Dehydrogenation and oxydehydrogenation of paraffins to olefins
M.M. Bhasina,∗, J.H. McCaina, B.V. Vorab, T. Imaib, P.R. Pujadob
a Union Carbide & UOP LLC, Subsidiary of Dow Chemical Co., P.O. Box 8361, 25303 South Charleston, WV, USAb 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, whilecatalytic paraffin oxydehydrogenation for olefin production has not yet been commercialized. However, there are someinteresting 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 productionof butenes that were then dimerized to octenes and hydrogenated to octanes to yield high-octane aviation fuel. Dehydrogenationemploys chromia-alumina catalysts and, more recently, platinum or modified platinum catalysts. Important aspects in dehydro-genation 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 of30–60%, while the catalytic dehydrogenation of C10–C14 paraffins typically operates at conversion levels of 10–20%. In theyear 2000, nearly 7 million metric tons of C3–C4 olefins and 2 million metric tons of C10–C14 range olefins were producedvia 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 steamcracking. However, there are a number of issues related to safety and process upsets that need to be addressed. Importantobjectives in oxydehydrogenation are attaining high selectivity to olefins with high conversion of paraffin and minimizingpotentially dangerous mixtures of paraffin and oxidant. More recently, the use of carbon dioxide as an oxidant for ethaneconversion to ethylene has been investigated as a potential way to reduce the negative impact of dangerous oxidant–paraffinmixtures and to achieve higher selectivity.
World War II, catalytic dehydrogenation of butanesover a chromia-alumina catalyst was practiced forthe production of butenes, which were then dimer-ized to octenes and hydrogenated to octanes to yieldhigh-octane aviation fuel.
Dehydrogenation of butanes over a chromia-aluminacatalyst was first developed and commercialized atLeuna in Germany and was also independently de-veloped by UOP (then Universal Oil Products) in the
398 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
United States, together with ICI in England. The firstUOP-designed plant came on stream in Billingham,England, in 1940 and was soon followed by two otherunits in Heysham, England, in 1941 [1]. The primarypurpose of this butane dehydrogenation was to pro-duce butenes, which were then dimerized to octenesusing solid phosphoric acid catalysts discovered bySchaad and Ipatieff [2].
Other companies soon followed these pioneeringefforts. For example, Phillips Petroleum built a multi-tubular dehydrogenation reactor near Borger, TX, in1943 [1]. However, the most significant developmentwas made by Houdry using dehydrogenation at lessthan atmospheric pressure for higher per-pass conver-sions. This process, which came on stream toward theend of World War II, was also used for the productionof butenes. After the war, Houdry further developedand commercialized the chromia-alumina dehydro-genation system and extended it to the production ofbutadiene in what became known as the CatadieneTM
process [3]. Other companies, including Shell, Gulf,and Dow, also practiced similar dehydrogenationtechnologies.
In the dehydrogenation process using chromia-alumina catalysts, the catalyst is contained in a fixedshallow bed located inside a reactor that may be eithera sphere, a squat vertical cylinder, or a horizontal cylin-der. The actual design reflects a compromise betweengas flow distribution across a large cross-sectionalarea and the need to maintain a low pressure drop. Asignificant amount of coke is deposited on the catalystduring the dehydrogenation step, therefore, a numberof reactors are used in parallel—some for dehydro-genation while the rest are being purged or regen-erated. The dehydrogenation reactions are stronglyendothermic, and the heat is provided, at least inpart, by the sensible heat stored in the catalyst bedduring regeneration (carbon burn); additional heat isprovided by direct fuel combustion and also by heatreleased in the chromium redox cycle. The length ofthe total reactor cycle is limited by the amount of heatavailable, and can be as short as 10–20 min.
The Houdry Catadiene process was used exten-sively for the production of butadiene, either byitself (n-butane to butadiene) or in conjunction withcatalytic oxydehydrogenation ofn-butene to butadi-ene. The latter was commercialized by the Petro-TexChemical Corp. [3] and was called the Oxo-DTM
process. A similar oxydehydrogenation approach forthe production of butadiene was also practiced byPhillips Petroleum [3].
Large quantities of butadiene have become availableover the past 30 years, mostly as a by-product from thethermal cracking of naphtha and other heavy hydrocar-bons. This market shift has resulted in the shutdownof all on-purpose catalytic dehydrogenation units forbutadiene production in North America, western Eu-rope, and the far East.
In the late 1980s, the application of chromia-aluminacatalysts was extended by Houdry to the dehydrogena-tion of propane to propylene and isobutane to isobuty-lene. The new process application called CatofinTM
[4,5] operates on the same cyclic principle as in theformer Catadiene process. As of late 2000, a total ofeight Catofin units exist for the production of isobuty-lene (including two converted older Catadiene units)with an aggregate capacity of about 2.8 million metrictons per annum (MTA) isobutylene. In addition, twoCatofin units were built for the production of propy-lene, but it is understood that only one is operationalwith a nameplate dehydrogenation capacity of about250,000 MTA propylene, but usually operating onlyon a seasonal basis. Plans for another 450,000 MTACatofin propane dehydrogenation unit in Saudi Ara-bia have also been announced. The Catofin processtechnology is currently owned by Sud-Chemie and isoffered for license by ABB Lummus.
In 1959, an alternative chromia-alumina catalyticdehydrogenation process was developed in the for-mer Soviet Union. This process avoided the use of thecyclic operation by using a fluidized bed reactor con-figuration similar to the fluidized catalytic cracking(FCC) process used in refineries [6]. However, back-mixing common to dense fluidized bed operations re-sults in poor selectivity and increases the formationof heavies, sometimes called “green oils”. Circulatingregenerated catalyst is used to provide the heat of re-action in the riser and spent catalyst is reheated bycarbon burn in the regenerator. During the 1990s, alarge scale fluid bed isobutane dehydrogenation unitfor about 450,000 MTA isobutylene was commercial-ized by Snamprogetti in Saudi Arabia based on tech-nology from Yarsintez in Russia [6], but it is under-stood that this unit has only operated at lower than de-sign capacity. Recent literature articles report furtherimprovements by Snamprogetti [7,8].
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 399
2. Noble metal dehydrogenation catalysts
A different approach to catalytic dehydrogenationwas first introduced in the mid-1960s for the sup-ply of long-chain linear olefins for the production ofbiodegradable detergents.
Synthetic detergents, based on the use of branchedalkylbenzene sulfonates derived from propylenetetramer and benzene, had been introduced in the1940s. By the early 1960s, however, it became appar-ent that branched dodecylbenzene-based detergents,though very active and offering excellent perfor-mance characteristics, did not biodegrade readily andwere accumulating in the environment. The need forbiodegradable detergents prompted the developmentof catalytic dehydrogenation of long-chain linearparaffins to linear olefins.
The work on catalytic reforming with noble metal(Pt) catalysts done in the 1940s by Haensel clearlydemonstrated that Pt-based catalysts had high activityfor the dehydrogenation of paraffins to the correspond-ing olefins [9]. In the 1960s, Bloch [10] further ex-tended this thinking by developing Pt-based catalyststhat could selectively dehydrogenate long-chain linearparaffins to the corresponding internal mono-olefinswith high activity and stability and with minimumcracking. This was the basis for the UOP PacolTM pro-cess for the production of linear olefins for the man-ufacture of biodegradable detergents [11]. In 1999,there were more than 30 commercial Pt-catalyzed de-hydrogenation units in operation for the manufactureof detergent alkylate.
Long-chain paraffins are both valuable and highlyprone to cracking. Therefore, in order to maintain highselectivity and yield, it is necessary to operate at rela-tively mild conditions, typically below 500◦C, and atrelatively low per-pass conversions. While this is eco-nomical for the production of heavy linear olefins, itis not for the production of light olefins.
