussed above are for the direct catalytic dehydrogenation of paraffins to the corresponding olefins or of olefins to diolefins. Other approaches have also been considered, although none has reached the level of commercialization. Some of the most notable are · halogen-assisted dehydrogenation and · oxydehydrogenation. Use of halogens for the dehydrogenation of paraffins has been proposed in different ways. For example, heavy paraffins were first chlorinated and then dehydrochlorinated to heavy olefins commercially in the past both by Shell and by H¨ ls, among others. Pyrolu ysis of methane in the presence of chlorine has been

proposed by Benson [22] for the production of acetylene and ethylene. Other chlorination/dehydrochlorination cycles have been proposed for the production of ethylene from ethane. Propane dehydrogenation in the presence of iodine via a propyl iodide intermediate has also been proposed [4,23]. Apart from the apparent corrosion problems associated with the use of halogens, other difficulties readily come to mind owing to the relatively high cost of chlorine, and even more so of iodine, and the need to either dispose of or recycle vast quantities of halogens. Oxydehydrogenation or oxidative dehydrogenation can be considered in at least two different ways. Use of oxygen to oxidize the hydrogen coproduct from dehydrogenation, and thus to displace the dehydrogenation equilibrium to higher conversions. This approach has been used commercially in the catalytic dehydrogenation of ethylbenzene to styrene as in the UOP Styro-PlusTM process or in the ABB Lummus/UOP SMARTTM process, but to date has not succeeded in the dehydrogenation of light or heavy paraffins. This technology has been used in a styrene unit at Mitsubishi Chemicals, Kashima, Japan. Although a similar approach has been proposed for the dehydrogenation of paraffins [24­27], it has not been commercialised. Direct use of oxygen as a means of dehydrogenating, for example, ethane to ethylene. Oxydehydrogenation has successful commercial applications in the conversion of n-butenes to butadiene (e.g. as in the Oxo-D process referred to earlier), but not yet for the production of ethylene or propylene. This subject is analyzed in more detail in the following section. Use of oxydehydrogenation relative to straight catalytic dehydrogenation must be viewed both in terms of safety issues and in an economic context. On the


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latter, even though oxydehydrogenation offers advantages as a means of overcoming thermodynamic equilibrium limitations, it also leads to the total or partial loss of byproduct hydrogen, which in some instances can have a very significant economic impact. Although less apparent, oxydehydrogenation also plays a role in the work done by BP Amoco, Asahi, and others to extend ammoxidation to the direct conversion of propane to acrylonitrile. It is believed that the ammoxidation of propane proceeds through a transient propylene intermediate from which acrylonitrile is derived through a conventional ammoxidation pathway [28]. 11.1. Oxydehydrogenation of ethane and propane Oxydehydrogenation of ethane and propane as a route to ethylene and propylene, respectively, has the attractive feature of removing the equilibrium conversion restriction of dehydrogenation. Table 3 contrasts the calculated equilibrium conversions of oxydehydrogenation of ethane and propane with the calculated equilibrium conversions for dehydrogenation. The formation of water rather than hydrogen in oxydehydrogenation effectively removes the equilibrium constraint on conversion at all temperatures of interest. At the same time, oxydehydrogenation is not without its own set of challenges that, in the case of ethane and propane, have kept oxydehydrogenation from being practiced on a commercial scale. Challenges in oxydehydrogenation include handling mixtures of paraffins and oxygen, which can be explosive at certain compositions, suitable conversion of paraffins, which often is limited by maintaining a safe paraffin­oxygen composition, and suitable selectivity to olefins. The carbon oxides--carbon monoxide and carbon dioxide--are thermodynamically more stable

