KR920010009B1 - Dehydrogenation catalyst composition - Google Patents

Abstract

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Description

Hydrocarbon Dehydrogenation Reaction Catalyst

1 and 2 are graphs showing the performance in the dehydrogenation process of the catalysts D, E and F prepared in Example 1, respectively.

1 is a graph showing the conversion (% by weight) as a function of test run time,

2 is a graph showing the selectivity (mol%) of the catalyst as a function of operating time for propylene production.

The present invention relates to a process for converting hydrocarbons in the presence of a catalyst complex, in particular to dehydrogenation of hydrocarbons which can be dehydrogenated. The present invention also relates to novel catalyst composites.

The dehydrogenation of hydrocarbons is of commercial importance because dehydrogenated hydrocarbons are increasingly demanded in the manufacture of various chemical products such as detergents, high octane gasoline, pharmaceuticals, plastics, synthetic rubbers and other products known in the art. It is a process. One example of this process is the dehydrogenation of isobutane for the production of isobutylene, which can be prepared by polymerizing a viscous agent for adhesives, a viscosity index additive for automotive oils and an antioxidant additive for plastics.

In the prior art, various catalyst complexes are known which contain a group VIII metal component, an alkali metal or alkaline earth metal component and a component selected from the group consisting of tin, germanium, lead, indium, gallium, thallium or mixtures thereof. However, catalyst composites containing these components complexed on a θ-alumina support having a surface area of 80-100 m 2 / g and an apparent volume density (ABD) of at least 0.5 g / cm 2 are not yet known in the art.

US Patent No. 4,070,413 discloses a dehydrogenation process using a catalyst in which a Group VIII metal and lithium are impregnated onto an alumina support. The alumina support is another feature that the hydrothermal treatment in steam at a temperature of about 800-1200 ℃. The catalyst of the present invention is characterized in that the catalyst of the present invention contains a component selected from the group consisting of tin, germanium, lead, indium, gallium, thallium or mixtures thereof in addition to the Group VIII metal component and the alkali metal or alkaline earth metal component. Is distinguished from. In addition, the catalyst support of the present invention has a larger ABD than the support of this patent. The patent discloses a catalyst having a pre-hydrothermally treated ABD of about 0.25-0.45 g / cm 3. From Example 3, it can be seen that the final catalyst complex of the catalyst of this patent has an ABD of about 0.3. The catalyst of the present invention should have a final ABD of at least 0.5 g / cm 3.

US Pat. No. 4,608,360 discloses a catalyst composite consisting of a Group VIII noble metal component, a co-formed Group IVA metal component and an alkali or alkaline earth metal component on an alumina support having a surface area of 5-150 m 2 / g. In addition, the alumina support of the patent is characterized in that the average pore diameter is about 300 mm 3 or less, and the pore having a diameter of 600 mm 3 or more occupies about 55% of the total pore volume of the support. In contrast, the catalyst of the present invention is characterized by a surface area of 80-100 m 2 / g and an ABD of at least 0.5 g / cm 3. The catalyst of this patent can be seen from Example 2 that the ABD is about 0.3. In addition, in the catalyst of the present invention, pores with a diameter of 600 mm 3 or more occupy a very small percentage of the total pore volume, while the catalyst of the patent accounts for pores having an average diameter of about 600 mm 3 or more occupy more than 50% of the total pore volume.

US Pat. No. 4,717,779 discloses a process for dehydrogenation of a dehydrogenable hydrocarbon using a selective oxidation catalyst consisting of a Group VIII precious metal component, a Group IVA component and, if necessary, a Group IA or Group IIA component. These components have an ABD of less than about 0.6 g / cm 3 of alumina precursor, but after calcining at a temperature of about 9000-1500 ° C., the ABD of alumina is 0.3-1.1 g / cm 3, and a pore of 1500 Å or more has 40% of the pore volume. It is complex-processed on the alumina support which occupies the above. In contrast, the catalyst of the present invention has an ABD of at least 0.5 g / cm 3, preferably at least 0.6 g / cm 3. In addition, pores of at least 1500 mm 3 make up a very small proportion of the total catalyst pore volume, ie less than 40% of the total catalyst pore volume.

