PROCESS FOR THE STEAM DEHYDROGENATION OF DEHYDROGENATABLE HYDROCARBONS WITH SIMULTANEOUS OXIDATIVE REHEATING

Abstract: Dehydrogenatable hydrocarbons are subjected to steam dehydrogenation in the vapor phase in conjunction with oxidative reheating of the intermediate products, utilizing the same catalyst to perform both the selective oxidation and steam dehydrogenation functions. The particular catalyst employed comprises a Group VIII noble metal component, an alkali metal component and a third component selected from boron, gallium, indium, thallium, germanium, tin and lead, all supported on a refractory inorganic oxide such as alumina.

Description: This invention relates to a process for the catalytic dehydrogenation of dehydrogenatable hydrocarbons where steam and an oxygen-containing gas are co-feeds in order to perform dehydrogenation and oxidative reheat functions simultaneously. The dehydrogenation of hydrocarbons is an important commercial process. This is because of the great demand for dehydrogenated hydrocarbons as feedstocks for industrial processes. For example, dehydrogenated hydrocarbons are utilized in the manufacture of various products such as detergents, high octane gasolines, and pharmaceutical products among others. Plastics and synthetic rubbers are other products which may be produced through use of dehydrogenated hydrocarbons. One example of a specific dehydrogenation process is dehydrogenating isobutane to produce isobutylene which may then be polymerized to provide tackifying agents for adhesives, viscosity-index additives for motor oils and impact-resistant and anti-oxidant additives for plastics. The prior art recognizes numerous hydrocarbon conversion processes which utilize a single catalyst to perform dehydrogenation and selective oxidation functions. The catalysts disclosed as being useful in such a process cover a wide range of compositions. Catalysts which are composed primarily of metal oxides comprise one group of catalysts useful in a selective oxidation/dehydrogenation process. US-A-3,502,737 discloses the use of an iron oxide catalyst in a process to dehydrogenate ethylbenzene into styrene and also to selectively oxidize hydrogen produced in the dehydrogenation reaction. US-A-4,180,690 discloses the use of a silica-based catalyst containing oxides of cobalt, chromium, aluminum, and magnesium for the oxydehydrogenation of an alkylaromatic. US-A-4,672,146 describes a process utilizing a catalytic composite useful for the dehydrogenation of dehydrogenstable hydrocarbons. The catalyst contains, among other components, a Group VIII noble metal and an alkali metal component on an alumina support. However, from a process standpoint, that patent fails to mention the utility of such a catalyst in the selective oxidation of hydrogen in the presence of oxygen at oxidation conditions. In fact, it states a diluent may be introduced into the dehydrogenation process, but it mentions only hydrogen, steam, methane, ethane, carbon dioxide, nitrogen, and argon, and specifically hydrogen as the diluent. Oxygen addition as a process co-feed or diluent is never mentioned in that patent. US-A-4652,687, US-A-4,565,898 and US-A-4,418,237 similarly disclose the use of a Group VIII metal-containing oxidation catalyst in a dual catalyst oxidation/dehydrogenation process. US-A-4,652,687 and US-A-4,565,898 specifically disclose the use of an oxidation catalyst comprising a Group VIII metal, a Group IVA metal, and a Group IA or IIA metal on an alumina support which has been calcined at between 900° and 1500°C in conjunction with a dehydrogenation catalyst. They go on to further support the use of a dehydrogenation catalyst that is different in nature from the oxidation catalyst, and specifically a dehydrogenation catalyst comprising an alkali metal-promoted iron compound. US-A-4,418,237 discloses the use of an oxidation catalyst comprising a Group VIII metal and a second metal component selected from the group consisting of rubidium, cesium, barium, francium, and radium on a porous alumina support. This oxidation catalyst is utilized in conjunction with a different dehydrogenation catalyst, and preferably a dehydrogenation catalyst again comprising an alkali metal-promoted iron compound. All three patents therefore disclose the use of an oxidation catalyst only in conjunction with a separate and different dehydrogenation catalyst and especially a dehydrogenation catalyst comprising an alkali metal-promoted iron compound. The present invention seeks to provide an improved process for the steam-promoted