JPH03182592A - Method for continuous catalytic and selective production of aromatic hydrocarbon - Google Patents

Abstract

PURPOSE: To enable a selective and continuous preparation of aromatic- containing hydrocarbon product by subjecting hydrogen and an olefin-containing hydrocarbon starting material to hydrogenation reaction, then subjecting the resulting product to dehydrocyclodimerization reaction and recovering the obtained aromatic product.

CONSTITUTION: (A) hydrogen and (B) an olefin-containing 2-5C hydrocarbon starting material are passed through a hydrogenation reaction zone having a hydrogenation catalyst and reacted to obtain a hydrogenation reaction zone product flow containing an olefin-based hydrocarbon in an amount of less than 50 mole % or more of that of component (B). Then, at least part of this product flow is passed through a dehydrocyclodimerization reaction zone containing a dehydrocyclodimerization catalyst and reacted to form a dehydrocyclodimerization reaction zone product containing H, methane, ethane, ethylene, a 3-5C aliphatic hydrocarbon, a 6 or more C aliphatic hydrocarbon and an aromatic hydrocarbon and subsequently, this product is recovered.

COPYRIGHT: (C)1991,JPO

Description

Detailed Description of the Invention [Technical Field] The present invention relates to hydrogenation and dehydrogenation cyclodimerxation.
The present invention relates to a selective combination process effective for converting C2 to C5 olefinic hydrocarbons to aromatic hydrocarbons, respectively.

[Background of the invention]

Dehydrogenated cyclodimerization contains 2 to 5 carbons per molecule,
This is a reaction in which paraffin and olefinic reactants are reacted on a catalyst to obtain aromatics mainly containing H2 and light components as by-products. This process is quite different from conventional reforming or dehydrocyclization processes that convert C1 and above reactants, primarily paraffins and naphthenes, to aromatics.

These aromatics show that the number of carbon atoms per molecule is equal to or less than the number of carbon atoms in the raw reactant, indicating that dimerization is slow. On the other hand, the aromatic products obtained by dehydrogenation cyclodimerization always have a higher number of carbon atoms per molecule compared to C2-C5 reagents, and therefore dimerization is the main step in dehydrogenation cyclodimerization. It shows that. Usually, the dehydrogenation cyclodimerization reaction is carried out at a temperature exceeding 260° C. using a bifunctional catalyst containing an acidic component and a dehydrogenation component. Such catalysts include acidic amorphous aluminas containing metal promoters.

Recently, crystalline aluminosilicates have been effectively used as catalyst components for dehydrogenation cyclic dimerization reactions. Crystalline aluminosilicates are generally called zeolites and have the empirical formula % (where n is the valence number of M, 5M is generally hydrogen or an element of group Ⅰ or Ⅰ, particularly Na, K, Mg).

Ca, Sr or Ba, x is generally 2 or more.

). Zeolites are S i04 and AQO4
It consists of a three-dimensional network structure of tetrahedrons, the corners of which have a skeletal structure connected to each other by shared oxygen atoms. The larger the ratio of genus 5i04 to genus AQO4,
Zeolites become suitable as dehydrogenation cyclodimerization catalysts.

Such zeolites include the mordenite and ZSM varieties. In addition to the zeolite component, certain metal promoters and an inorganic oxide matrix are included in the dehydrogenation cyclodimerization catalyst formulation. Examples of inorganic oxides include silica, alumina and mixtures thereof. A metal component such as a Group I or II metal of the Periodic Table is used to provide a dehydrogenation function. Acidic functionality can be provided by an inorganic oxide matrix, a zeolite, or both.

When olefins are used as raw materials for dehydrogenation cyclodimerization reactions, aromatic selectivity decreases compared to paraffinic raw materials.
It has also been found that catalyst deactivation (coking) increases. Rapid deactivation is believed to cause excessive carbon formation (coking) on the catalyst surface. The tendency toward coking often makes catalyst regeneration expensive and laborious. It is a particular object of the present invention to reduce the tendency of the catalyst to coke and thereby increase the life of the catalyst.

Conventionally, many methods have been known for producing aromatics from hydrocarbon raw materials containing C2 to C5 olefinic hydrocarbons through two-stage reactions. None of these methods embodies all aspects of the two-step reaction conversion process of the present invention, nor do they have the same advantages and benefits as the unique process of the present invention.

U.S. Pat. No. 4,554,393 (Liberts et al.) discloses a process for dehydrogenating a paraffin-based feedstock and then cyclodimerizing the alkenes in a second reactor to produce aromatics. The method of the present invention is at least different from the method of the above-mentioned patent in that the second reaction is dehydrogenation and cyclic dimerization in which an alkane is dehydrogenated to form an alkene, and then the obtained dehydrogenated hydrocarbon is cyclically dimerized. One means is different. The second reaction of the present invention realizes both reaction steps of the patent. Therefore, the method of the above-mentioned patent is similar in function only in the second dehydrogenation cyclodimerization reaction zone of the present invention.