The present invention relates to a catalyst component and a process using a catalyst having a surface area of 80-100 m 2 / g and an ABD of at least 0.5 g / cm 3. Anywhere in the prior art finds a substrate suggesting that such alumina catalyst substrates according to the invention have been used in the dehydrogenation of hydrocarbons which can be dehydrogenated, together with platinum group metal components, Group IVA metal components, and alkali or alkaline earth metal components. Can't.

It is an object of the present invention to provide an improved catalyst complex and a process for converting hydrocarbons using the catalyst complex, in particular a method for dehydrogenation of hydrocarbons which can be dehydrogenated.

Thus, in a broad embodiment, the present invention relates to tin, germanium, lead, indium, gallium, thallium or a second component selected from the group consisting of a first component, an alkali metal or an alkaline earth metal or a mixture thereof, which is a Group VIII precious metal It relates to a catalyst composite consisting of a combination of a third component selected from the group consisting of and a θ-alumina support having a surface area of 80-100 m 2 / g and an ABD of at least 0.5 g / cm 3.

In a more preferred embodiment, the catalyst composite of the present invention comprises a catalytically effective amount of a platinum and cesium and a third component selected from the group consisting of tin, germanium, lead, indium, gallium, thallium or mixtures thereof, with a surface area of 80 -100 m 2 / g and having an ABD of 0.6 g / cm 3 or more in combination with a θ-alumina support.

In another embodiment, the present invention is directed to a process for converting hydrocarbons using one of the catalysts described above. As a most preferred embodiment, the hydrocarbon conversion process is carried out under dehydrogenation conditions with a temperature of 400-900 ° C., a pressure of 0.1-10 atm, and an hourly liquid space velocity of 0.1-100 hr ⁻¹ .

An essential feature of the present invention lies in the nature of the support of the catalyst of the present invention. In particular, it is important that the surface area of the alumina catalyst support is 80-100 m 2 / g and the corresponding apparent volume density (ABD) is at least 0.5 g / cm 3. The support consists of a plurality of catalyst components, including Group VIII precious metal components, alkali metal or alkaline earth metal components and a component selected from tin, germanium, lead, indium, gallium, thallium or mixtures thereof. Such catalysts exhibit improved catalyst conversion and selectivity in hydrocarbon dehydrogenation processes as compared to conventional similar dehydrogenation catalysts. This improvement is believed to be due in particular to the increased pore size of the catalyst. Since the pores of the catalysts are larger than similar catalysts of the prior art, they are not easily coked, and the stability of the catalyst and access to the pores is improved.

As mentioned above, one of the essential features of the catalyst composite of the present invention lies in the first component selected from Group VIII precious metals or mixtures thereof. Group VIII precious metals may be selected from the group consisting of platinum, palladium, iridium, rhodium, osmium, ruthenium or mixtures thereof. However, platinum is a more preferred group VIII precious metal component. Virtually all Group VIII precious metal components are believed to exist in the elemental metal state in the catalyst.

Preferably, the Group VIII noble metal component is well dispersed throughout the catalyst. In general, Group VIII precious metal components, particularly platinum, are preferred and comprise about 0.01-5.0% by weight of the final catalyst composite, calculated on an elemental basis.

The Group VIII noble metal component may be incorporated into the catalyst composite by any suitable method, such as coprecipitation or cogelling, ion exchange or impregnation, or vapor deposition or deposition from an atomic source, or other catalyst components may be incorporated. It may be incorporated into the catalyst composite by a method similar to the above method before, during or after incorporation. A preferred method of incorporating the Group VIII precious metal component is a method in which the alumina support is impregnated with a solution or suspension of the decomposable compound of the Group VIII precious metal. For example, platinum may be added into the support by mixing the support with aqueous platinum chloride solution. In addition, for example, nitric acid or an acid of other optional components can be added to the impregnation solution to help the Group VIII noble metal component to be uniformly dispersed or fixed in the final catalyst composite.