than the olefins, and so catalysts must be found that can stop the reaction at the olefin rather than allowing it to proceed on to the oxides. Also, to compete with steam cracking, the selectivity to olefins must be quite high. Selectivity to ethylene from ethane in steam cracking is reported to be about 84% at 54% ethane conversion (800 C, 0.3 kg steam per kg feed, 0.79 s residence time, and 154 kPa hydrocarbon partial pressure), and 78% at 69% ethane conversion (833 C, 0.3 kg steam per kg feed, 0.75 s residence time, and 154 kPa hydrocarbon partial pressure) [29]. 11.1.1. Ethane oxydehydrogenation Ethane oxydehydrogenation as a route to ethylene has been examined at temperatures in the range of 300­500 C with reducible metal oxide catalysts, and at higher temperatures, above about 600 C, with largely non-reducible and reducible metal oxide catalysts. The latter have evolved primarily out of investigations of methane oxidative coupling [80­82]. In the oxidative coupling of methane, ethane and ethylene are major products, and ethylene has been shown to be derived from ethane [80­82]. The dehydrogenation or oxydehydrogenation of ethane to ethylene, must occur over these catalysts, and since the overall selectivity is high, the selectivity to ethylene must be good. In addition to these two lines of research, recent work at high temperatures using Pt and Pt-Sn on a monolith support has been reported.

12. Lower temperature ethane oxydehydrogenation Ethane oxydehydrogenation at temperatures in the range of 300­400 C is conducted with reducible metal oxide catalysts usually containing vanadium.

Table 3 Percentage conversions at equilibrium for dehydrogenation and oxydehydrogenation of ethane and propane at atmospheric pressurea Temperature (K) Ethane Dehydrogenation 400 600 800 1000


Propane Oxydehadrogenation 100 100 100 100 Dehydrogenation <1 1 25 87 Oxydehydrogenation 100 100 100 100

<1 <1 7 51

Stoichiometric mixture of alkane and oxygen for oxydehydrogenation.

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Early work in this area was conducted by Thorsteinson et al. using a vanadium molybdenum niobium oxide catalyst [30,31], and a number of papers and patents based on that work have been issued by Union Carbide Corp. and other laboratories [32­39]. Ethane is thought to react with molybdenum or vanadium in the catalyst to form surface ethoxide, which can then undergo a beta-elimination process to form ethylene. The surface ethoxide can be oxidized further to make surface acetate, which leads to acetic acid on hydrolysis. Acetic acid in varying amounts is coproduced with ethylene at pressures greater than one atmosphere. Later work has shown that catalysts can be made selective for either ethylene or acetic acid by modifications in the elemental makeup of the catalyst and suitable adjustments of reaction conditions [38,39]. Selectivity to ethylene with these catalysts is in the range of approximately 70% at approximately 70% conversion of ethane. Ethane oxydehydrogenation at temperatures in the range of 500 C may be conducted with phosphorous/molybdenum/antimony oxide catalysts [40] or iron-containing solid solution catalysts stabilized with one or more metal oxides [41]. Phosphorous/ molybdenum/antimony oxide catalysts, suitably modified with other elements give a selectivity to ethylene in the range of 78% at 20% conversion of ethane. Ironcontaining solid solution catalysts stabilized with metal oxide are optimally used with a flow of hydrogen chloride and water in addition to ethane and water. At a contact time of 12 s at 550 C over an iron in -alumina solid solution catalyst stabilized with lanthanum oxide, ethane conversion is in the range of 90%. Products are ethylene (80­93% selectivity) and vinyl chloride (1­4% selectivity). As a route to ethyleneonly, this route seems unattractive because of the production of vinyl chloride, which must be separated from ethylene. For vinyl chloride producers, however, where the stream may be able to be used without separation, such a high selectivity/conversion process to ethylene may possibly be economically attractive.

13. Higher temperature ethane oxydehydrogenation Ethane oxydehydrogenation at temperatures in the range of 650­800 C is conducted with a variety