European Patent No. 310.162,636 describes a process for producing aromatics from light olefin-containing feedstocks. The process involves contacting an olefin-containing feedstock with two reaction zones containing the same catalyst. Here, the catalyst in the first reaction zone has been deactivated by coke deposits, and the catalyst in the second reaction zone is essentially fresh. The main reaction that takes place in the first reaction zone is the dehydrocyclodimerization of olefins, while the main reaction that takes place in the second reaction zone is the dehydrocyclodimerization of paraffins. The European patent application differs from the present invention in that essentially the desired liquid product is not obtained in the first reaction zone. In the process of the invention, olefins are converted to paraffins and not to aromatics as in the prior art.

[Purpose and implementation mode]

The primary object of this invention is to provide a process for selectively converting olefinic hydrocarbon feedstocks to aromatic-containing hydrocarbon products. Furthermore, this method reduces the percentage of catalyst deactivation compared to conventional single-step methods using olefin feedstocks.

The conversion selectivity of the dehydrogenation cyclic dimerization reaction zone catalyst is improved. Accordingly, a broad embodiment of the present process relates to a selective continuous catalytic process for the production of aromatic hydrocarbons from C2-C5 olefin-containing feedstocks.

The process first includes an olefin-containing C2-C5 hydrocarbon feedstock along with hydrogen and a hydrogenation catalyst.

and passing at least 50 mole percent less olefins than the C2-C5 feed through a hydrogenation reaction zone operated at hydrogenation reaction conditions effective to obtain a hydrogenation reaction zone product stream. Hydrogen, methane, ethane, ethylene, C3 to C5 are passed through a dehydrogenation cyclic 2'Ik formation reaction zone which contains a dehydrogenation cyclic dimerization catalyst and is operated under dehydrogenation cyclic dimerization conditions. A dehydrogenated cyclic dimerization reaction zone product is produced that includes aliphatic hydrocarbons, C1 aliphatic hydrocarbons, and aromatic hydrocarbons. Finally, the products of the dehydrogenation cyclodimerization reaction zone are a fraction containing light gases such as hydrogen, methane, ethane, and ethylene, a C3-C5 aliphatic hydrocarbon recycle stream, and a C6+
Separated into aliphatic and aromatic hydrocarbon product streams. In a narrower embodiment, the continuous multistage catalytic process of the present invention comprises hydrogenating an olefin-containing 02-C5 hydrocarbon feedstock with a hydrogenation catalyst and at a temperature of 50-250°C, preferably 50°C.
~150°C, pressure 1.0~50 atm.

An essentially olefin-free hydrogenation reaction zone, preferably operated under hydrogenation conditions of 1.0 to 25 atm and a liquid hourly space velocity of 0.1 to 20/hr. First, the hydrogenation reaction zone product, which is substantially free of olefins, is passed through a dehydrogenation cyclodimerization reaction zone, and the C2 to C5 aliphatic hydrocarbons recovered in the separation zone are recycled. This dehydrogenation cyclic dimerization reaction zone contains a dehydrogenation cyclic dimerization catalyst and has a temperature of 350 to 700°C.

Preferably 400-650°C, pressure 0.25-20 atm, preferably 0.25-10 atm, and hourly space velocity of liquid 0.5-20/hr, preferably 0.5-1
0/hr dehydrocyclodimerization reaction zone product containing hydrogen, methane, ethane, ethylene, C3-C5 aliphatic hydrocarbons, and CI aliphatic and aromatic hydrocarbons. get. Finally, the fraction is collected and the C2
~C5 aliphatic hydrocarbons are recycled to the inlet of the dehydrogenation cyclodimerization reaction zone.

[Description of drawings]

FIG. 1 is a process flow sheet showing a typical conversion technique of the present invention.

Although FIG. 1 illustrates the production of aromatic hydrocarbons from a feed stream containing C2 to C5 olefins and paraffins according to the process of the present invention, the invention is not limited thereby.

The first step of the process of the invention is to pass a mixed hydrocarbon feed stream 4 through a hydrogenation reaction zone 5 containing a hydrogenation catalyst, this mixed feed stream 4 comprising a C2 to C5 olefinic and paraffinic hydrocarbon stream 1 and Contains a portion of hydrogen feed stream 3. This hydrogen feed stream is further supplied to line 14 as recovered hydrogen from the separation zone.
It is characterized in that it can be supplied from hydrogen, it can be supplied as fresh hydrogen from line 2, or it can be supplied as a combination of both. No matter how the hydrogen is fed to the hydrogenation reaction zone, the hydrogen in feed stream 3 is transferred to the hydrogenation reaction zone elution stream 7.
and a C2-C5 olefinic and paraffinic hydrocarbon feed stream to form a mixed hydrocarbon feed stream 4 to hydrogenation reaction zone 5.

The hydrogenation reaction zone effluent stream 6 is divided into two parts: a first part containing the hydrogenation reaction zone recycle stream 7 and a second part containing fresh feed to the dehydrogenation cyclodimerization reaction zone 10. The dehydrogenation cyclodimerization reaction zone IO receives the blended feedstock from line 9 containing part of the hydrogenation reaction zone effluent stream 8 along with the C2-C5 aliphatic hydrocarbons circulated from the separation zone 2 via line 13. .