Paraffin dehydrogenation is an endothermic reactionthat is limited by chemical equilibrium and, accord-ing to Le Chatelier’s principle, higher conversion willrequire either higher temperatures or lower pressures.In a somewhat abbreviated form for the production ofmono-olefins, this can be expressed as follows:
x2e = Kp
Kp + P
Fig. 1. Propane dehydrogenation equilibrium at 1.00 atm abs. pres-sure.
wherexe is the equilibrium conversion,P the total ab-solute pressure andKp is the equilibrium constant forthe dehydrogenation reaction. The equilibrium con-stant can be easily calculated from Gibbs free energiesas tabulated in the API 44 report or in similar sourcesof thermodynamic data. Figs. 1 and 2 illustrate theequilibrium conversion levels that can be obtained forpropane at 1 and 0.23 atm abs. (175 Torr), respectively.
Fig. 2. Propane dehydrogenation equilibrium at 0.23 atm abs. pres-sure.
400 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
The equilibrium constant for paraffin dehydrogena-tion increases significantly as the carbon number in-creases. Fig. 3 shows the equilibrium constant for thedehydrogenation ofn-paraffins ranging from ethaneto pentadecane [12]. Fig. 4 shows the temperaturesrequired to achieve 10–40% equilibrium conversionbased on these equilibrium constants. Fig. 4 indicatesthat the temperature required for the dehydrogenationof light paraffins is much higher than for heavy paraf-fins. For 40% conversion, for example, the dehydro-genation of propane requires a temperature of at leastabout 580◦C, while dodecane can be theoreticallydehydrogenated to the same extent at only 450◦C.The equilibrium conversion increases at higher
Fig. 4. Temperatures required to achieve 10 and 40% conversionof C2–C15 n-paraffins at 1 atm.
temperatures, but side reactions, coke formation, andcatalyst deactivation are also accelerated. Thus, ex-trapolation directly from heavy olefins to light olefinscannot be done without taking other factors intoconsideration.
Production of light olefins by the catalytic dehydro-genation of light paraffins must be able to maintainreasonable per-pass conversion levels and high olefinselectivity. Very importantly, it must be able to pro-duce olefins in high yields over long periods of timewithout shutdowns.
In the early 1970s, UOP introduced continuous cata-lyst regeneration (CCR) technology that enabled noblemetal catalysts to remain at their most desirable stableactivity for several years without having to shut downthe reactor for catalyst regeneration. The combinationof noble metal catalysts operating at high severity inconjunction with CCR technology made it possible todesign, build, and economically operate large catalyticdehydrogenation units that can produce light olefins,in particular, propylene and isobutylene, at high selec-tivities while still operating at superatmospheric pres-sures. This technology is known as the UOP OleflexTM
process. As of late 2000, there were four propane de-hydrogenation units, five isobutane dehydrogenationunits, and one combined propane/isobutane dehydro-genation unit of this type in commercial operation,with an aggregate operating capacity of 900,000 MTApolymer grade propylene and 2.3 million MTAisobutylene. In addition, another propane dehydro-genation unit for 350,000 MTA polymer grade propy-lene was under design and construction.
The world propylene production capacity, based onthe use of catalytic dehydrogenation of propane hasincreased steadily over the past 10 years [13] and is ex-pected to grow even further under the right economicconditions relative to the availability of propane; onthe other hand, environmental concerns on the use ofMTBE are expected to adversely impact the future ex-pansion of isobutane dehydrogenation applications.
Although production of ethylene via catalytic dehy-drogenation over Pt catalysts is very selective (about95%), extension of this dehydrogenation technologyto ethane has not taken place due to the need for evenmore severe operating conditions; higher temperaturesand lower pressures. Such conditions cause excessivecoking of the catalyst or require costlier operationunder vacuum.
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 401
Fig. 5. Reactions by platinum and acid sites in light paraffin dehydrogenation with unmodified catalyst.
Practically, all existing catalytic dehydrogenationcapacity based on Pt catalysts is based on the Oleflexprocess with CCR; however, there are also two smallerunits for isobutane dehydrogenation for 118,000 and13,000 MTA isobutylene, respectively, both based onthe STAR technology developed by Phillips Petroleumand derived from their earlier multitubular reactor de-sign experience. This reactor design resembles a typ-ical steam reformer that is operated until the catalystdeactivates as a result of coke deposition; banks oftubes are sequentially taken out of service for cata-lyst regeneration. The STAR technology is currentlyowned and licensed by Krupp–Uhde.
Fig. 6. Reactions by platinum and acid sites in heavy paraffin dehydrogenation with unmodified catalyst.
3. Process chemistry
The main reaction in catalytic dehydrogenation isthe formation of mono-olefins from the correspond-ing feed paraffin. Other reactions include consecutiveand side reactions. The reaction pathways involved inheavy paraffin dehydrogenation (e.g. detergent-rangeC10–C14 n-paraffins) are more complicated than thosein light paraffin dehydrogenation (e.g. propane andisobutane). The main difference in reaction pathwaysis that a significant amount of cyclic compounds canform via dehydrocyclization from heavy paraffins;this is not the case for light paraffins. Figs. 5 and 6
402 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
illustrate possible reactions that take place on Pt andacid (A) sites, respectively, in the dehydrogenation oflight and heavy paraffins when the catalyst is not se-lective, e.g. unmodified platinum catalysts supportedon alumina.
The consecutive reactions, the dehydrogenation ofmono-olefins to diolefins and triolefins, are catalyzedon the same active sites as the dehydrogenation ofparaffins to mono-olefins. The consecutive reactionsthat form triolefins, aromatics, dimers, and polymersmust be suppressed kinetically or by catalyst modifi-cations.
4. Role of catalysts and supports
The discussion in this section pertains to alumina-supported platinum catalysts. The work by Poole andcoworkers [14,15] provides an extensive review ofchromia-alumina catalysts.
The key role of dehydrogenation catalysts is toaccelerate the main reaction while controlling otherreactions. Unmodified alumina-supported platinumcatalysts are highly active but are not selective to de-hydrogenation. Various by-products, as indicated inFigs. 5 and 6, can also form. In addition, the catalystrapidly deactivates due to fouling by heavy carbona-ceous materials. Therefore, the properties of platinumand the alumina support need to be modified to sup-press the formation of by-products and to increasecatalytic stability.
Fig. 7. Paraffin dehydrogenation on modified Pt catalyst.
The reaction of olefins on platinum is faster thanthat of paraffins, because olefins interact with platinummore strongly than do paraffins. The role of platinummodifiers is to weaken the platinum–olefin interactionselectively without affecting the platinum–paraffin in-teraction. Arsenic, tin, germanium, lead, bismuth areamong metals reported as platinum activity modifiers.The consecutive dehydrogenation rate of mono-olefinsand diolefins is decreased by this modification withoutlowering the rate of paraffin dehydrogenation signifi-cantly. The modifier also improves the stability againstfouling by heavy carbonaceous materials.
Platinum is a highly active catalytic element and isnot required in large quantities to catalyze the reactionwhen it is dispersed on a high surface-area support.The high dispersion is also necessary to achieve highselectivity to dehydrogenation relative to undesirableside reactions, such as cracking.
The typical high surface area alumina supports em-ployed have acidic sites that accelerate skeletal isomer-ization, cracking, oligomerization, and polymerizationof olefinic materials, and enhance “coke” formation.Alkali or alkaline earth metals assist in the control ofthe acidity. Also,�-alumina supports that have essen-tially no acidity can be utilized; however, the challengeis to obtain high dispersion of platinum on such verylow surface area supports. Therefore, acidity must beeliminated by using suitable modifiers.
Modified catalysts possess high activity and highselectivity to mono-olefins. The major by-productsare diolefins that can be controlled kinetically. Coke
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 403
formation is also suppressed and, therefore, stabil-ity is greatly improved. Over modified catalysts,the major reaction pathways for both light andheavy paraffin dehydrogenation systems are simpler(Fig. 7).
Alumina has excellent thermal stability and me-chanical strength under processing, transport, andcatalyst regeneration conditions. However, the mostimportant reason alumina is used as support ma-terial is its superior capability to maintain a highdegree of platinum dispersion, which is essentialfor achieving high dehydrogenation activity andselectivity.