oxidation/dehydrogenation of hydrocarbons. It has been found that catalysts of a particular composition are not only useful as the dehydrogenation catalyst but are also surprisingly selective towards the oxidation of hydrogen in the presence of oxygen, hydrogen, and hydrocarbons, resulting in an increase in the yield of the desired dehydrogenated product. According to the present invention there is provided a process for the steam oxidation/dehydrogenation of a C₂-C₃₀ dehydrogenatable hydrocarbon by contacting the hydrocarbon steam and an oxygen-containing gas in a reaction zone with simultaneous oxidative reheating by selective combustion of hydrogen derived at least in part from the hydrocarbon using only one catalyst, said catalyst comprising a Group VIII noble metal component, a second component selected from alkali metals (lithium, sodium, potassium, rubidium, cesium, francium and mixtures thereof) and a third component selected from boron, gallium, indium, thallium, germanium, tin, lead and mixtures thereof, all on a refractory inorganic oxide support. In a preferred embodiment, the features of which are preferred separately or together, a steam oxidation/dehydrogenation process is provided which solely utilizes a catalyst comprising from 0.01 to 10 wt.% platinum or palladium, from 0.01 to 20 wt.% potassium, rubidium, or cesium and from 0.01 to 5 wt % tin on an alumina support having a surface area of from 5 to 120 m²/g to promote the oxidative steam dehydrogenation of a feed stream comprising an oxygen-containing gas, steam, and a C₂-C₃₀ dehydrogenatable hydrocarbon. In an alternative preferred embodiment, the features of which are preferred separately or together, the oxidative steam dehydrogenation process is further characterized in that a dehydrogenatable hydrocarbon comprising one or more C₂-C₁₅ paraffins and steam at a steam to hydrocarbon molar ratio of from 0.1:1 to 40:1 are introduced into the first reaction zone of a reactor containing a plurality of reaction zones at conditions including a pressure from 10 to 1,000 kPa (0.1 to 10 atmospheres), a temperature of from 400° to 900°C, and a liquid hourly space velocity of from 0.1 to 100 hr⁻¹. An oxygen-containing gas is then introduced into the second, and thereafter every other, reaction zone of the plurality of reaction zones such that the overall oxygen rate to all of the reaction zones ranges from 0.01 to 2 moles of oxygen per mole of C₂-C₁₅ paraffin. Figures 1 and 2 of the accompanying drawing are graphical representations of the oxidation performance of a catalyst used in the process of the present invention. Figure 1 is a graph of the oxygen conversion percentage as a function of hours on-stream. Figure 2 is a graph of the hydrogen oxidation selectivity of the catalyst of the process as a function of hours on-stream. The use of only one catalyst to perform both the selective oxidation and dehydrogenation reaction in a selective steam oxidation/dehydrogenation process overcomes many of the problems inherent with a two-catalyst process where each catalyst has only one function. Specifically, the use of only one catalyst will overcome the catalyst handling problems encountered in loading and unloading a two-catalyst system. In a two-catalyst system, the catalysts must be carefully loaded if contained in a single bed, or two or more reactors must be used when the catalyst is loaded in separate reactors. With only one catalyst, the reactor and catalyst loading become simpler and more economical. Additionally, a process using only a single catalyst is much more useful in a moving catalyst system where the catalyst is continuously being regenerated. Thus, the use of a single dual functional catalyst is an important aspect of the overall process. As indicated above, one feature of the catalytic composite of the process of this invention is a noble metal component from Group VIII of the Periodic Table of Elements. The Group VIII noble metal may be platinum, palladium, iridium, rhodium, osmium, ruthenium, or mixtures thereof. Platinum or palladium are, however, the preferred Group VIII noble metal components. It is believed that substantially all of the Group VIII noble metal component exists within the catalyst in the elemental metallic state. The Group VIII noble metal component generally will constitute 0.01 to 10 wt.%, calculated on an elemental basis, of the final catalytic composite. Preferably, the catalyst comprises 0.1 to 5 wt.% Group VIII noble metal component, especially 0.1 to 5 wt.