The product of dehydrogenated cyclodimerization zone 10 enters separation zone 12 through line 11. Separation zone 12 is aromatic-containing dehydrogenated ring 2
The quantified eluate stream 11 was converted into dihydrogen, methane, ethane, ethylene,
Any means useful for separating product fractions containing C3-C5 aliphatic hydrocarbons and C6° aliphatic and aromatic hydrocarbons may be used. Hydrogen is separated from separation zone 12 into line 15 where a total hydrogen stream 18 and a recycled hydrogen stream 15 are used as part of the total feed of the hydrogenation reaction zone.
It can be divided into From separation zone 12, C61 aliphatic and aromatic hydrocarbons are recovered through line 17, while
Through line 16 methane and possibly any ethane and a portion of ethylene are recovered. The recycle stream 13 of this dehydrocyclodimerization reaction zone contains essentially all C3 to C5 aliphatic hydrocarbons, as well as a portion of ethane and ethylene, each produced as a result of the dehydrocyclodimerization reaction. .

[Detailed explanation]

The method for converting light aliphatic hydrocarbons to aromatic or non-aromatic C6+ hydrocarbons is as clear from the previously cited literature.
This is an important issue of development efforts. The fundamental utility of this process is to convert low cost and highly useful C2-C5 hydrocarbons into more useful aromatic hydrocarbons and hydrogen, or to convert feedstock hydrocarbons into higher molecular weight aliphatic products. It is to be. It is desirable to correct the excess amount of C2-C5 hydrocarbons or meet the demand for aromatic hydrocarbons. Aromatic hydrocarbons are extremely useful in the production of a wide variety of petrochemicals, of which benzene is the most widely used base material for hydrocarbon chemicals. This hydrocarbon feedstock is particularly useful as a blending component in high octane motor fuels.

The feed compound for this two-step process is light hydrogen having 2 to 5 carbon atoms per molecule. The feed stream may be a single or a mixture of two or more of these compounds. Preferred starting compounds are propane, propylene, butanes, and butylenes, with saturated compounds being particularly preferred. The feedstream of this process may contain varying amounts of C2 and C5 hydrocarbons. It is preferred that the concentration of C5 hydrocarbons in the feed stream of the dehydrogenated cyclodimerization process be maintained at a minimum practical level, preferably below 5 mole percent.

One aspect of the invention is that all C2-C fed to the method of the invention
Part 5 of the aliphatic hydrocarbons is an olefinic hydrocarbon. Since the coking observed in the dehydrogenation cyclodimerization reaction zone increases in conjunction with increasing olefin feedstock content, the feedstock of the present invention contains at least 25 wt% C2-C
5 olefin is preferably included. Feedstocks containing very high amounts of olefins are therefore particularly detrimental to the dehydrogenation cyclodimerization catalyst under normal conditions which makes this process even more attractive.

The preferred products of this process are C6+ aromatic hydrocarbons. However, the dehydrogenated cyclodimerization process is not 100% selective, and several types of non-aromatic C6° hydrocarbons are produced even from saturated feedstocks. When raw materials consisting of C2-C5 aliphatic hydrocarbons are processed, most of the C6° hydrocarbon products are benzene, toluene, and various xylene isomers. Small amounts of C1+ aromatics are also obtained. The presence of olefins in the feed stream increases the amount of C6° long chain aliphatic hydrocarbons. The olefin concentration in the feedstock is quite high and significantly reduces aromatics production. Both of these problems and the coking problems discussed above are solved by the two-step process of the present invention.

The subject of the present invention is to increase the amount of valuable C6° alkyl aromatics and to perform dehydrogenated cyclodimerization by first hydrogenating an olefin-containing C2-C5 hydrocarbon and then dehydrogenating the hydrogenated feedstock. The purpose is to increase the life of the catalyst.

The first step of the process of the invention is a hydrogenation step, which step comprises a hydrogenation catalyst and is operated under hydrogenation conditions such that at least 50 mol% of the olefin content of the feedstock can be converted to paraffins. The reaction takes place in a reaction zone. In a preferred case, essentially all the olefins in the C2-C5 aliphatic feed are converted to paraffins so that the hydrogenation feed going to the second stage is essentially free of olefins.

A preferred embodiment is to have little sulfur and other impurities in the first stage feedstock.

If the raw material is saturated with water, there is no need to dry it;
Free water should be avoided. This dehydration step can be omitted if the feedstock is not saturated with water.

Once the water has been removed, the feedstock is typically combined with a slight excess of hydrogen compared to the stoichiometric amount of hydrogenation. In order to limit the heat generated during the reaction by this exothermic process by limiting the maximum concentration of solenoid in the blended raw materials, it is preferable to circulate the reactor. The heat of reaction is removed by placing suitable cooling means on this circulation stream. The hydrogenation reaction proceeds under very mild pressure and temperature conditions on a fixed bed catalyst system.

Due to the very high activity of the catalyst, the space velocity is high and therefore only relatively small reaction vessels and catalyst quantities need to be provided. If there are no impurities in the feedstock, the catalyst will be very stable and its cost will not be an economic issue.