The catalytic reaction rate is limited by the intra-particle mass transfer rate. If the mass transfer rate isrelatively slow, both activity and selectivity are low-ered. As a result, the support must have a low porediffusional resistance (high effectiveness factor). Fora given pore volume, the surface area and the strengthof the support increase as the pore diameter decreases,and the pore diffusional resistance decreases as thepore diameter increases. Thus, an appropriate porestructure must be determined for the support to achieveoptimal catalytic performance.
5. Dehydrogenation catalyst evaluation
In paraffin dehydrogenation, the rate of paraffin con-version (x) and mono-olefin production (sx) are givenby Eqs. (1) and (2) respectively:
dx
dt= f1(ki, Kipj ) (1)
d(sx)
dt= f2(ki, Ki, pj ) (2)
wheres is the selectivity ton-mono-olefins,t the con-tact time,f the rate function,ki the rate constant forreaction stepi, Ki the equilibrium constant for reactionstepi, andpj is the partial pressure of thej compound.
The following relationship between selectivity andconversion can be derived from Eqs. (1) and (2):
d(sx)
dx= f2(ki, Ki, pj )
f1(ki, Ki, pj )(3)
Theki andKi are a function of temperature andpj
is a function of conversion, total pressure (P), and feed
ratio (R). Therefore, Eq. (3) can be rewritten as
d(sx)
dx= f2[ki(T ), Ki(T ), x, P, R]
f1[ki(T ), Ki(T ), x, P, R](4)
As Eq. (4) is the ratio of two functions, the rateconstants become relative values and can be expressedas ki(T)/k0(T), where k0(T) is the rate constant forthe forward reaction of paraffin dehydrogenation tomono-olefins.
Eq. (4) can be written in a functional form inF asfollows:
d(sx)
dx= F
[ki(T )
k0(T ), Ki(T ), x, P, R
](5)
Eq. (5) indicates that selectivity is a function of con-version for the catalyst used (relative rate constants)and the given reaction conditions (temperature, pres-sure, feed ratio).
Selectivity decreases as the conversion increasesbecausen-mono-olefins are consecutively convertedinto by-products. Selectivity decreases sharply asconversion approaches equilibrium because the maindehydrogenation process is limited by equilibrium,but other reactions continue to occur. Therefore,if side reactions are controlled, the selectivity isimproved as the equilibrium conversion becomeshigher by increasing the temperature and by decreas-ing the pressure and the feed ratio of hydrogen toparaffin.
The relationship between selectivity and conver-sion can be simulated according to Eq. (5), if ratefunctions, relative rate constants, and equilibriumconstants are known. Fig. 8 shows simulated selectiv-ities to n-heptene andn-heptadiene for the dehydro-genation ofn-heptane. In this simulation, the relativerate constants used are unity, which represents thatthe catalyst possesses perfect selectivity regardingconsecutive dehydrogenation; the dehydrogenationrate of paraffin is equal to that of mono-olefin anddiolefin. Experimental selectivities obtained over aUOP dehydrogenation catalyst show good agreementwith the predicted values.
The rate of light paraffin conversion (Eq. (1)) overa Pt catalyst (Oleflex type process) can be expressedas a modified first order equation according to aLangmuir–Hinshelwood mechanism. The rest of theequations may be derived accordingly.
404 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
Fig. 8. Simulation of selectivity for dehydrogenation ofn-heptane.
6. Catalyst stability and regeneration
The dehydrogenation of long-chain paraffins is per-formed under relatively mild temperature conditionsof 400–500◦C. Thus, the catalyst can maintain a longlife even at high space velocity, and high catalyst pro-ductivity. Therefore, it is not economical to build fa-cilities for catalyst regeneration.
Because of equilibrium limitations, the dehydro-genation of light paraffins requires significantly highertemperatures above 600◦C to achieve economicallyattractive conversions. The catalyst deactivation isaccelerated under high-temperature conditions, andfrequent catalyst regeneration is necessary for lightparaffin dehydrogenation. For the dehydrogenation oflight paraffins, a number of different types of reactor-regeneration systems are commercially utilized.
Houdry’s Catofin and similar processes employ acyclic sequence of steps—process, purge, air regener-ation, purge, hydrogen reduction, and back to process.
The Phillips STAR process also regenerates the cat-alyst on a cyclic basis, but while the Houdry regener-ation is actually a mechanism to provide the heat forthe reaction even when coke build-up is still very low,the catalyst in the isothermal STAR process is onlyregenerated after coke has accumulated to appreciablelevels that result in low catalyst activity.
UOP’s Oleflex process uses multi-stage adiabaticreactors with CCR.
Snamprogetti’s dehydrogenation process consistsof 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 supplyingheat to the reactor.
7. Heat of reaction
The heat of reaction for paraffin dehydrogenationis about 30 kcal/mol (125 kJ/mol). In a cyclic adi-abatic operation (e.g. Houdry), heat is provided byreheating the catalyst to a high temperature duringthe regeneration step, so that the catalyst cools downand conversion decreases during the reaction step; be-cause several reactors are used in parallel, an averageconversion is obtained. In an isothermal process (e.g.STAR), the catalyst is loaded inside vertical tubesinside a furnace and the heat is introduced throughthe tube walls. In a fluidized reactor, the temperatureprofile can be maintained uniformly in the backmixedzone of the bed, while heat is provided by intro-ducing hot regenerated catalyst. In Oleflex adiabaticreactors, a significant temperature drop occurs acrossthe catalyst bed which lowers the equilibrium conver-sion level; a multistage reactor system with interstagereheating is used for higher paraffin conversions.
Fig. 9 illustrates conversion, equilibrium conver-sion, and temperature along the catalyst bed in a
Fig. 9. Temperature profile and conversions of three-stage isobutanedehydrogenation process.
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 405
Fig. 10. Isobutane dehydrogenation.
three-stage adiabatic reactor system for the dehydro-genation of isobutane. For propane dehydrogenation,a four-stage reactor system becomes more economi-cal because higher average temperatures are needed.A multi-stage reactor system also affords lower inlettemperatures, relative to a single stage reactor system
Fig. 11. Catofin process flow diagram.
(Fig. 10). Thus, thermal cracking and catalyst deacti-vation, 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 cyclicalprocesses make use of parallel reactors that contain ashallow bed of chromia-alumina catalyst. Fig. 11 il-lustrates a schematic of such a process. This technol-ogy has been used extensively for the production ofbutadiene and, in more recent years, for the produc-tion of isobutylene and propylene [16,17]. The feed ispreheated through a fired heater before being passedover the catalyst in the reactors. The hot reactor ef-fluent is cooled, compressed, and sent to the productfractionation and recovery section. The dehydrogena-tion reactors are refractory-lined carbon steel vessels
406 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
Fig. 12. Typical timing cycle for a five reactor system.
(i.e. cold wall design). In order to accommodate con-tinuous flow of the main streams (hydrocarbons andregeneration air), the reactors are operated on a timingcycle 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 proratedtime fraction of the total cycle time. Thus, with fivereactors, two reactors can be on stream simultaneously,two on regeneration, and one on purge, evacuation, andvalve changes. Fig. 12 provides a typical timing cyclefor a five-reactor unit [1], but as many as eight reactorsin parallel have been provided in some units. The totalcycle time is usually in the range of 15–30 min.
The on-stream period at sub-atmospheric pressureis followed by a purge. Next comes regeneration at es-sentially atmospheric pressure, followed by purge, hy-drogen reduction, and evacuation to reaction pressure,after which the reactor is ready for another on-streamperiod. Process streams enter and leave the reactorsthrough fast-acting gate valves. The gate valves can
range up to 40 in. (1 m) in diameter, are designedfor high-temperature service, and are equipped with apressurized inert seal in the bonnet to prevent leakageof air into the process gas when the valve is closed.Overall, this mechanical design has proven to be veryreliable over many years of operation.
The regeneration is done with air that has been pre-heated through a direct fired burner or, alternatively,with the exhaust of a gas turbine. The regenerationstep is intended to preheat the catalyst to the on-streamtemperature necessary to initiate the next process cycleand to remove coke deposits on the catalyst. Flue gassensible heat may be recovered in a waste heat boiler.The hydrogenation step prepares the catalyst for thedehydrogenation phase and also contributes additionalheat from the reduction of Cr6+–Cr3+.