of materials, most not containing the elements molybdenum and vanadium, which are common in lower-temperature catalysts. As opposed to the lower-temperature catalysts which tend to produce coproducts, the higher temperature processes tend to produce only ethylene as C2 product, though some do produce methane from cracking reactions. Oxygenated products other than carbon oxides largely are absent, likely because of their instability at high temperature, especially in contact with the catalyst and oxygen. Li and coworkers at the Lanzhou Institute of Chemical Physics, China, have shown that a Na2 WO4 -Mn/SiO2 catalyst (a high-selectivity methane coupling catalyst) at 700 C is capable of giving greater than 70% selectivity to ethylene (and 10% selectivity to methane) at >70% conversion of ethane [42]. Wang and coworkers at the National Institute of Materials and Chemical Research, Japan, found that lithium chloride on sulfated zirconia at 650 C yielded 70% selectivity to ethylene (and 2% to methane) at 98% conversion of ethane [43,44]. Although the catalyst did show some deactivation with time, it can still maintain an ethylene yield as high as 50% after 24 h. Lin et al. at the Tokyo Institute of Technology, Japan, and Chonnam National University, Korea, found that a SrBi3 O4 Cl3 catalyst at 640 C is capable of giving 90% selectivity to ethylene at a 25% conversion of ethane [45]. Co-feeding HCl did not change the conversion or the selectivity, but did slow the activity decrease along with the loss of chlorine observed in its absence. Dang et al. at the Lanzhou Institute of Chemical Physics, China, have examined a vanadium oxide on barium carbonate catalyst at 650 C that gave good activity [46]. At 34% conversion of ethane, they obtained 76% selectivity for ethylene. Au and coworkers at Hong Kong Baptist University examined SrFeO3- Cl and Ho2 O3 with BaCl2 [47,48]. Of the two, SrFeO3- Cl gave the higher yield of ethylene. At 680 C, SrFeO3- Cl yielded 90% ethane conversion and 70% selectivity. Over 40 h, there was no drop-off in yield or in chlorine content of the catalyst. Longer-duration tests were not conducted. The Ho2 O3 with BaCl2 catalyst at 640 C yielded 57% conversion and 68% selectivity to ethylene. Choudhary et al. at the National Chemical Laboratory, India, have examined strontium and other rare earth oxides deposited on sintered low surface area


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supports precoated with lanthanum and other rare earth oxides [49,50]. A Sr-Nd2 O3 catalyst showed the highest activity and selectivity for ethane. At 800 C, the Sr-Nd2 O3 catalyst was capable of giving 60% conversion of ethane and greater than 80% selectivity to ethylene. Workers at Phillips Petroleum Company have examined catalysts comprising lithium, titanium, and manganese, and catalysts comprising cobalt, phosphorous, and at least one promoter selected from a list of elements [51,52]. At 650 C, the lithium/titanium catalyst with manganese gave an ethane conversion of 47% and a selectivity to ethylene of 75%. At 670 C, a catalyst containing cobalt, phosphorous potassium, and zirconium where the cobalt was introduced as cobalt sulfide, and calcining of the catalyst was accomplished in the absence of oxygen at 670 C, gave an 85% conversion of ethane and a 86% selectivity to ethylene (25 vol.% ethane/75 vol.% air at 3 psig and 2400 GHSV). Addition of a halogen-containing compound, such as methyl chloride, increases ethylene yield. There is an indication that the catalyst loses activity with use over time for oxydehydrogenation of hydrocarbons. If catalyst activity and selectivity can be maintained over long periods of time, this catalyst seems like a good potential candidate for use in an economical process for ethane oxydehydrogenation to ethylene.

ethane of 67% [56,57]. The authors' consideration on mechanism follows: "Qualitatively, this process must involve first the oxidation of H2 to H2 O, which generates heat and removes O2 , followed by dehydrogenation of C2 H6 to produce C2 H4 and H2 , and all of these reactions occur within 10-3 s. Possible mechanisms to explain the results are (i) purely catalytic reactions on the Pt­Sn surface, (ii) purely homogeneous reaction and (iii) catalytic H2 oxidation followed by homogeneous ethane decomposition. We will consider each of these mechanisms and show that, while each of these gives partial interpretation of results, none appears to be totally satisfactory" [56]. Hydrogen is produced in the process in greater amount than fed, so that hydrogen required for the process can be recycled from downstream equipment. Such high selectivity to ethylene coupled with low contact time and high conversion of ethane makes this process an attractive possibility as an alternative to steam cracking of hydrocarbons as a commercial route to ethylene. While the possibility is bright, a number of questions need to be resolved. Among these are the questions of long-term catalyst activity, long-term selectivity, and whether such mixtures of ethane, hydrogen, and oxygen in the ratios needed, reported to be 2:2:1 at 950 C, can be handled safely on a commercial scale. In addition, issues related to process start-ups, shutdowns, and re-starts have to be addressed.