The operating conditions under which the hydrogenation reaction is effective depend on the olefins constituting the feedstock. Hydrogenation process temperature is 50-250℃
, pressure 1 to 50 atm, and hourly space velocity of liquid (
It is thought that this occurs at 0.1 to 20.0/hr (hereinafter referred to as LH5V). However, due to the exothermic nature of this method, the hydrogenation reaction step is carried out at a temperature of 50 to 150°C and a pressure of 1 to 25 atm.
Conducted under the conditions of LH3V0.1-20.0/hr,
It is also preferred to circulate the intermediate or final hydrogenated product to the inlet of the reaction zone in a weight ratio of recycled feed to fresh feed of 1 to 20.

Hydrogen is co-fed to the hydrogenation reaction zone. Preferably, the proportion of hydrogen feed is slightly greater than the stoichiometric amount required to saturate the feed olefin. However, the proportion of hydrogen can be significantly greater than the stoichiometric amount without significantly affecting the hydrogenation reaction process. The proportion of the hydrogen raw material is preferably 1 to 10 times the stoichiometric amount based on the olefin content in the hydrocarbon raw material. This corresponds to a raw material molar ratio of hydrogen to olefin of 1:1 to 10:1.

A larger amount of hydrogen is not used since this would require the use of an intermediate product separation zone in the hydrogenation reaction zone.

However, the use of intermediate product separators is within the scope of this invention.

If an intermediate product separation zone is used, any known means for separating hydrogen from hydrocarbons may be used. Such means include membrane separation zones, more preferably gas-liquid flash separation zones. Regardless of the means used, intermediate separation yields gaseous products containing hydrogen. This gaseous product can be recycled to the hydrogenation reaction zone or recovered as product and as a liquid product to the dehydrogenation cyclodimerization reaction zone.

Furthermore, this hydrogenation reaction step is characterized in that at least 50 mole percent of the olefins in the hydrocarbon feedstock are hydrogenated to paraffins. However, it is preferred that the hydrogenation of the olefin in the hydrogenation reaction step be essentially complete. That is, the hydrocarbon feedstock after the hydrogenation step is essentially olefin-free. ``Essentially free of olefins'' means that the olefin content in the hydrogenated hydrocarbon is 2.0 wt.
%, preferably less than 0.5 wt%.

Catalysts useful in the hydrogenation process of the present invention are those capable of completely hydrogenating the olefins in the olefin-containing feed to essentially olefin-free hydrogenated products as defined above. Any conventional catalyst of this type having known hydrogenation performance and capable of producing an essentially olefin-free hydrogenation product is satisfactory. Particularly useful catalysts are Group I noble metal components, especially pb on an inorganic oxide support.

The inorganic oxide support material useful as a catalyst in this process can be any known support material useful as a support for a catalyst. However, alumina is the most preferred support material. The most preferred inorganic oxide support of the present invention is 1 to 50
It is alumina with a surface area of 0rrf/g. The alumina support material may be prepared by any suitable method from synthetic or naturally occurring raw materials. The carrier may be in any desired shape such as spheres, pills, blocks, extrudates, powders, particles, etc. The preferred shape of alumina is spherical. Particles on the order of 1732 inches (0.08 cm) or less in diameter can be used, as can particles larger than 1/16 inch (0.16 cm), but the preferred particle size is about 1/16 inch (0.16 cm).
, 16 cm).

In the most preferred method, the alumina is spherical. To make alumina spheres, aluminum metal is reacted with a suitable colloidizable acid and water to form an alumina sol, and the mixture of the sol and gelling agent is then dropped into a hot bath. This mixture forms spherical particles of alumina gel in a thermal bath and is readily converted to the preferred α-alumina or τ-alumina by known methods including further aging, drying and calcination. Other shapes of alumina support material can be made by conventional methods.

As mentioned above, one feature of the catalyst compositions useful in the hydrogenation step of the present process is a noble metal component from Group I of the Periodic Table of the Elements. The precious metal of group Ⅰ is Pt.

It can be selected from the group consisting of Pd, Ir, Rh, Os, Ru, or mixtures thereof. However, Pt
Alternatively, Pd is a preferred Group I noble metal component, with Pd being the most preferred. Substantially all of the Group I noble metal components are believed to be present in the catalyst in the form of elemental metals.

The Group I noble metal component generally comprises about 0.01=10 wt% of the final catalyst composition on an elemental basis.

Preferably, the catalyst contains 0.1 to 5 wt% Group I noble metal component, particularly about 0.1 to about 1.0 wt% Pd.

The Group I noble metal component can be introduced into the hydrogenation catalyst by any suitable method, such as co-precipitation, co-gelation, ion exchange or immersion, or by deposition from the gas phase or from an atomic source, or other catalyst components. It may be introduced by similar methods before, during, or after. The preferred method for introducing the Group ■ precious metal component is
Platinum can be added to the support by soaking the refractory oxide in a solution or suspension of a decomposable compound of a Group 0 noble metal, for example by mixing the support with an aqueous solution of chloroplatinic acid. .