Another cyclical process is the Phillips STAR(steam active reforming) process [18]. This processuses a fixed-bed fired-tube reactor operating at a pos-itive superatmospheric pressure. In many respects, itis similar in design to a steam reforming furnace withthe heat of reaction provided by firing outside thetubes, thus operating at near isothermal conditions.
M.M
.B
hasinet
al./Applied
Catalysis
A:
General
221(2001)
397–419407
408 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
Steam is used as a diluent to lower the partial pres-sure of the reactants and, thus, to achieve reasonableconversion levels of about 30–40% for propane and45–55% for butanes. Steam also helps slow down thedeposition of carbon (coke) on the catalyst, thereby,extending cycle time from minutes to hours.
Periodic catalyst regeneration or carbon burnoff isrequired to maintain the activity of the catalyst. Typ-ical cycle time is reported to be at least 8 h, with 7 hof process time and 1 h of regeneration time. For con-tinuous operation, various furnace modules can be op-erated such that, for example, seven operate in theprocess mode while one is in the regeneration mode.Fig. 13 shows a schematic diagram of a STAR processunit [18].
9. Continuous processes
Snamprogetti, an Italian company, has commer-cialized fluidized bed dehydrogenation (FBD) for thecatalytic dehydrogenation of light paraffins using achromia-alumina catalyst with an alkaline promoter[6–8], which is used primarily for the dehydrogena-tion of isobutane to isobutylene during the manufac-ture of MTBE. The catalyst is microspheroidal with
Fig. 14. Snamprogetti’s FBD process scheme.
an average diameter<100�m and an apparent bulkdensity <2000 kg/m3 [19]. The heat of reaction isprovided by circulating hot regenerated catalyst backto the reactor. In all concepts, the FBD process isvery similar to the FCC process units commonly usedin petroleum refineries. However, because backmix-ing has a negative effect on the yields, horizontalbaffles with suitable openings are inserted within thefluidized bed to limit the back-flow of solids, suchthat the fluidized bed is split into a series of stages,each comparable to a CSTR [19]. A typical processscheme is shown in Fig. 14.
Fresh feed is vaporized, mixed with the recycle ofunconverted paraffins, and fed to the fluidized reactorthrough a distributor for optimal even distribution. En-trained catalyst is removed from the product off-gas bymeans of cyclones. Catalyst circulates continuouslyfrom the reactor to the regenerator and vice-versa bymeans of transfer lines. Coke deposited on the catalystis burnt off in the regenerator; however, because theamount of coke is relatively small, additional fuel mustbe burnt in the regenerator in order to satisfy the ther-mal requirements of the endothermic dehydrogenationreaction. However, while this approach is similar tothat in the Houdry process, FBD does not have a cat-alyst reduction step with hydrogen before proceeding
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 409
Fig. 15. UOP Pacol dehydrogenation process.
to the dehydrogenation cycle; lack of this step is be-lieved to be detrimental to the overall performance ofthe process.
UOP’s catalytic dehydrogenation processes typi-cally make use of radial flow adiabatic fixed-bed (orslowly moving bed) reactors with modified Pt-aluminacatalysts.
The UOP Pacol process for selective long-chainparaffin dehydrogenation to produce linear mono-olefins is shown in Fig. 15 in combination with theUOP detergent alkylation process. The Pacol processconsists of a radial-flow reactor and a product recov-ery section. Worldwide, more than 2 million MTA oflinear alkyl benzene (LAB) is produced employingthis process [20].
The flow diagram of the UOP Oleflex processis shown in Fig. 16. The process consists of a re-actor section and a product recovery section. Thereactor section consists of three or four stages ofradial-flow reactors, charge and interstage heaters,reactor feed-effluent exchangers, and the CCR sec-tion (Fig. 17). As noted earlier, today more than 1million metric tons propylene and 2 million metrictons isobutylene are produced via this route [13]. The
performances of the Pacol and Oleflex processes aresummarized in Table 1.
Use of the Oleflex process for the dehydrogenationof ethane to ethylene has also been investigated but, todate, the economics do not appear to be favorable be-cause of the low equilibrium conversion and the needto operate at a pressure lower than atmospheric if a rea-sonable ethane conversion is to be expected. The costof fractionating ethylene in an ethane–ethylene split-ter is otherwise too high. Dow Chemical has recentlybeen awarded a patent [21] for the dehydrogenationof ethane over a metal-mordenite catalyst complexat relatively low conversions in which the product
410 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
Fig. 16. UOP Oleflex process.
Fig. 17. Oleflex regeneration section.
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 411
Table 2Characteristics of five reactor systemsa
Downflow Radial flow Tubular Fluidized bed
Low pressure drop – –Plug flow – – –Catalyst addition or removal – –High heat transfer, near isothermal – –Variable heat transfer coefficient
a The choice of the right reactor depends on the catalyst and the selection of operating conditions. The ‘dash’ represents beneficialcharacteristics of each reactor type.
ethylene is selectively recovered from the diluteethylene–ethane stream by alkylating it with benzene.
10. Reactor design options
The choice of reactor design plays a very importantrole in the success of catalytic processes. The follow-ing types of reactor design are commercial today forendothermic 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 thefour reactor systems.
11. Other dehydrogenation technologies
The processes discussed above are for the direct cat-alytic dehydrogenation of paraffins to the correspond-ing olefins or of olefins to diolefins. Other approacheshave also been considered, although none has reachedthe level of commercialization. Some of the most no-table are
Use of halogens for the dehydrogenation of paraf-fins has been proposed in different ways. For example,heavy paraffins were first chlorinated and then dehy-drochlorinated to heavy olefins commercially in thepast both by Shell and by Huls, among others. Pyrol-ysis of methane in the presence of chlorine has been
proposed by Benson [22] for the production ofacetylene and ethylene. Other chlorination/dehydro-chlorination cycles have been proposed for theproduction of ethylene from ethane. Propane dehydro-genation in the presence of iodine via a propyl iodideintermediate has also been proposed [4,23]. Apartfrom the apparent corrosion problems associated withthe use of halogens, other difficulties readily cometo mind owing to the relatively high cost of chlorine,and even more so of iodine, and the need to eitherdispose of or recycle vast quantities of halogens.
Oxydehydrogenation or oxidative dehydrogenationcan be considered in at least two different ways.
Use of oxygen to oxidize the hydrogen coproductfrom dehydrogenation, and thus to displace the dehy-drogenation equilibrium to higher conversions. Thisapproach has been used commercially in the catalyticdehydrogenation of ethylbenzene to styrene as inthe UOP Styro-PlusTM process or in the ABB Lum-mus/UOP SMARTTM process, but to date has notsucceeded in the dehydrogenation of light or heavyparaffins. This technology has been used in a styreneunit at Mitsubishi Chemicals, Kashima, Japan. Al-though a similar approach has been proposed for thedehydrogenation of paraffins [24–27], it has not beencommercialised.
Direct use of oxygen as a means of dehydrogenat-ing, for example, ethane to ethylene. Oxydehydro-genation has successful commercial applications in theconversion ofn-butenes to butadiene (e.g. as in theOxo-D process referred to earlier), but not yet for theproduction of ethylene or propylene. This subject isanalyzed in more detail in the following section.
Use of oxydehydrogenation relative to straight cat-alytic dehydrogenation must be viewed both in termsof safety issues and in an economic context. On the
412 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
latter, even though oxydehydrogenation offers advan-tages as a means of overcoming thermodynamic equi-librium limitations, it also leads to the total or partialloss of byproduct hydrogen, which in some instancescan have a very significant economic impact.
Although less apparent, oxydehydrogenation alsoplays a role in the work done by BP Amoco, Asahi,and others to extend ammoxidation to the direct con-version of propane to acrylonitrile. It is believed thatthe ammoxidation of propane proceeds through a tran-sient propylene intermediate from which acrylonitrileis derived through a conventional ammoxidation path-way [28].