% platinum component. The Group VIII noble metal component may be incorporated in the catalytic composite in any suitable manner such as, for example, by coprecipitation or cogelation, ion exchange or impregnation, or deposition from a vapor phase or from an atomic source or by like procedures either before, while, or after other catalytic components are incorporated. The preferred method of incorporating the Group VIII noble metal component is to impregnate the refractory oxide support with a solution or suspension of a decomposable compound of a Group VIII noble metal. For example, platinum may be added to the support by commingling the latter with an aqueous solution of chloroplatinic acid. Another acid, for example nitric acid or other optional components may be added to the impregnating solution to further assist in dispersing or fixing the Group VIII noble metal component in the final catalyst composite. The Group VIII noble metal component may be located upon the catalyst in a variety of useful manners known in the art including uniformly dispersed, surface-impregnated or surface concentrated among others. An alkali metal component, also known as a Group IA component, is also an important ingredient of the catalyst of the process of this invention. The alkali metal component of the present invention may be cesium, rubidium, potassium, sodium, lithium, or francium, or mixtures thereof. Potassium, rubidium, cesium, or mixtures thereof, are the preferred alkali components. Preferably, the alkali metal component is well dispersed throughout the catalytic composite. The alkali metal component generally constitutes 0.01 to 20 wt.%, calculated on an elemental basis, of the final catalytic composite. The alkali metal component is preferably present in an amount ranging from 0.1 to 5.0 wt.% of the catalytic composite. The alkali metal component may be incorporated in the catalytic composite in any suitable manner such as, for example, by coprecipitation or cogelation, by ion exchange or impregnation, or by like procedures either before, while, or after other catalytic components are incorporated into the catalyst composition. A preferred method of incorporating the preferred alkali metal components is to impregnate the carrier material with a solution of potassium chloride or cesium nitrate. In a most preferred embodiment, the potassium and/or cesium is present in the catalytic composite in an amount ranging from 0.1 to 1.5 wt.%. The third component of the catalyst for the process of the present invention is selected from boron, gallium, indium, thallium, germanium, tin, and lead and mixtures of them. A Group IVA metal component is the preferred third component, with tin being the most preferred third component. The third component is typically present in amounts ranging from 0.1 to 20.0 wt.%. It is especially desired that the third component, e.g. a Group IVA metal component, is a co-formed component. Preferably, the co-formed Group IVA component is well dispersed throughout the catalyst. The co-formed Group IVA component generally constitutes 0.01 to 5 wt.%, calculated on an elemental basis, of the final catalyst composite. Preferably, the catalyst comprises 0.1 to 2 wt.% co-formed Group IVA component, especially 0.1 to 2 wt.% co-formed tin. A method of incorporating a co-formed tin component into a catalyst is by cogelling it during preparation of the catalyst support. For example, a soluble tin compound such as stannous or stannic chloride may be mixed with a hydrosol. Thereafter, a gelling agent such as hexamethylenetetramine is admixed with the hydrosol and tin compound and the resulting mixture is dropped into a hot oil bath forming spheres comprising an intimate mixture of the catalyst support and tin component. The other catalytic modifier (third) components mentioned may be incorporated into the catalyst by any method known in the art such as surface impregnation or deposition, uniform impregnation or any other known distribution. Such methods include impregnation, coprecipitation, and vapor condensation among others. Optional components which may be incorporated into the catalyst of the process of the present invention, e.g. by the above methods, include Group IIA metal components including beryllium, magnesium, calcium, strontium, and barium. A halogen component may also be incorporated into the catalyst of the process of the present invention. Suitable halogen components include fluorine, chlorine, bromine, and iodine. The halogen component may be added to the catalyst composite during the incorporation of the carrier material with other