Another essential component of the catalyst of the present invention is an alkali metal or alkaline earth metal component. The alkali metal or alkaline earth metal component of the present invention is a group consisting of cesium, rubidium, potassium, sodium and lithium or a group consisting of barium, strontium, calcium and magnesium or any one or both of these groups of metals. It can be chosen from mixtures. However, when only one component is selected for the catalyst composite of the present invention, the preferred second catalyst component is cesium. The alkali metal and alkaline earth metal components are present in the final catalyst composite in an oxidation state above that of the elemental metal. The alkali metal and alkaline earth metal components may be present as compounds such as oxides or may be present in admixture with the carrier material or other catalyst components, for example.

Preferably, the alkali metal and alkaline earth metal components are well dispersed throughout the catalyst composite. Alkali or alkaline earth metal components generally comprise about 0.01-10.0% by weight of the final catalyst composite, calculated on an elemental basis. Particularly preferably, the second catalyst component is cesium and constitutes about 0.01-4.0% by weight of the final composite, calculated on an elemental basis.

The alkali metal or alkaline earth metal component is incorporated into the catalyst composite by a suitable method such as co-precipitation or cogelling, ion exchange or impregnation, or before, during or after incorporation of other catalyst components. It can be incorporated into the catalyst composite by a similar method. A preferred method of incorporating the first and second alkali components is by impregnating the carrier material in the cesium nitrate solution.

The third essential component of the catalyst of the invention is a modifier metal component selected from the group consisting of tin, germanium, lead, indium, gallium, thallium and mixtures thereof. The third modifier component is preferably impregnated with an effective amount uniformly. In general, the catalyst, when calculated on an elemental basis, comprises about 0.01-10% by weight of the third modifier metal component of the final composite. Preferably, the catalyst comprises about 0.01-5.0 wt% of the third modifier metal component.

Any third modifier metal component of the present invention is preferably tin. All tin components exist in an oxidation state above the oxidation state of the elemental metal in the catalyst. This component may be present in the complex as a compound such as oxides, sulfides, halides, oxychlorides, aluminates, or in admixture with the carrier material or other components of the complex. Preferably, the tin component is used in an amount sufficient to contain about 0.01-10% by weight, most preferably about 0.01-5.0% by weight, of the final catalyst composite on an elemental basis.

Examples of suitable tin salts or water soluble tin compounds that can be used include tin bromide, tin tin chloride, tin chloride, tin chloride pentahydrate, ditin chloride tetrahydrate, ditin chloride trihydrate, tin chloride diamine, tin tribromide, Tin chromate, tin fluoride, ditin fluoride, tin iodide, tin sulfate, tin tartarate, and the like. Particular preference is given to tin chloride compounds such as ditin chloride or tintin chloride.

The third component of the catalyst may be combined with the support in any order. Thus, FIG. 1 may incorporate a second component onto a support and then form or evenly incorporate a surface with one or more third components. Alternatively, the third component may form a surface on the support or may be uniformly incorporated and then other catalyst components may be incorporated.

The catalyst composition of the present invention may also contain a halogen component. Halogen components include fluorine, chlorine, bromine, iodine or mixtures thereof. Chlorine is the preferred halogen component. The halogen component may generally be present in a mixed state with the porous carrier material and the alkaline component. Preferably, the halogen component is well dispersed throughout the catalyst composite. The halogen component can make up 0.01-15% by weight of the final catalyst composite, calculated on an elemental basis.