11.1. Oxydehydrogenation of ethane and propane
Oxydehydrogenation of ethane and propane as aroute to ethylene and propylene, respectively, hasthe attractive feature of removing the equilibriumconversion restriction of dehydrogenation. Table 3contrasts the calculated equilibrium conversions ofoxydehydrogenation of ethane and propane with thecalculated equilibrium conversions for dehydrogena-tion. The formation of water rather than hydrogenin oxydehydrogenation effectively removes the equi-librium constraint on conversion at all temperaturesof interest. At the same time, oxydehydrogenationis not without its own set of challenges that, in thecase of ethane and propane, have kept oxydehydro-genation from being practiced on a commercial scale.Challenges in oxydehydrogenation include handlingmixtures of paraffins and oxygen, which can be ex-plosive at certain compositions, suitable conversion ofparaffins, which often is limited by maintaining a safeparaffin–oxygen composition, and suitable selectivityto olefins. The carbon oxides—carbon monoxide andcarbon dioxide—are thermodynamically more stable
Table 3Percentage conversions at equilibrium for dehydrogenation and oxydehydrogenation of ethane and propane at atmospheric pressurea
a Stoichiometric mixture of alkane and oxygen for oxydehydrogenation.
than the olefins, and so catalysts must be found thatcan stop the reaction at the olefin rather than allowingit to proceed on to the oxides. Also, to compete withsteam cracking, the selectivity to olefins must be quitehigh. Selectivity to ethylene from ethane in steamcracking is reported to be about 84% at 54% ethaneconversion (800◦C, 0.3 kg steam per kg feed, 0.79 sresidence time, and 154 kPa hydrocarbon partial pres-sure), and 78% at 69% ethane conversion (833◦C,0.3 kg steam per kg feed, 0.75 s residence time, and154 kPa hydrocarbon partial pressure) [29].
11.1.1. Ethane oxydehydrogenationEthane oxydehydrogenation as a route to ethylene
has been examined at temperatures in the range of300–500◦C with reducible metal oxide catalysts, andat higher temperatures, above about 600◦C, withlargely non-reducible and reducible metal oxide cata-lysts. The latter have evolved primarily out of investi-gations of methane oxidative coupling [80–82]. In theoxidative coupling of methane, ethane and ethyleneare major products, and ethylene has been shown tobe derived from ethane [80–82]. The dehydrogena-tion or oxydehydrogenation of ethane to ethylene,must occur over these catalysts, and since the overallselectivity is high, the selectivity to ethylene mustbe good. In addition to these two lines of research,recent work at high temperatures using Pt and Pt-Snon a monolith support has been reported.
12. Lower temperature ethaneoxydehydrogenation
Ethane oxydehydrogenation at temperatures in therange of 300–400◦C is conducted with reduciblemetal oxide catalysts usually containing vanadium.
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 413
Early work in this area was conducted by Thorstein-son et al. using a vanadium molybdenum niobiumoxide catalyst [30,31], and a number of papers andpatents based on that work have been issued by UnionCarbide Corp. and other laboratories [32–39]. Ethaneis thought to react with molybdenum or vanadiumin the catalyst to form surface ethoxide, which canthen undergo a beta-elimination process to form ethy-lene. The surface ethoxide can be oxidized furtherto make surface acetate, which leads to acetic acidon hydrolysis. Acetic acid in varying amounts is co-produced with ethylene at pressures greater than oneatmosphere. Later work has shown that catalysts canbe made selective for either ethylene or acetic acidby modifications in the elemental makeup of the cat-alyst and suitable adjustments of reaction conditions[38,39]. Selectivity to ethylene with these catalysts isin the range of approximately 70% at approximately70% conversion of ethane.
Ethane oxydehydrogenation at temperatures inthe range of≥500◦C may be conducted with phos-phorous/molybdenum/antimony oxide catalysts [40]or iron-containing solid solution catalysts stabilizedwith one or more metal oxides [41]. Phosphorous/molybdenum/antimony oxide catalysts, suitably modi-fied with other elements give a selectivity to ethylenein the range of 78% at 20% conversion of ethane. Iron-containing solid solution catalysts stabilized withmetal oxide are optimally used with a flow of hydro-gen chloride and water in addition to ethane and wa-ter. At a contact time of 12 s at 550◦C over an iron in�-alumina solid solution catalyst stabilized with lant-hanum oxide, ethane conversion is in the range of 90%.Products are ethylene (80–93% selectivity) and vinylchloride (1–4% selectivity). As a route to ethylene-only, this route seems unattractive because of the pro-duction of vinyl chloride, which must be separatedfrom ethylene. For vinyl chloride producers, however,where the stream may be able to be used withoutseparation, such a high selectivity/conversion processto ethylene may possibly be economically attractive.
13. Higher temperature ethaneoxydehydrogenation
Ethane oxydehydrogenation at temperatures in therange of 650–800◦C is conducted with a variety
of materials, most not containing the elementsmolybdenum and vanadium, which are commonin lower-temperature catalysts. As opposed to thelower-temperature catalysts which tend to producecoproducts, the higher temperature processes tend toproduce only ethylene as C2 product, though some doproduce methane from cracking reactions. Oxygenatedproducts 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 Instituteof Chemical Physics, China, have shown thata Na2WO4-Mn/SiO2 catalyst (a high-selectivitymethane coupling catalyst) at 700◦C is capable ofgiving greater than 70% selectivity to ethylene (and10% selectivity to methane) at >70% conversion ofethane [42]. Wang and coworkers at the National In-stitute of Materials and Chemical Research, Japan,found that lithium chloride on sulfated zirconia at650◦C yielded 70% selectivity to ethylene (and 2%to methane) at 98% conversion of ethane [43,44]. Al-though the catalyst did show some deactivation withtime, it can still maintain an ethylene yield as highas 50% after 24 h. Lin et al. at the Tokyo Instituteof Technology, Japan, and Chonnam National Uni-versity, Korea, found that a SrBi3O4Cl3 catalyst at640◦C is capable of giving 90% selectivity to ethy-lene at a 25% conversion of ethane [45]. Co-feedingHCl did not change the conversion or the selectivity,but did slow the activity decrease along with the lossof chlorine observed in its absence.
Dang et al. at the Lanzhou Institute of ChemicalPhysics, China, have examined a vanadium oxideon barium carbonate catalyst at 650◦C that gavegood activity [46]. At 34% conversion of ethane,they obtained 76% selectivity for ethylene. Au andcoworkers at Hong Kong Baptist University examinedSrFeO3−δClσ and Ho2O3 with BaCl2 [47,48]. Of thetwo, SrFeO3−δClσ gave the higher yield of ethylene.At 680◦C, SrFeO3−δClσ yielded 90% ethane con-version and 70% selectivity. Over 40 h, there was nodrop-off in yield or in chlorine content of the cata-lyst. Longer-duration tests were not conducted. TheHo2O3 with BaCl2 catalyst at 640◦C yielded 57%conversion and 68% selectivity to ethylene.
Choudhary et al. at the National Chemical Labo-ratory, India, have examined strontium and other rareearth oxides deposited on sintered low surface area
414 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
supports precoated with lanthanum and other rare earthoxides [49,50]. A Sr-Nd2O3 catalyst showed the high-est activity and selectivity for ethane. At 800◦C, theSr-Nd2O3 catalyst was capable of giving 60% con-version of ethane and greater than 80% selectivity toethylene.
Workers at Phillips Petroleum Company have ex-amined catalysts comprising lithium, titanium, andmanganese, and catalysts comprising cobalt, phos-phorous, and at least one promoter selected from a listof elements [51,52]. At 650◦C, the lithium/titaniumcatalyst with manganese gave an ethane conversion of47% and a selectivity to ethylene of 75%. At 670◦C,a catalyst containing cobalt, phosphorous potassium,and zirconium where the cobalt was introduced ascobalt sulfide, and calcining of the catalyst was ac-complished in the absence of oxygen at 670◦C, gavean 85% conversion of ethane and a 86% selectivity toethylene (25 vol.% ethane/75 vol.% air at 3 psig and2400 GHSV). Addition of a halogen-containing com-pound, such as methyl chloride, increases ethyleneyield. There is an indication that the catalyst losesactivity with use over time for oxydehydrogenation ofhydrocarbons. If catalyst activity and selectivity canbe maintained over long periods of time, this catalystseems like a good potential candidate for use in aneconomical process for ethane oxydehydrogenationto ethylene.
