The method of obtaining saturated oligomers and method for producing motor fuel

 

(57) Abstract:

The method of obtaining saturated oligomers is in the oligomerization of light olefins to heavier olefins and heavy saturation of olefins. In doing so, the recycling of heavy paraffins in the oligomerization zone. The method can improve the selectivity of the oligomerization relative to C8-products and to reduce contamination of the catalyst. As a result, the oligomerization zone operates at a lower pressure. 2 C. and 8 C.p. f-crystals, 1 Il., table 2.

The use of the invention

The invention relates primarily to the receipt of hydrocarbons boiling within the gasoline by the oligomerization of light olefins and saturation of the resulting heavier olefins.

The background to the invention

There is a continuing need in the conversion of light hydrocarbon compounds into high-octane motor fuel. The alkylation of isobutane with propylene, butylenes and amylene using hydrofluoric acid (HF) as a catalyst, often referred to as HF-alkylation was highly effective way to produce high octane motor Topley is a mini selection of hydrofluoric acid (HF) of facilities for the HF-alkylation, led to research on the improvement or creation of alternative HF alkylation processes for the production of motor fuels. One existing alternative is similar to the process of alkylation, in which the catalyst used sulfuric acid. Although the use of sulfuric acid may reduce the degree of hazard, which to some extent connected with the use of hydrofluoric acid (HF), a process using sulfuric acid still perceived as representing the ability of the same dangers, and it not so cost-effective, as the process HF-alkylation. Thus, the search continues processes, replacing HF-alkylation.

Known and practiced other methods of connection of isobutane with light olefins for the production of motor fuels, but they do not provide a gasoline product of similar quality, or more capital-intensive in the creation and operation. One such alternative method is to dehydrogenate isobutane and oligomerization of the resulting olefins in the production of hydrocarbons within the gasoline boiling. Oligomerization of light olefins in motor epimerase before process HF-alkylation (see U.S. patent USA-2526966). Such oligomerization was also called catalytic condensation and polymerization, and the resulting motor fuel was often referred to as polymer gasoline.

Also known as process hydrogen olefinic hydrocarbon streams obtained by oligomerization, to saturate olefins. In British patent GB-A-2186287 described method of dehydrogenization and oligomerization fractions containing 4 carbon atoms, (C4), in the production of fuel for jet aircraft, which may hydrogenizing into high-quality gasoline. The method of processing the hydrogen jet fuels, diesel fuels and lubricating oils, the resulting dehydrogenization and oligomerization of light paraffins, disclosed in U.S. patent US-A-4678645. However, the treatment with hydrogen is not always favorable for gasoline fractions obtained by oligomerization, and can lead to a lower octane number, but it is known that particularly beneficial when saturated isooctane to isooctanol.

Summary of invention

The aim of the present invention is to improve integrated action oligomerization oligomerization zone and the zone of saturation, that in combination with the area of dehydrogenization can provide an attractive alternative process HF-alkylation.

The present invention combines the oligomerization of light olefins with the saturation of the resulting oligomers by re-loop part of the saturated oligomers produced as recycling in the oligomerization zone. Recirculation saturated oligomers in the oligomerization zone prevents contamination of the catalyst in the oligomerization zone in the deposition of carbon and unexpectedly also provides improved selectivity of the oligomerization zone in relation to octane WITH8-isomers of a higher order. Reduced pollution, provided by recirculation of heavy paraffins from the saturated exhaust zone oligomerization can be used to extend the life of the catalyst in the oligomerization zone or work zone oligomerization at a lower pressure.

Operation at lower pressures is particularly advantageous when the oligomerization zone and the zone of saturation combined with the area of dehydrogenization to create a power source light olefins. Usually the process of dehydrogenization at low pressure bn operating pressure in the dehydrogenation reactor may be 136-1136 kPa, while working pressures in a typical area of oligomerization usually exceed 2170 kPa and often are at the level of about 3550 kPa. By recirculation of saturated oligomers to reduce contamination of the oligomerization zone pressure can be reduced gradually below 2170 kPa. The difference in operating pressure between the area of dehydrogenization and oligomerization zone may be reduced to less than 1034 kPa, and it eliminates the need for any two-stage compression between these zones. Thus, the device is integrated zone dehydrogenization, oligomerization and saturation of benefits from this recycling both from the point of view of improved selectivity and reduce the cost of production.

In accordance with this, in a generalized form, the present invention is a process for the production of saturated oligomers. The process passes through the feeding site of the oligomerization zone that contains3-C5-olefins, oligomerization conditions in contact with an oligomerization catalyst at oligomerization conditions, usually characterized by the temperature 93-260oWith pressure 790-6996 kPa and hourly volumetric flow rate of 0,5 - 5,0 h-1. Recycle stream containing C8pair the Otok, contains7-olefins and heavier olefins and paraffins, which return from the oligomerization zone. At least part of the exit stream from the oligomerization zone and a stream containing hydrogen, is injected into the zone of saturation, and they are in contact with the catalyst saturation under conditions of saturation, in order to saturate the olefins in the output stream oligomerization. At least part of the exit stream from the zone of saturation, containing hydrocarbons having at least 8 carbon atoms, arrives in the area of oligomerization as mentioned recirculation flow, at the same time as part of the exit stream from the zone of saturation, WITH8and heavier paraffins, send back.