catalyst components, for example by using chloroplatinic acid to impregnate the platinum component. Also, the halogen component or a portion thereof may be added to the catalyst composite by contacting the catalyst with the halogen or a compound or solution containing the halogen before or after other catalyst components are incorporated with the carrier material. Suitable compounds containing the halogen include acids containing the halogen, for example hydrochloric acid. Or, the halogen component or a portion thereof may be incorporated by contacting the catalyst with a compound or solution containing the halogen in a subsequent catalyst regeneration step. In the regeneration step, carbon deposited on the catalyst as coke during use of the catalyst in a hydrocarbon conversion process is burned off the catalyst and the platinum group component on the catalyst is redistributed to provide a regenerated catalyst with performance characteristics much like the fresh catalyst. The halogen component may be added during the carbon burn step or during the platinum group component redistribution step, for example by contacting the catalyst with a hydrogen chloride gas. Also, the halogen component may be added to the catalyst composite by adding the halogen or a compound or solution containing the halogen, such as propylene dichloride for example, to the hydrocarbon feed stream or to the recycle gas during operation of the hydrocarbon conversion process. The refractory inorganic oxide carrier material useful for the catalyst of this process may be any known carrier material of this type which is useful as a catalytic support. However, alumina is the most preferred support material. The most preferred inorganic oxide support of the present invention is alumina having a surface area of from 1 to 500 m²/g, and preferably from 5 to 120 m²/g. The alumina carrier material may be prepared in any suitable manner from synthetic or naturally occurring raw materials. The carrier may be formed in any desired shape such as spheres, pills, cakes, extrudates, powders, granules, etc., and it may be utilized in any particle size. A preferred shape of alumina is the sphere. A preferred particle size is abort 1.6 mm (1/16-inch) in diameter, though particles as small as about 0.8 mm (1/32-inch), and smaller, may also be utilized, as well as particles larger than 1.6 mm (1/16-inch) diameter. In a most preferred method, the alumina is in the form of spheres. To make alumina spheres, aluminum metal is converted into an alumina sol by reacting it with a suitable peptizing acid and water, and then dropping a mixture of the sol and a gelling agent into a hot oil bath. The mixture forms spherical particles of an alumina gel in the hot oil bath which are easily converted into the preferred gamma- or eta-alumina carrier material by known methods including aging, drying, and calcining. Other shapes of the alumina carrier material may also be prepared by conventional methods. A Group IVA metal component as previously mentioned may be added to the alumina sol or alumina carrier material prior to substrate formation. The preferred alumina support having a surface area of between 5 and 120 m²/g of the present invention may be prepared by subjecting a precalcined alumina support to an additional calcination step. As indicated heretofore, in preparing an alumina support, the support is typically calcined and dried. However, it has been discovered that the novel pore distribution of the preferred catalyst of the present invention may be induced by a calcination step performed subsequent to the initial calcination. Accordingly, an alumina support optionally containing a co-formed Group IVA metal component is subjected to a calcination step conducted at from about 800° to about 1200°C. The calcination may be conducted in air, steam, or a mixture thereof. When a mixture is employed, generally the steam will comprise from 1% to 50%. The catalyst previously described as useful in the process of the instant invention is very useful in both the dehydrogenation of dehydrogenatable hydrocarbons in the presence of steam, and also in the selective oxidation of hydrogen produced by the dehydrogenation reaction. Any dehydrogenatable hydrocarbon may be utilized as feed to the present invention. However, it is preferred that the hydrocarbons which may be dehydrogenated include dehydrogenatable hydrocarbons having from 2 to 30 or more carbon atoms including paraffins, alkylaromatics, naphthenes, and olefins. One group of hydrocarbons which can be dehydrogenated with