Ethane oxydehydrogenation over Pt and Pt-Sncoated monolith catalysts at high temperatures andextremely short contact times has been the subject ofwork by Schmidt and coworkers at the University ofMinnesota for the past several years [53–58]. Workalong similar lines has also been reported to be un-der way in Russia. Similar results are reported overplatinum gauze and platinum coated pellets [59–61].Conversions of ethane on the order of 70% and se-lectivities on the order of 65% are obtained over aPt-coated ceramic foam monolith at approximately1000◦C at contact times on the order of 1 ms. Withthe addition of tin to the catalyst and with the addi-tion of hydrogen to the feed, selectivity to ethylenecould be increased to above 85% at a conversion of
ethane of 67% [56,57]. The authors’ considerationon mechanism follows: “Qualitatively, this processmust involve first the oxidation of H2 to H2O, whichgenerates heat and removes O2, followed by dehy-drogenation of C2H6 to produce C2H4 and H2, andall of these reactions occur within∼10−3 s. Possi-ble mechanisms to explain the results are (i) purelycatalytic reactions on the Pt–Sn surface, (ii) purelyhomogeneous reaction and (iii) catalytic H2 oxidationfollowed by homogeneous ethane decomposition. Wewill consider each of these mechanisms and showthat, while each of these gives partial interpretation ofresults, none appears to be totally satisfactory” [56].
Hydrogen is produced in the process in greateramount than fed, so that hydrogen required for theprocess can be recycled from downstream equipment.Such high selectivity to ethylene coupled with lowcontact time and high conversion of ethane makes thisprocess an attractive possibility as an alternative tosteam cracking of hydrocarbons as a commercial routeto ethylene. While the possibility is bright, a numberof questions need to be resolved. Among these are thequestions of long-term catalyst activity, long-term se-lectivity, and whether such mixtures of ethane, hydro-gen, and oxygen in the ratios needed, reported to be2:2:1 at 950◦C, can be handled safely on a commercialscale. In addition, issues related to process start-ups,shutdowns, and re-starts have to be addressed.
15. Ethane oxydehydrogenation withcarbon dioxide
Another approach that is being pursued in sev-eral laboratories is to use CO2 as a mild oxidant foroxydehydrogenation of hydrocarbons. [42,62–68]. Inaddition to ethane, oxidative dehydrogenation of ethyl-benzene [67] and of propane has also been demon-strated [65,66,68]. Extensive work has been done onoxidative dehydrogenation of ethane with CO2 byXu et al. [62] of the Dalian Institute of ChemicalPhysics. Thus, reaction temperature was significantlylowered from that used in steam cracking, though itis still quite high. High selectivity was obtained at∼800◦C, 1000 h−1 GHSV and 0.1 MPa (∼1 atm) us-ing several catalysts containing oxides of Cr, Cr–Mn,Cr–Mn–Ni, Cr–Mn–La onto a Silicalite-2 (Si-2) ze-olite. Highest conversion/selectivity observed were
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 415
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 con-versions approached equilibrium for these two condi-tions; which were calculated to be 72.8 and 77.6%.
Oxydehydrogenation using CO2 according to theauthors [62] has the following special features.
• CO2 acts as a mild oxidant for the oxidative dehy-drogenation of C2H6 to yield C2H4.
• A significant lowering of the reaction temperatureas compared to the steam cracking process, resultingin lower coke formation on the catalyst.
• C2H6 conversion and C2H4 selectivity are higherin this process than in all other processes of C2H4production from C2H6. Furthermore, no C3+ andC2H2 products are formed.
The dehydrogenation of ethane is facilitated by thereaction of CO2 with H2 (reverse water-gas shift) tomake CO+H2O and also with ethane and methane asshown as follows:
C2H6 = C2H4 + H2
CO2 + H2 = CO+ H2O
C2H6 + 2CO2 = 4CO+ 3H2
CH4 + CO2 = 2CO+ 2H2
In addition, ethane hydrogenolysis to methane alsoprovides favorable free energy of reaction.
C2H6 + H2 = 2CH4
The authors suggest that ethane oxydehydrogena-tion with carbon dioxide takes place according to thefollowing overall stoichiometry:
16C2H6 + 9CO2
= 14C2H4 + 12CO+ 6H2O + 12H2 + CH4
It can be seen from this reaction scheme that in or-der to increase the selectivity to ethylene, the forma-tion of methane must be suppressed and that the keyproblem is to develop a catalyst that can suppress thethermodynamically favorable side reactions.
The work of Xu et al. [62] also demonstratedthat the inclusion of steam in the feed leads to athree-fold reduction in coke formation at 800◦C.These authors also studied the C2H6 + CO2 reac-tion in FCC tail-gas to increase the ethylene content.
Thus, using a tail-gas containing 18.8% C2H6 and19.2% C2H4, the ethylene content was increased to25.4–27.2% depending on the composition of thecatalyst. This reaction was carried out at 1073 K,0.1 MPa, 1000 h−1 GHSV and CO2:C2H6 mole ratioof 1:1. The Cr/Si-2 catalyst gave the lowest ethyleneenhancement (25.4%) while Cr-Mn-La/Si-2 catalystgave the highest ethylene enhancement to 27.2%.The corresponding ethane conversions/selectivity are60.6%/79.6% and 63.6%/85.8%, respectively.
Some of the other early work in this area wasdone using highly selective catalysts for the oxidativecoupling of methane by Liu et al. of the LanzhouInstitute of Chemical Physics [42]. Their work,employing a selective methane coupling catalyst(Na2WO4-Mn/SiO2), showed that >70% conversionand selectivity could be achieved at 700–750◦C andspace velocities of >30,000 h−1, employing O2 as theoxidant. This is in contrast with 53% conversion/97%selectivity at 800◦C and 69.5% conversion/90.5%selectivity at 850◦C employing C2H6:CO2 = 1:1 and3600 h−1 space velocity. These catalysts were sta-ble for 100 h of operation. The authors propose thatsurface lattice oxygen is responsible for selective oxy-dehydrogenation while the bulk lattice oxygen is re-sponsible for deep (non-selective) oxidation of ethane.
Nakagawa et al. [63] studied a series of metal ox-ides for ethane dehydrogenation in the presence ofCO2 at 650◦C, C2H6:CO2 = 5:25 ml/min and SV=900 h−1 ml (g-cat)−1. The order of activity was asfollows:
Ga2O3 > Cr2O3 > V2O5 > TiO2
> Mn3O4 > In2O3 > ZnO > La2O3.
The ethylene selectivity was generally >85% in thepresence of CO2. Ethylene yields for Ga2O3 in thepresence of CO2 was approximately twice than in itsabsence. The ethylene yield enhancement with Cr2O3was only a modest 1.2%, while in the case of V2O5,there was a 2.7% detrimental effect. The authors pro-pose involvement of acid sites in the dehydrogenationreaction.
Wang et al. [64] of Japan studied the effect of sul-fate and Na-treatment silica on the dehydrogenationof ethane with CO2 at 650◦C and 1 atm. Sulfatingsilica (348 m2/g) with (NH4)2SO4 (and calciningat 700◦C/3 h to give∼2 wt.% SO4
2−), and adding
416 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
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% Crcatalyst (3.34 m2/g on silica gave 56.1/92.9/52.1%conversion/selectivity/yield. However, Na-Cr/SiO2gave very poor performance of 4.9/98.2/4.8%. Further-more, the surface area dropped to 3.1 m2/g. In com-parison, silica alone of 348 m2/g gave 2.2/97.2/0.2%conversion/selectivity/yield. Dehydrogenation is pro-posed to occur by abstraction of H by oxygen specieswhich in turn are formed from surface carbonatesdecomposition. At low temperatures, decompositionof carbonates absorbed on strong basic sites restrictsformation and mobility of oxygen resulting in loweractivity of Cr/Na-SiO2 catalyst. However, sulfationof silica favors reaction with hydrocarbon and hence,higher activity (along with formation of CO+ H2).The Cr/SiO2 and Cr/SO42−-SiO2 were shown to bestable for 5–6 h of operation.