The halogen component can be incorporated into the catalyst composition in a suitable manner during the preparation of the carrier material, or prior to, during, or after incorporation of other catalyst components. For example, the alumina sol used to form the desired aluminum carrier material may contain halogen and thus provide at least a portion of the halogen content in the final catalyst composition. In addition, the halogen component, or part thereof, may be added to the catalyst component when incorporating another catalyst component into the carrier material, for example, by using platinum chloride to contain the platinum component. The halogen component or part thereof may also be added to the catalyst composite by contact with a halogen-containing compound or solution or halogen before or after the other catalyst component is incorporated into the carrier material. Suitable compounds containing halogen include acids containing halogen, such as hydrochloric acid. In addition, the halogen component or portions thereof can be incorporated by contacting the catalyst with a compound or solution containing halogen in the subsequent catalyst regeneration process. In the regeneration process, the carbon deposited on the catalyst as coke during the use of the catalyst is burned in a process that is converted into hydrocarbons, and the catalyst and the platinum group components on the catalyst are redistributed to produce a regeneration catalyst having the same performance characteristics as the new catalyst to provide. The halogen component can be added, for example, by contacting the catalyst with hydrogen chloride gas during the carbon combustion process or during the platinum group component redistribution process. The halogen component can also be added to the catalyst composite by adding a halogen or a halogen containing compound or solution, such as propylene dichloride, to the hydrocarbon feed stream or recycle gas during the hydrocarbon conversion process. Halogen may also be added as chlorine gas (Cl ₂ ).

The carrier material of the present invention is alumina having a surface area of 80-100 m 2 / g. In addition, alumina as a catalyst carrier has an ABD of 0.5 g / cm 3 or more. Alumina carrier materials can be prepared from natural or synthetic raw materials in a suitable manner. The carrier may be molded into a predetermined shape, for example, spherical, pill, cake, extruded, powder, granular, or the like, and may be used having any particle size. The preferred shape of the alumina is spherical. Preferred particle sizes are about 1.59 mm (about 1/16 inch) in diameter, but small particles of about 0.79 mm (about 1/32 inch) or less can also be used.

To make the alumina spherical, the aluminum metal is reacted with a suitable peptising agent and water to convert it into an alumina sol, which is then dropped in an oil bath to form spherical particles of an alumina gel. According to another aspect of the present invention, the third modifier metal component can be added to the alumina sol before reacting with the peptizing agent and dropping it in a hot oil bath. In addition, other shapes of alumina carrier material can also be prepared by conventional methods. The alumina particles optionally containing the coformed third component are shaped, then dried and calcined.

The most important in imparting the preferred properties of the invention to the catalyst substrate is the drying and calcining of the alumina based component. It is important that the alumina based catalyst of the present invention has a surface area of 80-100 m 2 / g and a corresponding ABD of at least 0.50 g / cm 3. These properties are obtained by the final firing of alumina in the temperature range of 950-1200 ° C. The final firing process is preferably carried out under conditions sufficient to convert the alumina into θ-alumina suitable for the desired properties of the alumina substrate of the catalyst of the invention. Such conditions include a calcination temperature strictly controlled in the range of 950-1100 ° C, preferably 975-1020 ° C.

The surface area of the catalyst described in the description of the invention and in the claims was measured by known mercury intrusion techniques. This method can be used to measure the porosity distribution and pore surface area of porous materials by mercury instruc- tion using the Micromeritics Auto Pore 9200 Analyzer. In this method, high pressure mercury is injected into the pores of the catalyst particles, increasing the pressure up to 413,700 kPa (60,000 psia). Record the void volume at the predetermined pressure. Up to 85 pressure points can be selected. Thus, in this way it is possible to determine the overall distribution of the void volume.

The baking effect of the aluminum substrate at high temperatures described in the present invention is that the densification of the alumina substrate can be obtained. Densification, ie an increase in ABD, is achieved by reducing the total pore volume of the catalyst. In addition, high firing temperatures expand the voids present. To achieve this contrasting mechanism, the catalyst must be shrunk to the appropriate size while the existing pores expand. As the pores expand, the openings of the existing pores become larger, thus lessening the tendency for the pores to be filled or blocked due to the build up of coke.

It is preferred that the alumina component is essentially θ-alumina. The term “essentially θ-alumina” means that at least 75% of the alumina microcrystals are θ-alumina microcrystals. The remaining alumina microcrystals may be of the α-alumina or r-alumina type. However, other forms of alumina microcrystals known in the art can also be used. Most preferably, the "essentially θ-alumina component" is at least 90% of θ-alumina microcrystals.