In a specific embodiment, the present invention proposed a method of producing motor fuel from the incoming stream containing isobutane. During the execution of the process input stream containing isobutane, served in the area of dehydrogenization and in the area of dehydrogenization the incoming stream is introduced into contact with the catalyst dehydrogenization in terms of dehydrogenization, usually characterized by the temperature 510-649oWith pressure 136-1136 kPa and volume hourly flow rate of 0.5 to 50.0 h-14-olefins. In the oligomerization zone food zone oligomerization and recycle stream containing paraffins, enter in contact with an oligomerization catalyst in the solid phase under the conditions of the oligomerization, usually at a temperature of 93-260oWith the pressure 790-6996 kPa and volume hourly flow rate 0,5-5,0 h-1effective to obtain the exit stream oligomerization, contains7-isoolefine and above. The exit stream from the oligomerization zone is served without separation in the saturation region, together with a stream of hydrogen. In the saturation zone is introduced into the contact output stream from the oligomerization zone and hydrogen with a catalyst saturation under conditions of saturation to saturation of the olefins in the output stream from the oligomerization zone. At least part of the exit stream from the zone of saturation is sent to a separation zone for regeneration of the saturated stream that contains8-paraffins and higher. At least part of the regenerated saturated flow is sent to the UE the lustration

The figure schematically depicts a process in accordance with the present invention and shows the principal technological zones and associated piping and equipment.

Detailed description of the invention

The main working area for the implementation of the present invention is the reaction zone oligomerization. Corresponding zones of oligomerization in the present invention is made in various forms. Oligomerization has several names, such as catalytic condensation and catalytic polymerization. Among the known catalysts for such reactions include phosphoric acid in the solid phase, known under the name SPA-catalysts and homogeneous catalysts, such as boron TRIFLUORIDE, are described in U.S. patent US-A-3906053, USA-A-3916019 and US-A-3981941.

The preferred catalyst for the oligomerization process is phosphoric acid in the solid phase - SPA-catalyst. SPA catalyst, as mentioned earlier, refers to solid catalysts, which contain as a main ingredient acid on the basis of phosphorus such as ortho-, pyro - or tetraphosphorus acid. The catalyst usually form puta can be calcined and then ground to obtain catalyst particles, moreover, the paste can be ekstradiroval or it can be processed into granules prior to calcination to obtain a more uniform catalyst particles. As the carrier preferably use natural porous siliceous material such as diatomaceous earth, kaolin, infusorial earth and diatomaceous earth. The minimum number of different additives, such as mineral talc, mullerova earth and compounds of iron, including iron oxide, may be added to the media in order to increase its strength and rigidity. The combination of media and additives is preferably 15-30% of the catalyst, and the rest is phosphoric acid. Supplements can be 3-20% of the total mass of material media. There are, however, variants of this composition, such as a lower content of phosphoric acid. Further details relating to the composition and manufacture of the SPA catalysts can be obtained from U.S. patent US-A-3050473 and US-A-3132109.

In the present invention the reaction zone oligomerization preferably operates at temperatures and pressures that increase the compatibility of the conditions of her exit stream with the input conditions located downstream of the entrance into the reaction zone of tokenization. The preferred temperature in the reaction zone of oligomerization is usually located within 93-260oWith, and more typically within 149-232oS, and especially preferably for some catalysts, 149-204oC. the pressure in the reaction zone of oligomerization is usually located within 790-6990 kPa, and more typically within 790-3549 kPa, when using the present invention to reduce the operating pressure in the oligomerization zone. When applying the present invention particularly preferred working pressure range for PSA-catalyst is 790-2170 kPa. It was also found that the temperature in the narrow limits 149-171oWith may selectively cause more8-isomers.

Feeding a stream of the reaction zone oligomerization typically contains3-C5the olefins and paraffins. In the reactor can be fed with steam or water to maintain a low content of water of hydration preferred SPA catalyst. The supply of olefins is usually a light gas stream that is returned from the gas separation section of the FCC process, the stream WITH a4from the steam cracking of gas and gas coking or from the exit stream from dehydrogenization. Supply popitaiusi stream of olefins contains4-olefins, but it can also be completely or contain a large number of3-olefins. Usually the power olefins may contain olefins 3-C5the concentration which is at least 30 wt.%. Preferred food may contain4-olefins, the concentration of which is at least 30 wt. %, and more preferably at least 50 wt.%. It is preferable that the supply flow of the olefins contained 20 wt.%, and more preferably 30 wt.% isobutene. Isobutan preferably may be at least 33% of all of the butenes. Olefins in the preferred embodiment, the supply flow mainly includes branched olefins with isobutane present in large quantities. According to the invention the reaction of normal pentanol and propylene is carried out by maintaining a high concentration of isobutene in the supply flow of the oligomerization zone. Oligomerization of pentene and propylene in high-octane isomers promotes the distribution of the olefins in the feed stream to the isomerization zone that contains at least 50 wt. % isobutene. When propylene is present in large quantities in the supply flow zone oligomerization, Oct is the power. When the area of propylene oligomerization is injected in large quantities, it is preferable that butenova fraction was 100% from isobutene.