14. Ethane oxydehydrogenation over monolith catalysts Ethane oxydehydrogenation over Pt and Pt-Sn coated monolith catalysts at high temperatures and extremely short contact times has been the subject of work by Schmidt and coworkers at the University of Minnesota for the past several years [53­58]. Work along similar lines has also been reported to be under way in Russia. Similar results are reported over platinum gauze and platinum coated pellets [59­61]. Conversions of ethane on the order of 70% and selectivities on the order of 65% are obtained over a Pt-coated ceramic foam monolith at approximately 1000 C at contact times on the order of 1 ms. With the addition of tin to the catalyst and with the addition of hydrogen to the feed, selectivity to ethylene could be increased to above 85% at a conversion of

15. Ethane oxydehydrogenation with carbon dioxide Another approach that is being pursued in several laboratories is to use CO2 as a mild oxidant for oxydehydrogenation of hydrocarbons. [42,62­68]. In addition to ethane, oxidative dehydrogenation of ethylbenzene [67] and of propane has also been demonstrated [65,66,68]. Extensive work has been done on oxidative dehydrogenation of ethane with CO2 by Xu et al. [62] of the Dalian Institute of Chemical Physics. Thus, reaction temperature was significantly lowered from that used in steam cracking, though it is still quite high. High selectivity was obtained at 800 C, 1000 h-1 GHSV and 0.1 MPa (1 atm) using several catalysts containing oxides of Cr, Cr­Mn, Cr­Mn­Ni, Cr­Mn­La onto a Silicalite-2 (Si-2) zeolite. Highest conversion/selectivity observed were

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69.8%/90.8% at 800 C and 800 h-1 and 72.2%/90.4% at 820 C/1000 h-1 over a Cr/Si-2 catalyst. These conversions approached equilibrium for these two conditions; which were calculated to be 72.8 and 77.6%. Oxydehydrogenation using CO2 according to the authors [62] has the following special features. · CO2 acts as a mild oxidant for the oxidative dehydrogenation of C2 H6 to yield C2 H4 . · A significant lowering of the reaction temperature as compared to the steam cracking process, resulting in lower coke formation on the catalyst. · C2 H6 conversion and C2 H4 selectivity are higher in this process than in all other processes of C2 H4 production from C2 H6 . Furthermore, no C3 + and C2 H2 products are formed. The dehydrogenation of ethane is facilitated by the reaction of CO2 with H2 (reverse water-gas shift) to make CO + H2 O and also with ethane and methane as shown as follows: C2 H6 = C2 H4 + H2 CO2 + H2 = CO + H2 O C2 H6 + 2CO2 = 4CO + 3H2 CH4 + CO2 = 2CO + 2H2 In addition, ethane hydrogenolysis to methane also provides favorable free energy of reaction. C2 H6 + H2 = 2CH4 The authors suggest that ethane oxydehydrogenation with carbon dioxide takes place according to the following overall stoichiometry: 16C2 H6 + 9CO2 = 14C2 H4 + 12CO + 6H2 O + 12H2 + CH4 It can be seen from this reaction scheme that in order to increase the selectivity to ethylene, the formation of methane must be suppressed and that the key problem is to develop a catalyst that can suppress the thermodynamically favorable side reactions. The work of Xu et al. [62] also demonstrated that the inclusion of steam in the feed leads to a three-fold reduction in coke formation at 800 C. These authors also studied the C2 H6 + CO2 reaction in FCC tail-gas to increase the ethylene content.