the catalyst is the group of paraffins having from 2 to 30 or more carbon atoms. The catalyst is particularly useful for dehydrogenating paraffins having from 2 to 15 or more carbon atoms to the corresponding monoolefins or for dehydrogenating monoolefins having from 2 to 15 or more carbon atoms to the corresponding diolefins or acetylene derivatives. The catalyst is especially useful in the dehydrogenation of C₂-C₆ paraffins into monoolefins. The dehydrogenation of a dehydrogenatable hydrocarbon such as, for example, those mentioned above, is effected in the instant process by contacting the dehydrogenatable hydrocarbon and steam, with the previously described catalyst at conditions useful for both the steam dehydrogenation of hydrocarbons and the selective oxidation of hydrogen. Such conditions comprise temperatures which range from 400° to 900°C, and a reaction pressure in the range of from 0.1 to 10 atmospheres; the exact dehydrogenation conditions are, however, a function of the particular dehydrogenatable hydrocarbon undergoing dehydrogenation. Other reaction conditions will include a liquid hourly space velocity based on the hydrocarbon charge of from 0.1 to 100 hr⁻¹ and steam to hydrocarbon molar ratios ranging from 0.1:1 to 40:1. The dehydrogenation of hydrocarbons is an endothermic process. In a system employing a dehydrogenation catalyst only, it is typically necessary to add superheated steam at various points in the process or to intermittently remove and reheat the reaction stream between catalyst beds. In an improvement, processes were developed which utilized a two-catalyst system with distinct beds or reactors of dehydrogenation or selective oxidation catalysts. The purpose of the selective oxidation catalysts was to selectively oxidize the hydrogen produced as a result of the dehydrogenation reaction with oxygen that had been added to the oxidation zone to generate heat internally in the process. The heat generated would typically be sufficient to cause the reaction mixture to reach desired dehydrogenation temperatures for the next dehydrogenation step. The instant invention is an improvement of this previously mentioned system. As has been previously mentioned, in this invention, one specific catalyst can be used to accomplish both the dehydrogenation and oxidation reactions. Before explaining the preferred reactor configurations, more details of the oxidation aspect of the invention are disclosed. The selective oxidation step of this process utilizes at least in part the hydrogen which has been produced in the dehydrogenation step of the process to supply heat to the next dehydrogenation reaction section. To accomplish this, an oxygen-containing gas is first introduced into the reactor, preferably at a point adjacent to the selective oxidative catalyst section. The oxygen in the oxygen-containing gas is necessary to oxidize the hydrogen contained in the reaction stream. Examples of oxygen-containing gases which may be utilized to effect the selective oxidation of the hydrogen which is present will include air, oxygen, or air or oxygen diluted with other gases such as steam, carbon dioxide and inert gases such as nitrogen, argon, helium, etc. The amount of oxygen which is introduced to contact the process stream may range from 0.01:1 to 2:1 moles of oxygen per mole of hydrogen contained in the process stream at the point where oxygen is added to the process stream. In the selective oxidation reaction, the process stream which comprises unreacted dehydrogenatable hydrocarbon, dehydrogenated hydrocarbon, and hydrogen is reacted with oxygen in the presence of the selective steam oxidation/dehydrogenation catalyst whereby hydrogen is selectively oxidized to produce water with very little of the oxygen reacting with the process stream hydrocarbons. The oxygen-containing reactant may be added to the instant process in various ways such as by admixing oxygen with a relatively cool hydrocarbon feed stream or with the steam diluent, or it may be added directly to the reactor independently of the feed hydrocarbons or the steam diluent. In addition, the oxygen-containing reactant can be added at one or more points in the reactor in such a fashion as to minimize local concentrations of oxygen relative to hydrogen in order to distribute the beneficial temperature rise produced by the selective hydrogen oxidation over the entire length of the reaction zone. In fact, using a plurality of injection points for introducing the oxygen-containing gas into the steam oxidation/dehydrogenation reaction zone