Macho and coworkers [65,66] studied oxy-dehydrogenation of propane and butane overMn-Cr-K/Y-Al 2O3 to give propene and butenes, re-spectively, according to the following reaction:
C3H8 + CO2 ⇔ C3H6 + CO+ H2O
C4H10 + CO2 ⇔ C4H8 + CO+ H2O
Splitting of the C–C bond proceeds simultaneouslyleading to the formation of both higher and lower hy-drocarbons as shown in the following equation. Thesereactions are preferred at relatively higher tempera-tures while the relatively lower temperatures preferstraight dehydrogenation with CO2. In addition, cokeformation is more prevalent at relatively higher tem-peratures.
2C3H6 ⇔ C4H8 + C2H4
4C3H8 + 4CO2 ⇔ 3C4H8 + 4CO+ 4H2O
C3H8 + 3CO2 ⇔ C2H4 + 4CO+ 2H2O
C3H8 ⇔ C3H6 + H2
C2H4 + CO2 ⇔ CH4 + 2CO
The authors propose that:
• at middle temperatures (850–1000 K) dehydrogena-tion of alkanes proceeds first and hydrogen gener-ated reacts with CO2 giving CO+ H2O;
• at higher temperatures (>950 K), alkanes may reactdirectly with oxygen left from CO2.
The authors also explain the reactivity of catalystsin terms of acidic and basic nature of supports andthe reactivity of CO2 with such supports. Typically,good active catalysts appear to approach thermody-namic equilibrium. However, temperatures requiredto achieve higher conversions and yields also lead tohigher levels of coking and by-product formation. Noeconomic assessment of the CO2-based dehydrogena-tion processes is available. The authors are not awareof any commercialization efforts using this reaction.
16. Propane oxydehydrogenation
Propane oxydehydrogenation combining highpropane conversion and high propylene selectivity hasproved an elusive goal, and propane oxydehydrogena-tion as a route to propylene appears to be far fromrealizing its commercialization potential. One diffi-culty is that propylene is more easily oxidized thanis propane, so that selectivity tends to decline rapidlywith conversion. Another is that at temperatures aboveabout 700◦C, propane cracking becomes significant,and a variety of products other than propylene areproduced. Efforts to increase propylene selectivityhave centered in the areas of more selective catalystsat temperatures below 700◦C, more selective cata-lysts and conditions at high temperatures, membranereactors, and cyclic-operation reactors.
A variety of catalysts for propane oxydehydrogena-tion have been examined recently at temperatures be-low 700◦C [69–75]. Particularly, notable for its highselectivity to propylene is a vanadia-silica-zirconiacatalyst of Rulkens and Tilley at the University ofCalifornia, Berkley, CA [75]. These workers usinga molecular precursor route, produced an 18/36/46V2O5-SiO2-ZrO2 catalyst which gave 81.5% selec-tivity to propylene at 8% conversion of propane at550◦C. The presence of zirconium is important in re-taining high selectivity, likely by stabilizing the dis-persion of vanadia. These catalysts also are among themost active for propane oxydehydrogenation.
Ranzi and coworkers at the Polytechnic Universityof Milan [76] and Choudhary et al. at the NationalChemical Laboratory of India [77] have examined
M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419 417
oxydehydrogenation of propane at high temperatureswith and without catalysts. Selectivities to olefins,comprising primarily ethylene and propylene, can beobtained in selectivities of 50–80%. Conversion ofpropane is on the order of 90% under these conditions.
Alfonso, Julbe, Farrusseng, Menendez, and Santa-marıa at the University of Zaragoza, Spain, and theLaboratoire des Materiaux et Procedes Membranaires,France, examined the oxidative dehydrogenation ofpropane over V-Al2O3 catalytic membranes which al-low separation of propane and oxygen feeds [78]. Se-lectivities of 51% to propylene were obtained at 8%conversion of propane at 550◦C, which is higher thanthe 44% selectivity obtained at the same temperatureand conversion with premixed feeds.
Creaser et al. at Lulea University of Technol-ogy, Sweden, University of Waterloo, Canada, andChalmers University of Technology, Sweden, exam-ined cyclic operation of the oxidative dehydrogena-tion of propane over a V-Mg-O catalyst [79]. In acyclic reactor where oxygen and propane were alter-nately passed over the catalyst, propylene selectivitywas considerably increased at 510◦C compared tosteady state operation with mixed feeds. At a 1:1propane:oxygen ratio, selectivity to propylene was78% at 4.3% conversion of propane with cyclic op-eration compared to 55% selectivity to propylene at5.5% conversion of propane with mixed feeds. Theseconversions/selectivities are still far from being com-mercially attractive.
17. Summary
It is interesting to note that though oxydehydro-genation is thermodynamically favored at even am-bient conditions, no catalyst has been discovered thatcan provide such a true oxidative conversion at tem-peratures below∼300◦C. Catalysts discovered byUnion Carbide that operate at 300–400◦C [80–82],are the best known catalysts for producing ethylene(with co-production of some acetic acid in varyingamounts). Thus, a true oxydehydrogenation catalystis yet to be discovered. Another explanation mayvery well be that all of the currently known cata-lysts proceed through a dehydrogenation step (C–Hactivation and bond breakage) followed by oxida-tion of the evolved H2 or a surface H-containing
intermediate. This certainly appears to be the case withthe majority, if not all, of the catalysts that are activefor oxydehydrogenation that operate at 500–800◦C.However, under these conditions, it is very difficultto avoid undesirable partial/total oxidation to CO andCO2 via gas phase and surface reactions. Therefore,the challenge for oxydehydrogenation catalysis isto develop highly active and selective catalysts fortotally selective oxidation of alkanes to alkenes.
References
[1] G.F. Hornaday, F.M. Ferrell, G.A. Mills, Manufactureof mono- and diolefins from paraffins by catalyticdehydrogenation, in: Advances in Petroleum Chemistry andRefining, Vol. 4, Interscience, Pans, 1961.
[4] A.H. Weiss, The manufacture of propylene (158th Meeting ofthe American Chemical Society, 10–12 September 1969), in:Refining Petroleum for Chemicals, Advances in ChemistrySeries, Vol. 97, American Chemical Society, Washington, DC,1970.
[5] R.G. Craig, D.C. Spence, Catalytic dehydrogenation ofliquefied petroleum gas by the Houdry Catofin and Catadieneprocesses, in: R.A. Meyers (Ed.), Handbook of PetroleumRefining Processes, McGraw-Hill, New York, 1986 (Chapter4.1).
[6] D. Sanfilippo, F. Buonomo, G. Fusco, I. Miracca, ParaffinsActivation through Fluidized Bed Dehydrogenation: theAnswer to Light Olefins Demand Increase, Elsevier,Amsterdam, Stud. Surf. Sci. Catal. 119 (1998) 919–924.
[7] R. Iezzi, A. Bartolini, U.S. Patent 5,633,421 (27 May 1997),assigned to Snamprogetti.
[8] E.C. Luckenbach, F.A. Zenz, Giovanni Papa, A. Bertolini,U.S. Patent 5,656,243 (12 August 1997), assigned toSnamprogetti.
[9] V. Haensel, U.S. Patent 2,602,772 (8 July 1952), and otherpatents on catalytic reforming with Pt catalysts assigned toUOP.
[10] H.S. Bloch, UOP discloses new way to make linearalkylbenzene, Oil Gas J. 79–81 (1967); U.S. Patent3,448,165 (3 June 1969), and other patents on the catalyticdehydrogenation of paraffinic hydrocarbons assigned toHerman Bloch and UOP.