In the practical implementation of the present invention heavy paraffinic components in contact with the catalyst in conjunction with conventional feed stream oligomerization zone. Heavy paraffin components contain hydrocarbons, comprising in their structure at least eight carbon atoms, and the number of carbon atoms can reach up to 20, and preferably they contain C8-C12-paraffins. Add heavy paraffins provides a significant amount of heavy paraffins in the oligomerization zone, and preferably generates at least 20 wt.% WITH8-paraffins and higher in output stream of the reactor, and a more typical is the formation of at least 25 wt.% WITH8-paraffins and higher at the inlet of each catalyst layer. C8-paraffins are particularly preferred, and they preferably comprise at least 5 wt.% the first catalyst layer, and their content can reach up to 50 wt.% in the mass flow through the reaction zone oligomerization.

Heavy paraffin components can be introduced into protein places. In one or more cylindrical, vertically oriented vessels preferably may be several different layers of catalysts, and the supply flow is preferable to introduce into the reactor from above. The catalyst preferably is placed in a stationary layers in the oligomerization zone in the device, which is known as the structure of the reactor chamber type. Typically, the reactor chamber contains about five layers of catalyst. In the reactor chamber type reagents pass through several layers of catalyst of large diameter. The temperature of the reactants can further be controlled by recycling an additional stream of relatively inert hydrocarbons, which acts as a flow of heat. The reaction zone oligomerization, as a rule, are equipped with so many layers of catalyst, which take an intermediate injection of the cooling material to control the temperature of the exothermic reaction. Significant advantages can be obtained, if add power heavy paraffins in the form of intermediate injected flows, which are also favorable for the process, so as to serve as a cooling stream.

With the addition of heavy paraffins in the Combi is the acidity of the paraffin in the feeding zone of the oligomerization reaction is, at least 50 wt.% and more typically at least 70 wt.%. A large portion of the olefins in the feed stream is reacted in the reaction zone oligomerization with Isobutanol. The conversion of olefins is usually 80-99%. The main products of oligomerization consist of a7+-olefins.

The output stream after oligomerization, containing unreacted heavy olefins and received heavy olefinic hydrocarbons, is sent to the reactor saturation. The corresponding reactors provide almost complete saturation the saturation of all olefins in the reactor saturation. The exit stream from the oligomerization is preferably sent directly to the zone of saturation without separation or recirculation of light residues. Work at a lower pressure reactor oligomerization allows direct transfer of the exit stream polymerization in the hydrogenation reactor. Esotericist is usually the reason why the saturation zone operates at higher temperatures than the oligomerization zone, so that the coolant and paraffins in the output stream from the oligomerization provide additional material for the heat flow to the heat absorption of the reaction zone of saturation.

Bereavem thread. The gas stream should contain at least 50 wt.% of hydrogen. Preferably, the hydrogen-containing gas stream has a concentration of hydrogen greater than 75 wt.%. The hydrogen can be recycled from the section dehydrogenization to serve a greater number of hydrogen power in the saturation region with the rest of the necessary hydrogen supplied from external sources, as additional hydrogen, or in cases where zone dehydrogenization no, all of the necessary amount of hydrogen can be submitted from external sources. Preferably, the additional hydrogen possessed high purity in order to increase the purity of the hydrogen supplied to the zone of saturation, thereby reducing the volume of light hydrocarbons. These light hydrocarbons are generally undesirable, as their presence increases, without having this, a massive amount in the reaction zone of saturation, and their relatively high vapor pressure can lead to increased losses of potential hydrocarbons in the discharge downstream. However, a preferred variant of the present invention, which integrates the processes of dehydrogenization, oligomerization and saturation, can facilitate the use of Sep'a is the gas flow is mixed with the output stream of oligomerization in the ratios which provide ratios of hydrogen to hydrocarbon in the range of 0.1:to 2.0, and more preferably in the range of 0.15:0,30.

The preferred reactor saturating the present invention provides substantially complete saturation of all unsaturated hydrocarbons. Conditions in the hydrogenation zone are usually characterized by the temperature within 93-316oWith pressure 791-4928 kPa and volume hourly flow rate of 1-20 h-1. Preferably, the reaction conditions were chosen so as to keep the feed stream of hydrocarbons in the gas phase. The processing unit hydrogen typically operates at temperatures making it possible to increase the temperature of the combined feed stream to the reaction temperature of saturation due to heat exchange with the outgoing flow of the processing device by hydrogen. Thus, any heat input into a sequence of oligomerization and hydrogenation can preferably be carried out with the help of gravity of the heater on the input stream into the reaction zone oligomerization.