is a preferred mode of operation for the instant process. This procedure minimizes the opportunity for local build-up of the concentration of oxygen relative to the amount of hydrogen, thereby minimizing the opportunity for undesired reaction of the oxygen-containing gas with either feed or product hydrocarbons. It is also an aspect of this invention that hydrogen may be a co-feed of the process. Hydrogen if needed would typically be injected into the oxidation section of the steam oxidation/dehydrogenation process. The hydrogen would be added as fuel, so to speak, to elevate the temperature of the process stream. The hydrogen could be added to the inlet of the reactor with the oxygen-containing gas to elevate the temperature of the feed stream to dehydrogenation reaction conditions, or anywhere desired and effective in the reaction zone. According to the unique process of the present invention, a mixture of dehydrogenatable hydrocarbons, steam, and an oxygen-containing gas is contacted with the catalytic composite of the present invention in a steam oxidation/dehydrogenation reaction zone maintained at steam oxidation/dehydrogenation conditions. This contacting step may be accomplished by using the catalyst in a fixed bed system, a moving bed system, a fluidized bed system, or in a batch type operation. However, in view of the fact that the attrition losses of the valuable catalyst should be minimized and of the well known operational advantages, it is preferred to use either a fixed bed catalytic system, or a dense phase moving bed system such as is shown in US-A-3,725,249. If a fixed bed catalytic reaction system is used for the process of the present invention, it is anticipated that the reaction system could take many forms. The first possibility is that the reaction would comprise a single reaction zone within a single reactor with single inlet and outlet ports. All co-feeds would enter the inlet of the reactor and products and by-products would leave the system through the reactor outlet port. A permeation of this flowscheme would be to have a reaction zone having one or more oxygen-containing gas inlet ports separate from the inlet or outlet ports so that oxygen could be injected at different points into the catalyst bed of the reaction zone. This type of configuration would mimic a reaction system with a plurality of catalyst beds only in this instance, it would be a single reaction zone in which the oxidation and dehydrogenation function of the catalysts was dispersed throughout the bed in a plurality of locations. Of course, the catalyst could be distributed in a plurality of distinct reactors each having a specific function, either steam dehydrogenation or oxidation, or the fixed bed reactor system could comprise a combination of the two possibilities. The choice of the best system would likely depend upon many factors including available equipment, particular reaction speed, and efficiency, as well as optimum reaction conditions for the two distinct reactions being performed. Regardless of the fixed bed catalyst system employed, the hydrocarbon may be contacted with the catalyst bed in either upward, downward, or radial flow fashion. Radial flow of the hydrocarbon through the catalyst bed is preferred for commercial scale reactors. The hydrocarbon must be in the vapor phase when it contacts the catalyst. It should be noted that in a fixed bed reaction system that there are two functions the catalyst is performing in different ares of the catalyst beds. Of course, these two catalytic functions are steam dehydrogenation and selective oxidation. The distinct areas where these reactions occur are referred to herein as reaction zones. These reaction zones may be distinct, wherein a mechanical means is present which separates a volume of catalyst which is used to dehydrogenate the reaction mixtures from a volume of catalyst that is used to selectively oxidize hydrogen. On the other hand, these zones may be indistinct where there is no physical means used to separate the reaction zones. In either situation, the function the catalyst is performing will effectively separate the reaction zones. Where there are distinct reaction zones within the steam oxidation/dehydrogenation reaction zone, it is preferable that every other reaction zone of a plurality of reaction zones performs the same catalytic function. That is to say, if the first reaction zone is used to dehydrogenate the hydrocarbon, then the second zone is used for selective oxidation, the third zone for dehydrogenation, the fourth zone for