Thus, using a tail-gas containing 18.8% C2 H6 and 19.2% C2 H4 , the ethylene content was increased to 25.4­27.2% depending on the composition of the catalyst. This reaction was carried out at 1073 K, 0.1 MPa, 1000 h-1 GHSV and CO2 :C2 H6 mole ratio of 1:1. The Cr/Si-2 catalyst gave the lowest ethylene enhancement (25.4%) while Cr-Mn-La/Si-2 catalyst gave the highest ethylene enhancement to 27.2%. The corresponding ethane conversions/selectivity are 60.6%/79.6% and 63.6%/85.8%, respectively. Some of the other early work in this area was done using highly selective catalysts for the oxidative coupling of methane by Liu et al. of the Lanzhou Institute of Chemical Physics [42]. Their work, employing a selective methane coupling catalyst (Na2 WO4 -Mn/SiO2 ), showed that >70% conversion and selectivity could be achieved at 700­750 C and space velocities of >30,000 h-1 , employing O2 as the oxidant. This is in contrast with 53% conversion/97% selectivity at 800 C and 69.5% conversion/90.5% selectivity at 850 C employing C2 H6 :CO2 = 1:1 and 3600 h-1 space velocity. These catalysts were stable for 100 h of operation. The authors propose that surface lattice oxygen is responsible for selective oxydehydrogenation while the bulk lattice oxygen is responsible for deep (non-selective) oxidation of ethane. Nakagawa et al. [63] studied a series of metal oxides for ethane dehydrogenation in the presence of CO2 at 650 C, C2 H6 :CO2 = 5:25 ml/min and SV = 900 h-1 ml (g-cat)-1 . The order of activity was as follows: Ga2 O3 > Cr 2 O3 > V2 O5 > TiO2 > Mn3 O4 > In2 O3 > ZnO > La2 O3 . The ethylene selectivity was generally >85% in the presence of CO2 . Ethylene yields for Ga2 O3 in the presence of CO2 was approximately twice than in its absence. The ethylene yield enhancement with Cr2 O3 was only a modest 1.2%, while in the case of V2 O5 , there was a 2.7% detrimental effect. The authors propose involvement of acid sites in the dehydrogenation reaction. Wang et al. [64] of Japan studied the effect of sulfate and Na-treatment silica on the dehydrogenation of ethane with CO2 at 650 C and 1 atm. Sulfating silica (348 m2 /g) with (NH4 )2 SO4 (and calcining at 700 C/3 h to give 2 wt.% SO4 2- ), and adding


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5 wt.% Cr gave the best catalytic performance. A conversion/selectivity/yield of 53.9/87.6/55.2% was observed at 650 C/1 atm. With a 10:50:40 = ethane:CO2 :N2 mixture. The corresponding 5% Cr catalyst (3.34 m2 /g on silica gave 56.1/92.9/52.1% conversion/selectivity/yield. However, Na-Cr/SiO2 gave very poor performance of 4.9/98.2/4.8%. Furthermore, the surface area dropped to 3.1 m2 /g. In comparison, silica alone of 348 m2 /g gave 2.2/97.2/0.2% conversion/selectivity/yield. Dehydrogenation is proposed to occur by abstraction of H by oxygen species which in turn are formed from surface carbonates decomposition. At low temperatures, decomposition of carbonates absorbed on strong basic sites restricts formation and mobility of oxygen resulting in lower activity of Cr/Na-SiO2 catalyst. However, sulfation of silica favors reaction with hydrocarbon and hence, higher activity (along with formation of CO + H2 ). The Cr/SiO2 and Cr/SO4 2- -SiO2 were shown to be stable for 5­6 h of operation. Macho and coworkers [65,66] studied oxydehydrogenation of propane and butane over Mn-Cr-K/Y-Al2 O3 to give propene and butenes, respectively, according to the following reaction: C3 H8 + CO2 C3 H6 + CO + H2 O C4 H10 + CO2 C4 H8 + CO + H2 O Splitting of the C­C bond proceeds simultaneously leading to the formation of both higher and lower hydrocarbons as shown in the following equation. These reactions are preferred at relatively higher temperatures while the relatively lower temperatures prefer straight dehydrogenation with CO2 . In addition, coke formation is more prevalent at relatively higher temperatures. 2C3 H6 C4 H8 + C2 H4 4C3 H8 + 4CO2 3C4 H8 + 4CO + 4H2 O C3 H8 + 3CO2 C2 H4 + 4CO + 2H2 O C3 H8 C3 H6 + H2 C2 H4 + CO2 CH4 + 2CO The authors propose that: · at middle temperatures (850­1000 K) dehydrogenation of alkanes proceeds first and hydrogen generated reacts with CO2 giving CO + H2 O;