Other acids, such as nitric acid, or other optional components can be added to the soak solution to further aid in dispersing or fixing the Group I noble metal component in the final catalyst composition. The Group 1 noble metal component can be disposed on the catalyst by various useful methods known in the art, such as homogeneous dispersion, surface immersion, surface concentration, and the like.

The C2-C5 aliphatic hydrocarbon feed streams used in the process of the invention are produced and recovered in a zero cracking or reforming process that is expected to be available as a product or by-product of a refinery or petrochemical process. Light aliphatic hydrocarbons are examples of feed streams obtained from such processes. The products of the syngas production process are light aliphatic hydrocarbons recovered at the refinery's thermal processing facility and can thus serve as another source of feedstock for the process.

In the process of the invention, the feed stream is contacted with a hydrogenation catalyst in a hydrogenation reaction zone maintained at hydrogenation conditions.

This contacting can be accomplished using fixed bed equipment, moving bed equipment, fluidized bed equipment, or batch operations; It is preferred to use a concentrated phase moving bed apparatus such as that shown in patent 3.725,249. It is anticipated that when a fixed bed catalyst apparatus is used to accomplish the process of the present invention, the catalyst of the present invention may be contained in one or more fixed bed reactors, preferably two or more reactors. Ru.

In a fixed bed apparatus or a concentrated phase moving bed apparatus containing two different reaction zones, the feed stream is temperature controlled by any suitable means and then passed to a first hydrogenation zone containing a bed of the catalyst composition of the invention. Ru. Of course, the hydrogenation zone is preferably two separate reactors with suitable means between the reactors in order to maintain the desired conversion temperature at the inlet of the first reactor. It is also important to note that the reactants can contact the catalyst upwardly or downwardly, or preferably in a radial flow manner. Furthermore, the reactants may be brought into contact with the catalyst in a liquid phase, a mixed gas-liquid phase, or a gas phase, although the best results are obtained in the gas phase.

In a preferred hydrogenation reactor consisting of two fixed bed reactors, the temperature of the first reaction zone is such that the slip stream of hydrocarbons removed from the hydrogenation step is circulated between the first and second reaction zones. It is preferable to control by. Further, in order to remove the heat generated by the exothermic hydrogenation reaction, it is preferable to use a cooling means on this circulating sliding flow.

As mentioned above, the weight ratio of the circulating slip stream to the fresh feed entering the hydrogenation zone ranges from 1.0 to 20.

The liquid product of the hydrogenation reaction step is then passed through a dehydrogenation cyclodimerization reaction step. This dehydrogenation cyclic dimerization reaction step is operated under dehydrogenation cyclic dimerization reaction conditions and includes a dehydrogenation cyclic dimerization reaction catalyst.

The raw material compound for the dehydrogenation cyclic dimerization reaction step is a light paraffinic hydrocarbon from the hydrogenation step having 2 to 5 carbon atoms per molecule. This feed stream may be one or a mixture of two or more of these compounds. Preferred starting compounds are propane and butane. A preferred feed stream for this process may include several C2 and C5 hydrocarbons. It is preferred that the concentration of C5 hydrocarbons in the feed stream of the dehydrocyclodimerization process be maintained at a minimum practical level, preferably below 5 mole percent. The preferred products of this process are C,+ aromatic hydrocarbons. However, the dehydrogenated cyclodimerization process is not 100% selective and several non-aromatic C-hydrocarbons are produced even from saturated feedstocks.

When processing feedstocks consisting of propane and/or butane, most of the C6° hydrocarbon products are benzene, toluene,
and various xylene isomers. A small amount of C, +aromatics is also produced.

The configuration of the reaction zone and the composition of the catalyst used within the reaction zone are not essential elements of the invention or limit the features of the invention. However, in order to provide some background on this process, we believe it is useful to indicate a preferred reactor for use in the present invention. Such devices include U.S. Pat. No. 3,652,231; 3,692,496.
; 3.706.536 ; 3.785.963 ;
3,825.116; 3,839. +96;
3.839.197 ; 3.854.887 ; 3,
856,662; 3,918,930; 3,9
g1,824; 4.094,8+4; 4.110
, 081; and 4,403,909.

These patents also describe catalyst regeneration systems and various catalyst bed operating systems and devices. This reactor is widely used industrially for reforming naphtha fractions. The use of this device for the dehydrogenation of light paraffins is also described.

A preferred moving bed reactor has a diameter of about 1/64 inch (0,
04 cm) to 1/8 inch (0.32 cm) spherical catalyst is used. This catalyst is preferably one in which a metal component is deposited on a support material by immersion or co-precipitation.

The above-cited references point out that it is current fashion to use zeolitic support materials, and for catalysts ZS is often identified as the preferred material in this field.
Reference is made to M-5 type zeolite. If properly formulated, the zeolite material itself exhibits considerable activity in the dehydrogenocyclodimerization reaction; it is preferred to use a metal component within the catalyst system to enhance the activity of the catalyst. A preferred metal component is Ga.