The preferred reactor for processing a hydrogen contains a fixed bed of the catalyst for handling hydrogen. The composition of the catalyst, which can be a combination of clay and aluminum-containing metal elements of group VIII and group VI(b) of the periodic table suitable for use. The group VIII elements include iron, cobalt, Nickel, ruthenium, rhenium, palladium, osmium, indium and platinum, and cobalt and Nickel are particularly preferred. Group VI(b) metal includes chromium, molybdenum and tungsten, and molybdenum and tungsten are particularly preferred. Metal components are placed on a porous support material. The carrier material may contain alumina, clay or silica. Particularly suitable catalysts are those containing a combination of cobalt and Nickel, comprising about 2-5 wt.%, and molybdenum, comprising 5-15 wt.% on an alumina carrier. Percentage (mass) of the metal content is calculated as if they are in the form of metals. Typical commercially supplied the catalyst is a composition based on alumina in the form of spheres or extruded particles impregnated with cobalt and molybdenum or Nickel and molybdenum in the proportions above. Other applicable compositions of catalysts containing 15-20 wt.% Nickel on alumina. The specific gravity of a commercially supplied catalysts typically 0.5-0.9 g/cm3while the surface area of 150-250 m2/, Usually, the higher the metal content in kata the button to obtain a recycle stream of heavy paraffins. Components of heavy waxes also contain hydrocarbons obtained. Department of the recirculation flow can be effected in a simple separator, which is a rough division of the exhaust flow from the zone of saturation, to ensure a sufficient number of heavy material recirculation flow oligomerization zone and a certain amount of cooling material that circulates in the saturation zone. Alternatively, the recycle stream may contain side stream product after fractionation of the lighter substances in the main separator.

Preferably, the exit stream from the reaction zone of saturation was filed in the cooling separator. In the cooling separator separates the most part WITH4-hydrocarbons and substances with a low boiling point from the exit stream of the zone of saturation to provide a cooling flow with a relatively high concentration of C8-hydrocarbons and higher. Regeneration of substances with a high molecular weight of the effluent stream of the reaction zone of saturation is feasible by combining the reaction zone oligomerization and the reaction zone of saturation. Regeneration C8-hydrocarbons and higher allows FDI oligomerization at a lower pressure.

The effect of oligomerization zone at a lower pressure, which can be achieved through the use of the present invention can also be applied to successfully merge the upper flow zone oligomerization zone of dehydrogenization. Usually the power zone dehydrogenization contains threads of light paraffins. Preferred foods rich WITH4-paraffins and contain large amounts of isobutane. (The term "rich" in this context means the thread in which the percentage (mass or volume) of the mentioned components is at least 50%, while the term "relatively wealthy" means a stream with a higher concentration of the mentioned components than power, from which it is produced. ) Food zone dehydrogenization preferably has a concentration of isobutane in the range 60-99 wt. %. Typical sources of this feed stream is a natural butane, streams, saturated WITH4hydrocarbons, oil refineries and butane getregistration installations. The flow Isobutanol can be obtained from the flow of butane oil refineries or from other sources where they get rich in Bhutan power. Prepost denaut zone dehydrogenization with the reaction zone oligomerization by using high pressure and a low degree of conversion. Low degree of transformation allows to reduce the deactivation of the catalyst and particularly in combination with higher bulk velocity provides the opportunity for most areas of the hydrogenation reaction to work at low regeneration requirements. In addition, higher pressures, especially to reduce requirements for compression to separate the exit stream and increase the efficiency of the process.

The process of catalytic dehydrogenization and the device of the reactor are well known. In the respective processes of dehydrogenization produce mixing raw materials with a stream containing hydrogen, and injected into the food contact with the catalyst in the reaction zone. The preferred feedstocks for catalytic dehydrogenization in the present invention should preferably contain isobutane and may also contain propane and pentane. The process of catalytic dehydrogenization used for processing, mainly paraffin hydrocarbons, so as to obtain olefinic hydrocarbon compounds. Corresponding zones of dehydrogenization for this process to provide the necessary degree of conversion of isobutane in isobutane at a relatively low speed contaminated the e in accordance with the present invention. The specific configuration of the dehydrogenation reactor depends on the technical characteristics of the catalyst from the reaction zone. Low degree of conversion for work zone dehydrogenization usually provide for the production of olefins in the range of 10-40 wt.% and more typically in the range of 20-30 wt.%. Working conditions in the area of dehydrogenization preferably selected to produce the exit stream of olefins, in which the ratio of isobutene and normal butene and propylene is greater than one. Regeneration of the catalyst may be accomplished by rocking the layer, plurigenera or by using the partition constant regeneration of the catalyst.