As noted above, the θ-alumina form in crystalline alumina is prepared from an amorphous alumina precursor by tightly controlling the maximum firing temperature experienced by the catalyst support. It is known that firing temperatures 800-950 ° C. produce alumina which essentially consists of gamma-alumina microcrystals. Firing temperatures of 1100 ° C. and higher are known to promote the formation of alpha-alumina microcrystals, whereas firing temperatures of 950-1100 ° C., especially 975-1020 ° C., promote the formation of θ-alumina microcrystals.

After mixing the catalyst components with the desired alumina support, the resulting catalyst composite is dried at a temperature of about 100-320 ° C., typically for about 1-24 hours or longer, and then at about 320-600 ° C. at about 0.5 Fire for -10 hours or longer. This final firing process typically does not adversely affect alumina microcrystals or ABD. However, if necessary, the support may be calcined at high temperature. Finally, calcined catalyst composites are typically subjected to a reduction step prior to use in hydrocarbon conversion processes. This reduction step is carried out in a reducing atmosphere, preferably in a dry hydrogen atmosphere, at about 230-650 ° C. for about 0.5-10 hours or longer, at which time the temperature and time are such that almost all platinum group components are reduced to the elemental metal state. Choose enough to.

As mentioned above, the catalyst of the present invention has particular utility as a hydrocarbon conversion catalyst. The hydrocarbon to be converted is brought into contact with the catalyst at hydrocarbon conversion conditions. Such conditions include a temperature of about 200-1000 ° C., a pressure of 0.25 absolute air pressure (ATMA) to about 25 gauge air pressure, an hourly liquid space velocity of about 0.1 to about 200 hr ⁻¹ , and the like.

According to an embodiment of the present invention, the hydrocarbon conversion process is a dehydrogenation reaction. In a preferred process, the hydrocarbon which is dehydrogenated in the dehydrogenation zone maintained under dehydrogenation conditions is brought into contact with the catalyst composite of the present invention. Such a contact process can be performed by a fixed catalyst bed system, a mobile catalyst bed system, a fluidized bed system, etc., or by a batch operation. Of these, the fixed catalyst layer system is better. In such a fixed catalyst bed system, the hydrocarbon feed stream is preheated to a predetermined reaction temperature and then passed through a dehydrogenation zone with the fixed catalyst bed. The dehydrogenation zone itself has one or more separate reaction zones, and it is desirable to provide heating devices between these reaction zones to maintain the desired reaction temperature at the inlet of each reaction zone. The hydrocarbon may be contacted with the catalyst in either upstream, downstream, or radial flow. Hydrocarbon radial flow through the catalyst bed is preferred for commercial scale reactors. Hydrocarbons may be in liquid, vapor-liquid mixed or vapor phase upon contact with the catalyst, with hydrocarbons in vapor phase being preferred.

Hydrocarbons that can be dehydrogenated include C2-C30 or more dehydrogenated hydrocarbons including paraffins, alkylaromatics, naphthenes and olefins. One group of hydrocarbons that can be dehydrogenated using the catalyst of the present invention is a common paraffin group having 2 to 30 carbon atoms or more. The catalyst is particularly useful for dehydrogenating paraffins having 2 to 15 carbon atoms or more to their corresponding monoolefins or dehydrogenating monoolefins having 3 to 15 carbon atoms or more to their corresponding diolefins. The catalyst is particularly useful for dehydrogenating C ₂ -C ₆ paraffins, mainly propane and butane, to monoolefins.

Dehydrogenation conditions include a temperature of about 400-900 ° C., a pressure of about 0.01-10 absolute atmosphere and a liquid hourly space velocity (LHSV) of about 0.1-100 hr ⁻¹ . In general, for ordinary paraffins, the smaller the molecular weight, the higher the temperature required for equivalent conversion. The pressure in the dehydrogenation zone is kept as low as practically possible, depending on the limitations of the device, in order to maximize the benefits of chemical equilibrium.