The dehydrogenation cyclodimerization conditions used in the method of the present invention are, of course,
It varies depending on factors such as raw material composition and desired conversion. Essentially all C2-C5
The preferred range of conditions for dehydrogenation and cyclic dimerization of raw materials containing paraffinic hydrocarbons is a temperature of about 350 to 700°C, a pressure of about 0.25 to 20 atm, and a LH3V of 0.5 to 20/
It is hr. Preferred process conditions are temperatures of about 400-6
50°C, pressure approximately 0.25-10 atm, and L HS V
o, 5 to 10.0/hr. As the average carbon number of the feedstock increases, a lower temperature range is required for optimal finishing; conversely, as the average carbon number of the feedstock decreases, a higher temperature in the reaction zone is required. I understand.

The feed stream for the dehydrogenation cyclodimerization process is herein defined as all streams introduced into the dehydrogenation cyclodimerization reaction zone. The feed stream contains C2-C5 paraffinic hydrocarbons. ° "C2-C5 paraffinic hydrocarbon" means one or more open chains having 2 to 5 carbon atoms per molecule;
It refers to straight chain or side chain isomers. Furthermore, the hydrocarbons in the feedstock are essentially saturated. That is, this hydrocarbon is 2
．． Contains less than 0 wt% olefin.

The C3 and/or C4 hydrocarbons are preferably selected from isobutane, n-butane and propane. A diluent may also be included in the feed stream.

Examples of such diluents include hydrogen, nitrogen, helium,
Argon, neon, CO, Co2. Examples include NH4, H2O or their precursors. The water precursor is defined as a compound that liberates H2O when heated to the dehydrogenation cyclodimerization reaction temperature.

In addition to the hydrogenation reaction zone product stream, the dehydrogenation cyclodimerization reaction zone is expected to be fed with a recycle stream containing C2-C5 aliphatic hydrocarbons. A C2 to C5 aliphatic hydrocarbon recycle stream is recovered from the product stream of the dehydrocyclodimerization reaction zone for further processing. Preferably, this recycle stream contains C3 to C5 aliphatic hydrocarbons, along with small amounts of ethane and ethylene. The rate of recycle flow varies depending on the selectivity and conversion power of the dehydrogenation cyclodimerization reaction zone.

In the present invention, the dehydrogenated cyclodimerization reaction zone feed and recycle stream are contacted with a catalyst composition in the dehydrogenated cyclodimerization reaction zone maintained at dehydrogenated cyclodimerization conditions. This contact can be distantly related to the use of the catalyst composition in fixed bed equipment, moving bed equipment, fluidized bed equipment, or batch operations. However, due to the risk of attrition of valuable catalyst and the known operational advantages, it is preferable to use a fixed bed system or a concentrated phase moving bed system as shown in US Pat. No. 3,725,249. It is contemplated that this contacting step can be carried out in the presence of any conventional dehydrocyclodimerization or physical mixture of catalyst particles exhibiting similar behavior.

In a fixed bed apparatus or a concentrated phase moving bed, the feed stream is preheated to the desired reaction temperature by suitable heating means and then passed through a dehydrocyclodimerization zone containing a bed of the desired catalyst composition. It will of course be appreciated that the dehydrogenated cyclodimerization zone may be one or more separate reactors with suitable means between the reactors in order to maintain the desired conversion temperature at the inlet of each reactor. It is also important to note that the reactants can contact the catalyst bed upwardly, downwardly, or preferably in a radial flow manner. Furthermore, the reactants may be brought into contact with the catalyst in a liquid phase, a mixed gas-liquid phase, or a gas phase, although the best results are obtained in the gas phase.

It is also preferred that the dehydrocyclodimerization device is a dehydrocyclodimerization zone comprising a fixed bed or a concentrated phase moving bed of one or more catalyst compositions of the present invention.

Of course, in the case of a multi-bed system, one or more beds may use one type of dehydrogenation cyclodimerization catalyst, and the remaining beds may use other types of dehydrogenation cyclodimerization catalysts or catalysts that behave similarly. It is within the scope of the present invention to use In a multi-reactor dehydrocyclodimerization zone there may be one or more separate reactors with suitable heating means between the reactors to compensate for the heat losses encountered in each catalyst bed. In the case of concentrated phase moving bed equipment, the catalyst is especially removed from the bottom of the reaction zone.

It is customary to regenerate it by conventional means known in the art and then return it to the top of the reaction zone.

Preferred catalysts useful in the dehydrogenated cyclodimerization step of the process of this invention are those comprising a phosphorous-containing alumina, a gallium component, and a crystalline aluminosilicate zeolite with a silica to alumina ratio of at least 412. Preferred catalysts further include the crystalline aluminosilicate being ZSM-5, and 3
It is characterized by being contained in an amount of 5 to 59.5% by weight. The most preferable catalyst contains 0.1 to 5.0 wt% gallium.
and 40 to 6 ht% of alumina containing 5 phosphorus. Such catalysts are described in U.S. Pat. No. 4,636,483, incorporated herein by reference.