The reaction of catalytic dehydrogenation is usually performed in the presence of catalyst particles containing one or more noble metals of group VIII (e.g., platinum, iridium, rhodium, palladium), United with a porous carrier such as a refractory inorganic oxide. Usually the preferred catalyst contains a platinum group metal, alkaline metal with a porous media. The catalyst may also contain metals promoters, which effectively improve the effect of a catalyst. Preferably, the porous material of the carrier to the soup 25-500 m2/, Porous material of the carrier should be relatively stable to the conditions used in the reaction zone, and may be selected from a range of materials-carriers, which are traditionally used in hydrocarbon conversion catalysts dual purpose. Porous media can be, thus, selected from a range that includes activated carbon, coke or charcoal, silica or silica gel, clays and silicates including synthetically derived and natural, which may or may not be treated with acid, as, for example, attapulgite clay, diatomaceous earth, kieselguhr, bauxite; refractory inorganic oxide, such as titanium dioxide aluminum (alumina), titanium dioxide, zirconium dioxide, magnesium oxide, silica alumina, alumina, Boria; crystalline alumina silicates, such as natural or synthetic mordenite or a combination of one or more of these materials. The preferred porous carrier is a refractory inorganic oxide, and the best results were obtained when using alumina as the carrier. Alumina is the most commonly used media. The preferred alumina known as "gam the jam must have media from metaglidasen in the form of spherical particles. Particles having a relatively small diameter of the order of 1, 6 mm, preferably, however, the particles may be of larger diameter, to 5.6 mm

In the final composition of the catalyst may be a component of the platinum group in the form of compounds, such as oxide, sulfide, halide, oxysulfide etc., elemental metal or in combination with one or more other ingredients of the catalyst. There is a belief that the best results are achieved when almost all of the components of the platinum group are present in the elemental state. Components of the platinum group usually be 0.01 - 2.0 wt.% the composition of the finished catalyst, calculated in the calculation of the basic composition. Preferably, the platinum component of the catalyst was 0.2 to 1.0 wt.%. The preferred platinum group component is platinum, and palladium is the preferred metal. The preferred platinum group component may be introduced into the catalyst composition in any suitable way, such as joint deposition or co-gelation with the preferred material carrier, or by ion exchange or impregnation of the material carrier. Preferred SPO is the lia platinum group for impregnation of the calcined material of the carrier. For example, platinum group component may be added to the support by mixing the support with an aqueous solution chloroplatinate or chloroplatinic acid. Acid such as hydrogen chloride, usually added in the impregnating solution, to facilitate the distribution of the platinum group component to the material to the media. In the preferred material of the catalyst, the concentration of chloride is 0.5 to 3.0 wt.%.

The preferred alkali metal is usually either potassium or lithium, depending on the feed hydrocarbon. The concentration of the alkali metal may be 0.1 to 3.5 wt.%, but preferably 0.2 to 2.5 wt.% in the calculation of the basic composition. This component can be added to the catalyst by using the methods described above, a separate step or simultaneously with the dissolution of another component.

The dehydrogenation catalyst may also contain metal-promoter. One such preferred metal-promoter is tin. Component tin should contain from 0.01 to 1.0 wt.% the tin. It is preferable that the ratio (atomic) of tin and platinum was in the range of 1:1 to 6:1. Component tin may be introduced into the catalyst composition in any suitable known way is a recreational way of introducing a tin component includes a joint precipitation during the preparation of the preferred material of the carrier. A more detailed description of the preparation of the material carrier, and adding a platinum group component and a tin component is described in U.S. patent US-A-3745112.

Work zones dehydrogenization not in very hard conditions can increase the life span of the catalyst. Depending on the catalyst system and power properties of the zone dehydrogenization in the reaction zone of dehydrogenization use solid catalyst, which can act as a stationary layer, polarieverywhere layer or constantly regenerated catalyst. The actual structure of the zone dehydrogenization can be relatively simple and include one reactor and one heater. Moreover, catalytic dehydrogenation reaction zone may consist of multiple layers of catalyst. In one such system, the catalyst is placed in a circular frame, through which it can move under its own weight. Preferred methods of dehydrogenization light hydrocarbons suitable for implementation of the permanent dehydrogenization isobutane with a system of continuous regeneration of the catalyst described in U.S. patent US-A-5227566, US-A-3978150, US-A-3856662 and US-A-3854887.

In the preferred form have successfully enables support for the production, at a time when the catalyst is removed or replaced. In a typical reaction zone with a movable layer of fresh catalyst particles are fed through the reaction zone under the effect of their gravity. The catalyst removed from the bottom of the reaction zone and sent to a regeneration zone, where they perform hereinafter described multi-stage regeneration process to fully restore the promoting ability of the catalyst. The catalyst moves under its own gravity through the various regeneration zone and then it is removed from the regeneration zone and transported to the reaction zone. The structure of a typical section of the combustion, drying and dispersion in the movable layer is described in U.S. patent US-A-3653231 and US-A-5227566.

Operating conditions for the preferred zones of dehydrogenization made in accordance with the present invention are typically characterized by an operating temperature in the range 510-649oSince, moreover, it is preferable operating temperature, comprising at least 593oWith, and especially preferred is the working temperature of 610oC. a Relatively high working pressure meets the low conversion of the preferred zones of dehydrogenization and is usually 377-1136 kPa. Over the m values of the pressure within 584-860 kPa especially preferred. Conditions low conversion also allows the work zone dehydrogenization with a low ratio (molar) of hydrogen and hydrocarbons in the range of 0.1 to 4.0 and more preferably about 0.2. Volumetric flow rate for the zone of dehydrogenization is 0.5-50,0 h-1and in normal conditions exceed 10,0 h-1and typically equal to approximately 15,0 h-1.