· at higher temperatures (>950 K), alkanes may react directly with oxygen left from CO2 . The authors also explain the reactivity of catalysts in terms of acidic and basic nature of supports and the reactivity of CO2 with such supports. Typically, good active catalysts appear to approach thermodynamic equilibrium. However, temperatures required to achieve higher conversions and yields also lead to higher levels of coking and by-product formation. No economic assessment of the CO2 -based dehydrogenation processes is available. The authors are not aware of any commercialization efforts using this reaction.

16. Propane oxydehydrogenation Propane oxydehydrogenation combining high propane conversion and high propylene selectivity has proved an elusive goal, and propane oxydehydrogenation as a route to propylene appears to be far from realizing its commercialization potential. One difficulty is that propylene is more easily oxidized than is propane, so that selectivity tends to decline rapidly with conversion. Another is that at temperatures above about 700 C, propane cracking becomes significant, and a variety of products other than propylene are produced. Efforts to increase propylene selectivity have centered in the areas of more selective catalysts at temperatures below 700 C, more selective catalysts and conditions at high temperatures, membrane reactors, and cyclic-operation reactors. A variety of catalysts for propane oxydehydrogenation have been examined recently at temperatures below 700 C [69­75]. Particularly, notable for its high selectivity to propylene is a vanadia-silica-zirconia catalyst of Rulkens and Tilley at the University of California, Berkley, CA [75]. These workers using a molecular precursor route, produced an 18/36/46 V2 O5 -SiO2 -ZrO2 catalyst which gave 81.5% selectivity to propylene at 8% conversion of propane at 550 C. The presence of zirconium is important in retaining high selectivity, likely by stabilizing the dispersion of vanadia. These catalysts also are among the most active for propane oxydehydrogenation. Ranzi and coworkers at the Polytechnic University of Milan [76] and Choudhary et al. at the National Chemical Laboratory of India [77] have examined

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oxydehydrogenation of propane at high temperatures with and without catalysts. Selectivities to olefins, comprising primarily ethylene and propylene, can be obtained in selectivities of 50­80%. Conversion of propane is on the order of 90% under these conditions. Alfonso, Julbe, Farrusseng, Men´ ndez, and Santae mar´a at the University of Zaragoza, Spain, and the i Laboratoire des Mat´ riaux et Proc´ d´ s Membranaires, e e e France, examined the oxidative dehydrogenation of propane over V-Al2 O3 catalytic membranes which allow separation of propane and oxygen feeds [78]. Selectivities of 51% to propylene were obtained at 8% conversion of propane at 550 C, which is higher than the 44% selectivity obtained at the same temperature and conversion with premixed feeds. Creaser et al. at Lulea University of Technology, Sweden, University of Waterloo, Canada, and Chalmers University of Technology, Sweden, examined cyclic operation of the oxidative dehydrogenation of propane over a V-Mg-O catalyst [79]. In a cyclic reactor where oxygen and propane were alternately passed over the catalyst, propylene selectivity was considerably increased at 510 C compared to steady state operation with mixed feeds. At a 1:1 propane:oxygen ratio, selectivity to propylene was 78% at 4.3% conversion of propane with cyclic operation compared to 55% selectivity to propylene at 5.5% conversion of propane with mixed feeds. These conversions/selectivities are still far from being commercially attractive.

intermediate. This certainly appears to be the case with the majority, if not all, of the catalysts that are active for oxydehydrogenation that operate at 500­800 C. However, under these conditions, it is very difficult to avoid undesirable partial/total oxidation to CO and CO2 via gas phase and surface reactions. Therefore, the challenge for oxydehydrogenation catalysis is to develop highly active and selective catalysts for totally selective oxidation of alkanes to alkenes. References

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