Effluent streams from the dehydrogenation zone generally contain unconverted dehydrogenable hydrocarbons, hydrogen and dehydrogenation reaction products. This effluent stream is typically cooled to separate the hydrogen rich vapor phase from the hydrocarbon rich liquid phase by passing through a hydrogen separation zone. Generally, the hydrocarbon-rich liquid phase is further separated by suitable selective absorbents, optional solvents, selective reaction (s) or using suitable fractionation apparatus. Unconverted dehydrogenable hydrocarbons can be recovered and recycled to the dehydrogenation zone. The product of the dehydrogenation reaction is recovered as the minimum product or as an intermediate in the preparation of other compounds.

The dehydrogenable hydrocarbon may be mixed with the diluent before, during or after passing through the dehydrogenation zone. The diluent may be hydrogen, water vapor, methane, ethane, carbon dioxide, nitrogen, argon, or the like or a mixture thereof. Hydrogen and water vapor are preferred diluents. Typically, when hydrogen or water vapor is used as the diluent, the amount is used in an amount sufficient so that the molar ratio of diluent to hydrocarbon is from about 0.1: 1 to about 40: 1, most at a molar ratio of about 1: 1 to about 10: 1 Excellent results are obtained. The diluent stream passed to the dehydrogenation zone may be a recycle diluent which is usually separated from the effluent from the dehydrogenation zone in the separation zone.

Diluent mixtures such as water vapor and hydrogen can be used. When hydrogen is the main diluent, substances that decompose under water or dehydrogenation conditions to form water, such as alcohols, aldehydes, ethers or ketones, can be added continuously or intermittently to the dehydrogenation zone, the amount being approximately in the hydrocarbon feed stream. 1 to about 20,000 ppm by weight (calculated based on equivalent water). When dehydrogenated paraffins dehydrogenate paraffins having 6 to 30 carbon atoms or more, the best results are obtained when the added water is about 1-10,000 ppm by weight.

To be commercially successful, dehydrogenation removal catalysts must have three properties: high activity, high selectivity and good stability. Activity is a measure of the ability of a catalyst to convert a reactant to a product under certain reaction conditions, i. For the activity of the dehydrogenation catalyst, the percentage of paraffin conversion or dissipation relative to the amount of paraffin in the raw material is measured. Selectivity is a measure of the catalytic ability of the catalyst to convert the reactant to the desired reactant relative to the amount of reactant converted. Catalyst selectivity is determined in mole percent by measuring the amount of olefins in the product relative to the total moles of converted paraffins. Stability is a measure of the rate of change of activity and selectivity parameters over time of operation, the lower this ratio the more stable the catalyst.

Dehydrogenation of hydrocarbons is an endothermic process. In systems using only dehydrogenation catalysts, it is usually necessary to add superheated steam at various points in the process or to intermittently remove and reheat the reaction stream between catalyst beds. As a further improved method, a method of using a bicatalyst system with a special bed or a dehydrogenation reactor or a selective oxidation catalyst has been developed.

The purpose of the selective oxidation catalyst is to selectively oxidize the hydrogen produced as a result of the dehydrogenation reaction with oxygen added to the oxidation zone to generate heat within the process itself. The heat generated is typically sufficient to allow the reaction mixture to reach the desired dehydrogenation temperature for subsequent dehydrogenation processes. This method may be performed with the system described above. When using this method, the catalyst of the present invention will consist of at least the dehydrogenation catalyst and other specific catalysts used to carry out the oxidation reaction. Prior to describing the preferred reactor configuration, the oxidation of the present invention is described in detail.

The selective oxidation step, when used, uses the hydrogen generated in the dehydrogenation step of the process to supply heat to the next dehydrogenation reaction step. To this end, an oxygen containing gas is first introduced into the reactor, preferably at a point proximate to the selective oxidation catalyst section. Oxygen in the oxygen containing gas is needed to oxidize the hydrogen contained in the reaction stream. Examples of oxygen-containing gases that can be used to perform the selective oxidation of hydrogen present are air, oxygen, or air or oxygen diluted with other gases such as water vapor, carbon dioxide and inert gases (e.g. nitrogen, argon, helium, etc.). Can be mentioned. The amount of oxygen introduced to contact this process stream is preferably about 0.01-2 moles of oxygen per mole of hydrogen included in the process stream at the time oxygen is added to the process stream. In the selective oxidation reaction, a process stream consisting of unreacted dehydrogenable hydrocarbons, dehydrogenated hydrocarbons and hydrogen is reacted with oxygen in the presence of a selective steam oxidation / dehydrogenation catalyst to react with hydrocarbons by selectively oxidizing hydrogen. It generates water and thermal energy with traces of oxygen.