Low conversion rates and lower pressures for the reaction zone of dehydrogenization also help to reduce hardware costs, if this area is combined with the oligomerization zone of low pressure in accordance with the present invention. For example, the compressor of the reciprocating action can be used in regeneration of recycled hydrogen from the exit stream of dehydrogenization. Higher pressure in the area of dehydrogenization and its Association with the reaction zone of saturation can also reduce hardware costs associated with the supply of hydrogen and its regeneration. The use of relatively high pressures in the area of dehydrogenization may also contribute to the regeneration of the hydrogen flow, the purity of which is 80% or higher with a minimum cooling. The low degree of Conviasa the amount of recycled isobutane, which reduces the detrimental effect any transfer of the olefin in the area of dehydrogenization. In addition, the supply of excess hydrogen from dehydrogenization in the zone of saturation leads to regeneration of excess isobutane in a primary distillation column, which otherwise would cause unacceptable loss of such hydrocarbons as the process is running.

The method and its various stages will hereinafter be described with reference to the accompanying drawing.

The stream of light hydrocarbons, rich WITH4-paraffins, served in deisobutanizer column 10 through line 12. In column 10 also serves recirculated material from here and further described in the exit stream in line 96 of the zone of saturation. From deisobutanizer column 10 away side stream 16, which contains the supply isobutane stream to the hydrogenation zone.

On line 16 power flow for zone dehydrogenization through the heat exchanger 98. The heat continue, when on line 16 first pass power to zone dehydrogenization through the heat exchanger 18 for heat recovery from exhaust 20 zone dehydrogenization. Depending on the composition of the dehydrogenation catalyst a small amount of sulfur may be added and for passivation of the catalyst. Line 22 connects the supply flow zone dehydrogenization with a hydrogen-containing stream to obtain a combined feed stream, which is passed by line 28 through boot heater 24 and the reactor 26 of dehydrogenization. Upon contact with the dehydrogenation catalyst is dehydrogenization parts of paraffin components of the feed stream that exits the reactor 26 zone dehydrogenization on line 30 and through an intermediate heater 32 to heat the reaction for conversion of isobutane in the second reactor 34. On line 36 direct the heated and partially converted power zone dehydrogenization through the reactor 34 and out of the reactor through line 20 to supply the exhaust stream.

The figure schematically shows the action zone dehydrogenization with regeneration of the catalyst, by means of which serves the recovered catalyst in the reactors 26 and 34 to line 38. Using the supply system of the catalyst (not shown) submit zakoksovanie catalyst from the lower zones of the reactor 26 and 34 on the pipes 42 and 44, respectively, and the pipe 46 to lift the catalyst. Pipeline 46 zakochany the catalyst is delivered to section 48 regenerate>/P>Output stream 20 from the reaction zone of dehydrogenization contains at least hydrogen, butane, butenes, a certain amount of light hydrocarbons and small amounts of heavy hydrocarbons. On line 20 output stream from the reaction zone of dehydrogenization is directed through the heat exchanger 18 and the separator 52 to remove various heavy hydrocarbon components from the exhaust stream is directed through line 57. On line 56 serves the exiting flow of steam from the separator 52, which removed various heavy hydrocarbon components in the hydrogen separator 54 for the regeneration of the hydrogen-containing stream, where hydrogen is returned to the supply flow zone dehydrogenization on line 22. Zone dehydrogenization in accordance with the present invention can operate using a flow of hydrogen is relatively low purity. The appropriate recycle stream of hydrogen for the implementation of the present invention can contain hydrogen at a concentration of less than 90 wt.%. Acceptable concentrations of hydrogen can be very low and reach up to 70%, but more typical are the concentrations of the order of 80%, with the remainder hydrogen-containing stream includes C1-C3-hydrocarbons and usually small concentration is on, if required, by cooling the separated product from the separator 54 to reduce the transfer of contaminants in stream 22. The remaining portion of the hydrogen-containing stream that is returned from the separator 54, serves as a supply of hydrogen reactor saturation, as will be described later.

Sedimentary stream 60 from the separator 54 is connected with registertask stream 58 to receive the combined power to the reaction zone oligomerization. Registergui stream supplied through line 58, delivers more olefins in the reaction zone oligomerization. Olefins supplied through line 58, comprise normal butenes and isobutene and can also include3- and5-olefins, as well as some waxes.

On line 60 serves a combined flow of power to the oligomerization zone, going through the heat exchanger 62 to heat the food for oligomerization relative to the exhaust flow of the zone of saturation. On line 64 serves hot meals for oligomerization through balancing the heater 66 to the reactor through line 68. In accordance with the present invention the heavy recycle stream containing8-paraffins and above, combined with the power of the oligomerization zone in l is divided into many stages. Additional amount of C8-olefins and higher let into the space between the layers in the reactor 70, using cooling distribution lines 72 to deliver the refrigerant in each of the inner reaction zone. The injection cooler receive heavy substances in accordance with the present invention to provide the advantages of washing of the catalyst and the selectivity to high-octane8-isomers. The pipe 74 to the exhaust stream connected to the output oligomerization reactor 70, provides regeneration effluent stream oligomerization.