The hydrocarbon products of the two-stage process of the present invention are directed to separation zones for separation into specific product fractions. The hydrocarbon products of this two-stage process can be separated into fractions consisting of hydrogen, methane, ethane, and ethylene; C3-C5 aliphatic hydrocarbons; and CG9 aliphatic and aromatic hydrocarbons. The recovered hydrogen may be partially recycled to the hydrogenation reaction zone as a hydrogen feedstock, and/or may be partially or fully recovered for use in hydrogen-consuming refinery processes such as hydrocracking or hydrotreating processes. Very good. C1,1 aliphatic and aromatic hydrocarbons are recovered as the desired products of this two-stage process. The 03-C5 aliphatic hydrocarbons are recovered as recycled feedstock to the dehydrogenation cyclodimerization zone as described above. Finally, the ethane and ethylene are usually split together and a portion of this ethane/ethylene product stream is combined with the C3-C5 aliphatic recycle stream of the dehydrocyclodimerization reaction step. The remaining ethane/ethylene portion is usually difficult to process (refra
ctory) is combined with the methane stream and recovered as a light byproduct stream.

The raw materials for this dehydrogenation cyclic dimerization reaction zone are 5.0 to 15
, 0 mol % of ethane.

The product stream of the dehydrogenation cyclodimerization reaction step may be sent to any type of separation system known in the art capable of separating and recovering the product and recycle stream. For example, U.S. Patent 4,642,4
No. 02 discloses a method of combining a reaction zone and a product recovery zone in order to completely obtain the xylene produced by the dehydrogenation cyclodimerization reaction. Furthermore, it is expected that the product obtained in the second reaction step of the method described herein can be recovered using any conventionally known method. For example, U.S. Patent 3,53
No. 7,978 and No. 3,574,089 describe the recovery of naphtha, hydrogen-rich recycle gas, and light hydrocarbon streams from the effluent of a catalytic reforming zone. US Patent 3,1
No. 01,261 describes a method for recovering light ends and naphtha from the effluent of a reforming reaction zone. These documents teach separation techniques such as partial concentration, stripping columns, and absorption.

Treatment systems disclosed in the prior art as methods of improving the process, ie, separation efficiency, are also expected to be useful as part of the method of the present invention.

For example, U.S. Patents 4,381,417 and 4.381.4
No. 18 describes a product recovery device for a dehydrogenation process in which the expansion of a gas stream yields a liquid useful as a cooling medium. According to the latter document, the eluate from the reactor is cooled, dried, further cooled, and then passed through the gas-liquid separation zone 28.
I am passing through.

The gas from this separation zone depressurizes (dep) in the turbine 32.
ressurize) and collected in separation zone 34 to provide a cold mixed phase stream.

U.S. Pat. No. 3,838,553 describes the use of low temperatures and high pressures for gas separation and the combination of low temperature separation zones and standard gas-liquid type separation zones. The second part of this document
In the figure, the still high-pressure eluate in the low-temperature separation zone undergoes pressure oscillations (sw
ing) flowing into the absorption band.

Selectively permeable membranes are described in US Pat. No. 4,180,388 and US Pat. No. 4,548,619. These documents also describe various series using circulation and interstage compression.
Various arrangements of two or more membrane separation units in the stream are shown.

Note that the drawings show preferred embodiments of the present invention. Those skilled in the art will appreciate that this process flow diagram includes several heat exchangers, process controls, pumps, fractionator overheads, reboilers, and many other fragmentary processes that are not necessary for an understanding of the invention. It will be appreciated that this is a simplified version with the equipment omitted. It is also easy to understand that the process flow shown in the figure can be modified in many ways without departing from all the basic concepts of the invention. For example, the depiction of the necessary heat exchanger in the figure has been simplified. Therefore, it has been omitted. Those skilled in the art will appreciate that the selection of heat exchange methods employed to provide the necessary heating and cooling at various points in the process involves many variations in implementation. Will. In such a complex process, there are many possibilities for indirect heat exchange between different process streams. Also, depending on the specific arrangement and circumstances of the equipment for this process, other processing units (not shown) may be used.
Preferably, heat exchange is performed on steam, hot oil, or process streams from.

The following examples will serve to illustrate certain embodiments of the invention disclosed herein.

However, this example should not be construed as limiting the scope of the invention as claimed, as those skilled in the art will appreciate that many changes are possible without departing from the spirit of the invention. should not do.

EXAMPLES The following examples are based on pilot plant data obtained from both a complete olefin hydrogenation process and a dehydrogenated cyclodimerization process.

An olefinic hydrocarbon feedstock is fed to a hydrogenation reaction zone from two reactors. The hydrogenation reaction zone had a reactor inlet temperature of 149°C, an absolute inlet pressure of 21.7 atm, and LH3V.
It was operated at approximately 3V5/hr.

The catalyst is divided into two reactors such that the contents in the first and second reactors are respectively 90% and 10% of the total catalyst capacity. A recycle stream is passed from the outlet of the first hydrogenation reactor to the inlet of the same reactor in a ratio of 8.9 moles of recycle hydrocarbon to 1 mole of fresh hydrocarbon feedstock. This hydrogenation catalyst has 0.5 wt% Pd on an aluminum substrate. The flow rates and compositions used in the first hydrogenation reaction zone of this process are shown in Table 1 below.