Selective steam oxidation / dehydrogenation catalysts are useful for the selective oxidation of hydrogen in the presence of hydrocarbons. Examples of such catalysts are described in US Pat. No. 4,418,237. Alternatively, the catalyst used in the selective oxidation step may be the same as the catalyst used in the dehydrogenation step. Such catalysts or methods of using them are described in US Pat. Nos. 4,613,715 and 3,670,044. The catalyst of the present invention exhibits both dehydrogenation and selective oxidation functions. Thus, the catalyst of the present invention can be used for the dehydrogenation and selective oxidation process of hydrocarbons containing a single catalyst.

The oxygen-containing reactant may be added to the process in various ways, such as by mixing oxygen with a relatively low temperature hydrocarbon feed stream or steam diluent, or may be added directly to the reactor independent of the hydrocarbon feed or steam diluent. In addition, oxygen-containing reactants may be added to the reactor at various locations in such a way as to minimize the local concentration of oxygen relative to hydrogen, in order to distribute the beneficial temperature rise caused by selective hydrogen oxidation over the entire length of the reaction zone. . In practice, it is a preferred method of operation to use multiple injection points for introducing an oxygen containing gas into the steam oxidation / dehydrogenation reaction zone. This method minimizes the possibility of local buildup of oxygen concentration over the amount of hydrogen, thereby minimizing the opportunity for undesirable reactions of the oxygen-containing gas with the raw materials or the hydrocarbons produced.

The following examples further illustrate the catalyst and process of the present invention. These examples are provided for purposes of illustration and should not be considered as limiting the comprehensive interpretation of the invention as set forth in the appended claims.

Example 1

In order to demonstrate the advantages achievable by the present invention, the catalyst of the invention and a number of catalysts different from the invention have been prepared. First, spherical alumina supports for all catalysts were prepared by well known hydrophobic methods. The tin component was incorporated into the support by mixing the tin component precursor with the alumina hydrosol and then gelling the hydrosol. In this case, the tin component was evenly distributed throughout the catalyst particles. The catalyst particles were then dried at 600 ° C. for about 2 hours and then calcined at various temperatures as shown in Table 1 below. The firing temperature shown in Table 1 is the maximum firing temperature used for each catalyst.

The calcined tin-containing particles were then impregnated with the chloroplatinic acid solution and cesium nitrate solution to homogenously incorporate platinum and cesium into the alumina substrate. After impregnation, the catalyst was dried in an oven at about 150 ° C. for 2 hours in the presence of 10% water vapor and then heated at 540 ° C. for 1/2 hour in the absence of water vapor.

Table 1 below shows the metal content, surface area and physical properties such as ABD of each of the resulting catalysts.

TABLE 1

Catalysts A, E and F are not catalysts according to the present invention. Catalysts B, C and D are all catalysts prepared according to the present invention.

Example 2

Catalysts A and B were evaluated for their ability to dehydrogenate propane feedstock in the experimental plant. The experimental plant was operated at an inlet temperature of 600 ° C., 1 atmosphere and a liquid hourly space velocity of 3 hr ⁻¹ . Water and propane were simultaneously introduced into the reactor of the experimental plant with a molar ratio of H ₂ O / C ₃ as 2. Data on C ₃ conversion and selectivity for both catalysts are listed in Table 2.

TABLE 2

It is clear from Table 2 that catalyst B, the catalyst of the present invention, exhibited higher selectivity and conversion than catalyst A, whereby catalyst B of the present invention is far superior to catalyst A of the prior art in its ability to dehydrogenate C ₃ hydrocarbons. Done In addition, the catalyst B was found to be θ-alumina by analysis by X-ray diffraction technique.