oxidation, and so on. The same holds true if the first reaction zone of a plurality of reaction zones is used to selectively oxidize hydrogen. This description holds true for a process of the instant invention which employs a moving bed of catalyst also. The process of the instant invention can best be accomplished in a moving bed catalytic system. Such a system is described in US-A-3,725,249 and is most useful for use in a reaction where the catalyst becomes deactivated by coke deposition thereon. In such a situation, the catalyst of the instant process would continuously move through the plurality of reaction zones of the process and once deactivated, be transported to a continuous catalyst regeneration system. Once regenerated, the catalyst would be returned to the reaction system. It is an aspect of this invention that the selective steam oxidation/dehydrogenation conversion process be a complete process. That is to say, the process will comprise a reaction section and other sections such as gas recycle, liquid recycle, product recovery, and the like such that the process is viable and efficient. Examples of some of the product recovery techniques that could be employed alone or in combination in the product recovery zone of a hydrocarbon conversion process are: distillation including vacuum, atmospheric, and superatmospheric distillation; extraction techniques including, for example, liquid/liquid extractions, vapor/liquid extractions, supercritical extractions and others; absorption techniques, adsorption techniques, and any other known mass transfer techniques which can achieve the recovery of the desired products. The following examples are introduced to further describe the process of the invention. The examples are intended as illustrative embodiments and should not be considered to restrict the otherwise broad interpretation of the invention as set forth in the claims appended hereto. EXAMPLE I Precalcined spheroidal theta-alumina particles having 0.5 wt.% tin and an average surface area of approximately 86 m²/g and averaging about 1.6 mm (1/16-inch) in diameter were consecutively impregnated with 0.75 wt.% platinum and 3.5 wt.% cesium. Precalcination of the alumina spheres was effected in flowing air at a temperature of about 1080°C over about a 3-hour interval. Impregnation was accomplished with a chloroplatinic acid solution in a rotary steam evaporator or dryer, while evaporating the solution to dryness. The platinum-impregnated spheres were subsequently calcined in air containing 10% steam for 2 hours at about 570°C. The calcined, platinum-impregnated spheres were next impregnated with cesium from an aqueous solution of cesium nitrate and further calcined in air for about 3 hours at 566°C followed by reduction with hydrogenation at 566°C for 2 hours. The resulting catalytic composition is hereinafter referred to as Catalyst A. EXAMPLE II In this example, Catalyst A was evaluated in a first test as a dehydrogenation catalyst, and in a second test as an oxidation/dehydrogenation catalyst. These comparison tests were carried out in a pilot plant comprising two reactors in series, and product separation facilities. All feed components were introduced into the inlet of the first reactor. All operating variables were kept constant between tests except that the first reactor feed temperature was lower in the second test so that the feed to the second reactor could be heated by selective hydrogen oxidation to the second reactor feed temperature. A detailed comparison of the operating variables is found in Table 1 below. Pilot plant test 2 was designed to model a first oxidation reaction zone and a second dehydrogenation reaction zone. That is to say, the feedstock for both tests mimics a dehydrogenation reaction zone intermediate product. Additionally, nitrogen was added to the first, dehydrogenation only, test to maintain the reactor space velocities. The second test used air to maintain these space velocities and to promote hydrogen oxidation. The following Tables 2 and 3 contain information about the pilot plant results of both tests. Propane Conversion (Mole %) Hours On-Stream Test #1 Test #2 Rx #1 Out Rx #2 Out Rx #1 Out Rx #2 Out 2 0.5 - 0.5 15.0 6 3.5 14.5 2.5 16.0 10 3.0 13.0 1.5 13.7 14 2.5 11.5 1.4 11.5 18 2.5 11.0 1.0 11.5 22 2.4 9.5 0.8 10.8 26 2.0 9.2 0.2 10.0 30 1.9 8.5 0.5 9.1 34 1.5 7.6 0.4 8.3 38 2.0 6.9 0.2 7.4 Selectivity to Propylene (Mole %) Hours On-Stream Test #1 Test #2 Rx #1 Out Rx #2 Out Rx #1 Out Rx #2 Out 2 - 98.6 100 98.5 6 98.0 96.4 - 97.0 10 97.5 96.0 100 96.2 14 98.0 