On line 76 serves hydrogen in the exit stream from the reaction zone oligomerization to get the combined power to the reaction zone of saturation. On line 76 take the hydrogen out of the process of dehydrogenization coming through the lines 78 and 80. Any external source of hydrogen is connected to the process via line 82. Preferably, all outflows from the oligomerization was held together with added hydrogen directly into the zone of saturation.

In the zone of saturation saturate unsaturated components boiling within the gasoline and unreacted light olefin oligomerization zone to get than necessary, sequentially, the food served through each reactor. The serial arrangement of reactors allows you to control the temperature.

On line 84 hold the combined power through indirect heat exchange relative to the exit stream from the zone of saturation in the heat exchanger 86 and heated power saturation zone served by the line 87 in the first reactor saturation 82. The preferred structure of the hydrogenation zone is a system of two-stage vododobyvayuschih reactor, in which the exit stream from the first reactor saturation 82 is passed through line 85 to the second reaction zone 86. For monitoring the temperature in an exothermic process in the first reaction zone a portion of the heavy hydrocarbons is passed as a cooling flow along the line 88 in combination with the output stream 85 of the first reaction zone. Now saturated output stream of the reaction zone oligomerization direct on line 90 through the heat exchanger 86 and 91 to the heat exchanger 62 for heat recovery by the reaction of saturation before entering into the cooling separator 92 through line 93.

Cooling separator 92 extract rich heavy hydrocarbons from the zone of saturation to create a recirculating flow in the liquid phase in the config hydrocarbon in the area of oligomerization on the lines 69 and 72 and the cooling flow in the reaction zone of saturation on line 88. The rest of the heavy hydrocarbons return to the top of the separator through line 95 to the output line 102 of the process in the form of settling products from deisobutanizer 10. The substance of the upper part of the separator 92 and any additional sludge separator supplied through the line 95, take line 96, and they contain mainly alkylate and unreacted butane. Additional heat recovery from the stream 96, coming from the separator is produced in the heat exchanger 98 to first raise the temperature of the incoming air stream supplied to the zone dehydrogenization from deisobutanizer 10.

In a preferred embodiment of the present invention use deisobutanizer 10 for simultaneous distillation product containing the upper stream of the separator, together with the initial receipt of the supply flow zone dehydrogenization. In the preferred deisobutanizer vessel to produce the separation of light residues, such as the upper flow 100, and simultaneously deliver the previously described feed stream 16 zone dehydrogenization as a by-product. Design deisobutanizer columns is to provide a high concentration of em to get isobutane with a purity of 80 wt.% and more preferably with a purity of at least 95 wt.%. The column also produce Department WITH the remaining4connections minus the materials from the inlet flow area of dehydrogenization supplied through the upper pipe 100. From the primary distillation column also serves saturated components of the product from the oligomerization zone in the bottom of the stream 102. From deisobutanizer this preferred structure may also be lodged with the normal side WITH4- the stream to remove excess unreacted butane from the process. Side stream divert on line 104, which can be used to supply additional heat, using indirect heat exchange with the input stream4in the heat exchanger 106.

For a more complete demonstration of the attendant advantages of the present invention were produced in the following experiments.

EXAMPLE 1

Feeding a stream containing 13 wt.% normal butene of 17.7 wt.% isobutene and 69.4 wt.% isobutane, enter in contact with approximately 50 cm3catalyst in the form of a solid phosphoric acid containing calcined mixture of phosphoric acid in siliceous basis. The catalyst consists of granules with a diameter and a length of approximately components of 6.4 mm, and galagedara the careful packing sand free space between the granules. Power was applied to the reactor at a temperature of 190oWith and under pressure 3549 kPa and it passed through the reactor with a volume hourly flow rate of 2 h-1. In the oligomerization reaction, the maximum temperature of the reactor was increased to 202oC. a Sample of the exit stream from the reactor was separated, person to distil and produced its analysis to determine conversion4-olefins and the selectivity to hydrocarbons. The results of the analysis are presented in table.1.

EXAMPLE 2

To demonstrate the benefits of recycling heavy hydrocarbons, added an additional 25 wt.% normal C8-paraffins in food in Example 1, and conducted an experiment under the same operating conditions and with the same catalyst as in Example 1. In the oligomerization zone maximum temperature reached 200oC. a Sample of the exit stream from the reactor also person to distil and made his analysis and found that the contents of C8olefin is significantly higher. The results of the analysis are also presented in table.1.

Comparison of Examples 1 and 2 shows that approximates the conversion of C4-olefin Appendix C8-paraffins shifts the selectivity of the exit stream reduces the formation of both isomers with a large number of carbon and with fewer carbon.

As this example shows, the recycled stream in accordance with the present invention provides a significant increase in the selectivity of the oligomerization against high-octane isomers WITH8.