Example 3

In this example, the effect of cesium concentration on the catalyst of the present invention was tested. In addition to the very different surface areas, the catalysts in Example 2 also differed in cesium concentrations. In this example, it is intended that the concentration of cesium has only a very small effect on the performance of the catalyst.

In this example, the ability to dehydrogenate the paraffins of the mixed paraffin / olefin raw materials in the experimental plant was evaluated for both C and D, the catalysts of the present invention. Catalysts C and D both contained 0.75 wt% platinum and 0.5 wt% tin on essentially the same support. However, catalyst C contained 3.5 weight percent cesium, while catalyst D contained 4.0 weight percent cesium.

Both catalysts were evaluated equally in both reactors in the same experimental plant. The raw materials supplied to the experimental plant consisted of 0.3 mol of propylene, 0.7 mol of propane, 2.07 mol of water, 0.3 mol of hydrogen, and 0.13 mol of nitrogen. Both reactors were operated at an inlet temperature of 600 ° C. The pressure in the experimental plant was adjusted so that the outlet pressure of the second reactor remained 1.34 atm. The hourly liquid space velocity of the first reactor based on the hydrogen feed rate was 80 hr ⁻¹ . The space velocity of the second reactor was 8 ^{hr −1} . The evaluation results are shown in Table 3 below.

TABLE 3

From Table 3, it can be seen that the two catalysts show similar conversion and selection performance. Clearly, this performance is not equivalent and cesium concentration has some effect. It can be seen that the activity reduction rates of the two catalysts are very similar. This is in contrast to the two catalysts of Example 2, where the prior art catalysts have reduced activity much more rapidly than the catalysts of the present invention. That is, catalysts with a large surface area of the support tend to clog their small pore inlets, whereas catalysts with a small surface area of the present invention do not exhibit such rapid decreases in activity.

Example 4

In this example, the catalysts D, E and F prepared in Example 1 were evaluated in the same experimental plant as described in Example 3. The purpose of this test is to evaluate the difference in the performance of the dehydrogenation catalyst due to the change in the surface area of the three catalysts. The investigation showed that the catalyst D had a surface area of 80 m 2 / g, the catalyst E was 107 m 2 / g, and the catalyst F was 45 m 2 / g. All three catalysts had an ABD of 0.5 g / cm 3 or more. The results of the activity and selectivity tests of the second reactor are illustrated in FIGS. 1 and 2.

1 shows the C ₃ molar conversion of each catalyst in a second reactor of two reactors as a function over time. FIG. 2 also shows the propylene selectivity of each catalyst in the second reactor in mol% as a function over time.

These figures show that catalyst D with a surface area of 80 m 2 / g also shows better conversion and selectivity than catalysts E and F which are not catalysts of the invention. Catalyst F with a surface area of 45 m 2 / g showed much lower conversion and selectivity compared to Catalysts D and E. Catalysts D and E showed similar C ₃ conversion performances, and the conversion stability of catalyst D was somewhat better than catalyst E. Catalyst D clearly showed better propylene selectivity than catalyst E.

The results show maximum propane conversion and propylene selectivity at a surface area of about 80 m 2 / g and show that the conversion and selectivity is reduced when the surface area approaches about 45 m 2 / g and 107 m 2 / g.

Claims (4)

0.01-5.0% by weight of platinum, 0.01-4.0% by weight of cesium and 0.01-5.0% by weight of tin on a θ-alumina support having a surface area of 80-100 m 2 / g and an apparent volume density (ABD) of not less than 0.5 g / cm 3 Catalyst composite, characterized in that. A dehydrogenated hydrocarbon is contacted with a catalyst according to claim ¹ at a temperature of 400-900 ° C., a pressure of 0.1-10 atm and a liquid hourly space velocity (LHSV) of 0.1-100 hr ⁻¹ . Method of dehydrogenation. The method of claim 2 wherein the dehydrogenable hydrocarbon is a dehydrogenable C ₂ -C ₃₀ hydrocarbon. 4. The process of claim 3 wherein the dehydrogenable hydrocarbon is C ₂ -C ₆ paraffins.

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