96.1 100 96.3 18 97.5 96.0 100 96.0 22 97.3 95.6 100 95.5 26 97.0 95.6 100 95.5 30 96.4 96.1 100 95.0 34 96.2 96.0 100 94.6 38 95.6 94.9 100 95.5 A review of the dehydrogenation results of the two tests indicates that the results are similar for both tests. The propane conversion (at Reactor #2 outlet) is almost identical for both tests with the propane conversion for the oxidation/dehydrogenation Test 2 being slightly better than Test 1 overall. The first reactor outlet conversion is generally lower in the oxidation case in Test 2 than in comparison to the dehydrogenation case in Test 1 because the average temperature is lower in the first reactor of the second test. The propylene selectivity data (that is the percent of the converted product that is propylene) indicates that the selectivity is similar for the reactor #2 in both tests. This is evidenced by comparing the reactor #2 outlet selectivity data for Tests 1 and 2. The reactor #1 outlet selectivity is higher in Test 1 than Test 2. Figures 1 and 2 represent the oxidation zone performance of the catalyst of reactor 1 in the second test. Figure 1 represents the oxygen (O₂) conversion (the percent of the oxygen converted to another form) versus time-on-stream. Figure 1 indicates that O₂ is totally oxidized in the oxidation zone during the test period. Figure 2 represents the selectivity of the oxygen to desirably combust hydrogen, that is, the amount of oxygen converted that oxidizes hydrogen. From Figure 2, it can be seen that the desirable hydrogen oxidation selectivity of the selective steam oxidation/dehydrogenation catalyst of the instant process is high at 83% and increases further to a very acceptable value of 98%. The catalyst of the process of the instant invention is therefore a very efficient oxidation catalyst as it completely oxidized all added oxygen while maximizing hydrogen oxidation, thus minimizing the combustion of valuable hydrocarbons with the oxygen. Overall the process produces a dehydrogenated product in greater amounts with a single catalyst than does a single catalyst-containing dehydrogenation process of the prior art.

Claims: A process for the steam dehydrogenation of a C₂-C₃₀ dehydrogenatable hydrocarbon by contacting the dehydrogenatable hydrocarbon steam and an oxygen-containing gas in a reaction zone with simultaneous oxidative reheating by selective combustion of hydrogen derived at least in part from the hydrocarbon, using only one catalyst, said catalyst comprising a Group VIII noble metal component, a second component selected from lithium, sodium, potassium, rubidium, cesium, francium and mixtures thereof, and a third component selected from boron, gallium, indium, thallium, germanium, tin, lead and mixtures thereof, all on a refractory inorganic oxide support. A process as claimed in cliam 1, characterized in that the dehydrogenatable hydrocarbon comprises one or more C₂-C₁₅ paraffins. A process as claimed in claim 1 or 2, characterized in that the Group VIII noble metal component is platinum or palladium. A process as claimed in any of claims 1 to 3, characterized in that the second component is cesium or potassium or a mixture thereof. A process as claimed in any of claims 1 to 4, characterized in that the third component is tin. A process as claimed in any of claims 1 to 5, characterized in that the catalyst comprises from 0.01 to 10 wt.% platinum or palladium, 0.01 to 20 wt.% potassium, rubidium, or cesium, or a mixture thereof, and 0.01 to 5 wt.% tin on an alumina support having a surface area of from 5 to 120 m²/g. A process as claimed in any of claims 1 to 6, characterized in that the reaction is carried out at a steam to hydrocarbon molar ratio of 0.1:1 to 40:1, a pressure from 10 to 1000 kPa (0.1-10 atmospheres), a temperature of 400° to 900°C, a liquid hourly space velocity (LHSV) of 0.1 to 100 hr⁻¹ and an oxygen to hydrocarbon mole ratio of 0.01:1 to 2:1. A process as claimed in any of claims 1 to 7. characterized in that the reaction is carried out in a reactor containing a plurality of reaction zones in series, the hydrocarbon and steam being introduced into the first reaction zone and an oxygen-containing gas being introduced into the second and any subsequent reaction zones. A process as claimed in any of claims 1 to 8, characterized in that the third component is a co-formed component well distributed throughout the catalyst.

Inventors: IMAI, TAMOTSU, JAN, DEN-YANG
Applicants: UOP
Application year: 1988
Main classification: C07C5
IPC codes: C10G27, B01J23, C10G35, C07C11, C07C15, C07C5, C07B61