EXAMPLE 3

Feeding a stream containing 8.4 wt. % of normal butene, of 21.9 wt.% isobutene and to 69.7 wt.% isobutane were introduced into contact with approximately 50 cm3catalyst in the form of a solid phosphoric acid of the same type used in Examples 1 and 2. The reactor vessel and Metodologicheskie catalyst were the same as in Examples 1 and 2. Power was applied to the reactor at a temperature of 191oWith and under pressure 3549 kPa and it passed through the reactor with a volume hourly flow rate of 2 h-1. In the oligomerization reaction, the maximum temperature of the reactor was increased to 206oC. a Sample of the exit stream from the reactor was removed, was hydrogenosomal, person to distil and produced its analysis to determine conversion4-olefins selectivity to hydrocarbons and theoretical (EXACT) and motor (MOC) octane numbers. The results of the analysis are presented in table.2.

EXAMPLES 4-6

To trademonster, added additional 25 wt. % the normal C12-paraffins in the feed that was used in Example 3, and perform the experiment in the range of low temperature and bulk velocity with the same catalyst as in Example 3. Data conversion, selectivity and theoretical (EXACT) and motor (MOC) octane numbers after distillation and hydrogenation output streams for the three experiments with added C12components in the diet are presented in table.2 number of Examples 4-6.

All three experiments conducted at low temperatures in comparison with Example 3, and show a significant increase in the selectivity to C8by adding a normal C12-paraffins. Examples 4 and 5 show a significantly higher selectivity to C8-isomers than in Example 3, with only a slightly reduced level of conversion. The result of the increased selectivity is higher octane number than that obtained without12components in Example 3. Example 6 shows that a higher level of conversion than in Example 3, can be achieved at relatively higher temperatures and lower flow rate while maintaining much b the oligomers using the oligomerization zone and the zone of saturation, characterized in that it includes the following steps: a) introducing a feed stream to the oligomerization zone that contains3-C5the olefin recycle stream containing paraffins, having the number of carbon atoms of at least 8, in interaction with acidic oligomerization catalyst located in the zone of oligomerization, which support the oligomerization conditions effective to obtain the exit stream containing olefins having a number of carbon atoms of at least 7; (b) removing the exit stream of oligomerization of the said oligomerization zone containing the said paraffins and olefins having the number of carbon atoms of at least 7; (c) direction, at least part of the exit stream from the oligomerization zone and a hydrogen-containing stream in the saturation zone containing the catalyst saturation under conditions of saturation, effective to saturate olefins in the above-mentioned part of the exit stream oligomerization, and removing the exit stream of the zone of saturation; (d) the direction of the above the exit stream of the zone of saturation, containing paraffinic hydrocarbons having at least 8 carbon atoms, said oligomerization zone as in is it WITH8-paraffins and higher.

2. The method according to p. 1, characterized in that the catalyst in the oligomerization zone contains a solid catalyst of phosphoric acid.

3. The method according to p. 1, characterized in that the oligomerization zone operates at a pressure less 2170 kPa.

4. The method according to p. 1, wherein the recycle stream contains paraffins8or12.

5. The method according to p. 1, characterized in that the feed stream contains4-olefins.

6. The method according to p. 1, characterized in that at least part of the exit stream of the zone of saturation is cooled and returned to the oligomerization zone and the zone of saturation in the form of a coolant, which is part of these recirculation flow.

7. The method according to p. 6, characterized in that the oligomerization zone contains a solid catalyst of phosphoric acid and in which the aforementioned coolant Inuktitut in the above-mentioned layer, at least one intermediate position.

8. The method according to p. 1, characterized in that the oligomerization zone operates at 149-204oC.

9. A method of producing motor fuel containing isobutane input stream using zone digid the input stream, containing isobutane, in the area of dehydrogenization and the introduction referred to the input stream into contact in the mentioned area dehydrogenization with a dehydrogenation catalyst under conditions of dehydrogenization, effective to obtain the exit stream containing hydrogen and C4-isoolefine; (b) separating the stream of hydrogen purity hydrogen which is 70-95 mol. % of said exit stream of dehydrogenization, and extracting a feed stream containing4-isoolefine from the rest of the mentioned exit stream of dehydrogenization, to obtain the supply flow oligomerization zone; (c) the direction referred to the supply flow of the process of oligomerization and paraphysomonas recirculation flow in said oligomerization zone and introducing into contact with the solid oligomerization a catalyst effective to obtain the exit stream oligomerization containing isoolefine having the number of carbon atoms of at least 8; (d) direction, at least part of the flow of hydrogen, selected on the stage), and the exit stream from the oligomerization zone without separation in the saturation zone containing the catalyst saturation, supported in terms of saturation and effective for proceduralise, at least part of the mentioned exit stream of the zone of saturation in the zone of separation to extract a busy stream containing paraffins, having the number of carbon atoms of at least 8; (f) returning at least part of the above mentioned saturated flow in step (C) as mentioned recirculation flow.

10. The method according to p. 9, characterized in that the concentration of olefin in the feed stream to the oligomerization zone is at least 30 wt. % isobutene.

 

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