[11] R.C. Berg, B.V. Vora, Detergent alkylate, in: Encyclopediaof Chemical Processing and Design, Vol. 15, Marcel Dekker,New York, 1982, pp. 266–284.
[12] T. Imai, B.V. Vora, J.C. Bricker, Development ofdehydrogenation catalysts and processes, Petroleum andPetrochemical College, Chulalongkorn University, Bangkok,December 1989.
418 M.M. Bhasin et al. / Applied Catalysis A: General 221 (2001) 397–419
[13] J.H. Gregor, R. Antonelli, T.D. Foley, E.C. Arnold, Increasedopportunities for propane dehydrogenation, DeWitt WorldPetrochemical Review, Houston, TX, 23–25 March 1999.
[14] C.P. Poole, D.S. MacIver, The physical–chemical propertiesof chromia-alumina catalysts, in: Advances in Catalysis, Vol.17, Academic Press, New York, 1967, pp. 223–314.
[15] C.P. Poole, W.L. Kehl, D.S. MacIver, Physical properties ofcoprecipitated chromia-alumina catalysts, J. Catal. 1 (1962)407–415.
[16] E.L. Tucci, J.M. Dufallo, R.J. Feldman, Commercialperformance of the Houdry Catofin process for isobutyleneproduction for MTBE, in: Proceedings of the Workshop onCatalysts and Catalytic Processes, Research Institute, KingFahd University of Petroleum and Minerals, Dhaharan, SaudiArabia, 6 November 1991.
[17] S. Gussow, R. Whitehead, Isobutane dehydrogenation byCatofin as feed for motor fuel ether, in: Proceedings of theNPRA Annual Meeting, San Antonio, TX, 17–19 March1991.
[18] R.O. Dunn, F.M. Brinkmeyer, G.F. Schuette, The PhillipsSTAR process for the dehydrogenation of C3, C4, and C5
paraffins, in: Proceedings of the NPRA Annual Meeting, NewOrleans, LA, 22–24 March 1992.
[19] F. Buonomo, G. Fusco, I. Miracca, G. Papa, D. Sanfilippo,Fluid bed dehydrogenation of light paraffins: the FBDtechnology, in: Proceedings of Petrotech ’96, Bahrain, 10–12June 1996.
[21] A.Q. Campbell, J.M. Garces, T.M. May, R.F. Pogue,Preparation of ethylbenzene and substituted derivatives byalkylation using unpurified reaction products of ethyleneprepared by dehydrogenation of ethane, World Patent WO96/34843 (7 November 1996), assigned to Dow Chemical Co.
[22] S. Benson, Conversion of methane using chlorine, U.S. Patent4,199,533 (22 April 1996), assigned to the University ofSouthern California.
[23] A.H. Weiss, Which propylene process is best? HydrocarbonProcess. 47 (1968) 123–127.
[24] T. Laegreid, M. Ronnekleiv, A. Sobbaken, in: Proceedings ofthe DGMK Conference, 11–12 November 1993, p. 147.
[25] R. Lodeng, P. Soaker, Reactor for catalytic dehydrogenationof hydrocarbons with selective oxidation of hydrogen, U.S.Patent 5,997,826 (7 December 1999), assigned to Den NorskeStats Oljeselskap A.S. (Statoil).
[27] P.A. Agaskar, R.K. Grasselli, J.N. Michaels, P.T. Reischman,D.L. Stern, J.G. Tsikoyiannis, Process for the catalyticdehydrogenation of alkanes to alkenes with simultaneouscombustion of hydrogen, U.S. Patent 5,430,209 (4 July 1995),assigned to Mobil Oil Corp.
[28] J.F. Brazdil, A.P. Cavalcanti, J.P. Padolewski, Method forpreparing vanadium antimony oxide-based oxidation andammoxidation catalysts, U.S. Patent 5,693,587 (2 December1997), assigned to The Standard Oil Company of Ohio(Sohio/BP).
[31] E.M. Thorsteinson, T.P. Wilson, F.G. Young, P.H.U.S. Kasai,U.S. Patent 4,250,346 (1981), assigned to Union Carbide.
[32] J.H. McCain, U.S. Patent 4,524,236 (18 June 1985), assignedto Union Carbide.
[33] R.M. Manyik, J.L. Brockwell, J.E. Kendall, U.S. Patent4,899,003 (6 February 1990), assigned to Union Carbide.
[34] K. Ruth, R. Kieffer, R. Burch, J. Catal. 175 (1998) 16.[35] K. Ruth, R. Burch, R. Kieffer, J. Catal. 175 (1998) 27.[36] M. Kitson, U.S. Patent 5,260,250 (9 November 1993),
assigned to BP Chemicals.[37] W. Ueda, N.F. Chen, K. Oshihara, Kinet. Catal. 40 (1999)
401 (English).[38] W. Ueda, N.F. Chen, K. Oshihara, Chem. Commun. 517
(1999).[39] C. Hailett, European Patent Application, Publication
0,480,594,A2 (15 April 1992), assigned to B.P. Chemicals.[40] M. Kutyrev, F. Cavani, F. Trifiro, European Patent
Application, Publication 0,544,372,A1 (2 June 1993),assigned to Eniricherche S.p.A.
[41] J. Magistro, U.S. Patent 5,087,791 (11 February 1992),assigned to B.F. Goodrich.
[42] Y. Liu, J. Xue, X. Liu, R. Hou, S. Li, in: A. Parmaliana, etal. (Eds.), Natural Gas Conversion V, Elsevier, Amsterdam,Stud. Surf. Sci. Catal. 119 (1998) 593,
[43] S. Wang, K. Murata, T. Hayakawa, S. Hamakawa, K. Suzuki,Catal. Lett. 59 (1999) 173.
[44] S. Wang, K. Murata, T. Hayakawa, S. Hamakawa, K. Suzuki,Chem. Commun. 103 (1999).
[45] S.W. Lin, Y.C. Kim, W. Ueda, Bull. Chem. Soc. Jpn. 71(1998) 1089.
[46] Z. Dang, J. Gu, J. Lin, D. Yang, Catal. Lett. 54 (1998) 129.[47] H.X. Dai, C.F. Ng, C.T. Au, Catal. Lett. 57 (1999) 115.[48] C.T. Au, K.D. Chen, H.X. Dai, Y.W. Liu, C.F. Ng, Appl.
Catal. A 177 (1999) 185.[49] V.R. Choudhary, S.A.R. Mulla, V.H. Rane, J. Chem. Technol.
[75] R. Rulkens, T.D. Tilley, J. Am. Chem. Soc. 120 (1998)9959.
[76] A. Beretta, P. Forzatti, E. Ranzi, J. Catal. 184 (1999) 469.[77] V.R. Choudhary, V.H. Rane, A.M. Rajput, AIChE J. 44 (1998)
2293.[78] M.J. Alfonso, A. Julbe, D. Farrusseng, M. Menendez, J.
Santamarıa, Chem. Eng. Sci. 54 (1999) 1265.[79] D. Creaser, B. Andersson, R.R. Hudgins, P.L. Silverston,
Chem. Eng. Sci. 54 (1999) 4437.[80] G.E. Keller, M.M. Bhasin, Synthesis of ethylene via oxidative
coupling of methane. 1. Determination of active catalysts, J.Catal. 73 (1999) 9.
[81] M.M. Bhasin, Feasibility of ethylene synthesis via catalyticoxidative coupling of methane, in: paper presented at theMethane Conversion Symposium, Auckland, New Zealand,27 April to 1 May 1987.
[82] M.M. Bhasin, Feasibility of Ethylene Synthesis via OxidativeCoupling of Methane, Elsevier, Amsterdam, Stud. Surf. Sci.Catal. 36 (1988) 343.
![Effect of Excess Iron on Oxidative Dehydrogenation of 1-Butene … · 2019-08-24 · catalyst stoichiometry for the oxidative dehydrogenation of n-butane [22]. They showed that by](https://static.documents.pub/doc/80x56/5faa538d87db295f1775edaf/effect-of-excess-iron-on-oxidative-dehydrogenation-of-1-butene-2019-08-24-catalyst.jpg)