Feces--1 Flow rate of hydrogenation reaction zone (kg/hr) Component (HCBN raw material) (H2 raw material) (Product)
H2O259,37,5527 c, 3024.02 0 3024.02
IC44844, 6805643, 32=IC4771
,10500,1644]5=N,C4589,668
00,16441513BD 272+,55
0 0.001511507NC4000,
164415 Total (kg/hr) +5818.6 259,
3 16077.9 As can be seen from Table 1, hydrogen was supplied to the hydrogenation reaction zone at a rate such that the amount supplied to the reaction zone was slightly greater than the stoichiometric amount required for hydrogenation of the olefin. Furthermore, it is clear that essentially all the solenoid hydrocarbons are hydrogenated to paraffinic hydrocarbons.

The hydrogenation reaction zone effluent (product) is then directed to the dehydrogenation cyclodimerization reaction zone for conversion to aromatic hydrocarbons. The dehydrogenation cyclic dimerization reaction zone has an average reaction temperature of 540°C and a pressure of 5.1
Operate at atmospheric pressure and LH5V2,35/br of blended raw materials. About 50% of the raw material is crystalline aluminosilicate zeolite.
wt%, phosphorus-containing alumina containing about 22 wt% phosphorus 4
9.0 wt% and about 1 wt% Ga. How to make such a catalyst is described in U.S. Patent No. 4.636.4.
83. The results of the dehydrogenation cyclodimerization process steps are shown in Table 2 below.

The following is a margin dipping knee-2 Flow rate (kg/hr) component of the dehydrogenation cyclodimerization reaction zone
Hydrocarbon raw material Hydrocarbon product H27,6618J C13158,0 C224111,3 =C292,7 C,,3024,0376,7 =C:+24,2n
C47402,665,7 iC45643,33,6 =C,l'S Trace 4.650 CI, Aliphatic 8836.2C
-Aromatic 409.6 Total naphthalenes (kg/hr) 16077, 5 1
6071.0 As is clear from the data in Table 2 above, paraffinic hydrocarbon feedstocks are readily converted to aromatic-containing products in the second reaction stage of the process of the invention.

[Brief explanation of drawings]

FIG. 1 is a process flow sheet illustrating a typical conversion technique according to the present invention. 2...Fresh hydrogen raw material stream 3...Hydrogen raw material stream 4...
Mixed raw material stream 5...Hydrogenation reaction zone 6.8...Hydrogenation reaction zone elution (product) stream 7...Hydrogenation reaction zone circulation flow 10...Dehydrogenation cyclic dimerization reaction zone 12...・Separation strip

Claims (1)

[Claims] 1. (a) Hydrogen and an olefin-containing C_2 to C_5 hydrocarbon feedstock containing a hydrogenation catalyst and C_2 to C_5
(b) through a hydrogenation reaction zone operated at hydrogenation reaction conditions effective to obtain a hydrogenation reaction zone product stream containing at least 50 mole percent less olefinic hydrocarbons than the hydrocarbon feed; At least a portion of the reaction zone product stream is passed through a dehydrogenation cyclodimerization reaction zone containing a dehydrogenation cyclodimerization catalyst and operated under dehydrogenation cyclodimerization conditions to produce hydrogen, methane, ethane, ethylene, C_3. ~C_5 aliphatic hydrocarbons, C_6^
+ producing a dehydrogenation cyclic dimerization reaction zone product containing an aliphatic hydrocarbon and an aromatic hydrocarbon, and then (c) recovering the aromatic product of the dehydrogenation cyclic dimerization reaction zone C_2~ A method for continuous catalytic and selective production of aromatic hydrocarbons from raw materials containing C_5 olefins. 2. A claim characterized in that the dehydrogenation cyclodimerization product fraction containing a portion of ethane and ethylene and a portion of C_3 to C_5 aliphatic hydrocarbons is recycled to the dehydrogenation cyclodimerization reaction zone. Method 1. 3. A claim further characterized in that the hydrogenation reaction zone product stream is passed through a gas-liquid separation zone prior to passing through the dehydrogenation cyclodimerization reaction zone for the purpose of removing unreacted hydrogen from the process stream of the hydrogenation reaction zone. Method of item 1. 4. Process conditions of hydrogenation reaction zone are temperature 50-250℃
, pressure 1 to 50 atm, and hourly space velocity of liquid 0
．． Claim 1 further characterized in that the rate is 1 to 20/hr.
the method of. 5. The process conditions of the dehydrogenation cyclodimerization reaction zone are temperature 35
2. The method of claim 1 further characterized in that the temperature is 0-700<0>C, the pressure is 0.25-20 atmospheres, and the hourly space velocity of the liquid is 0.5-20/hr. 6. Hydrogen: Olefin raw material molar ratio = 1:1-1
2. The method of claim 1 further characterized in that a 0:1 ratio is fed to the hydrogenation zone. 7. Claim 1, further characterized in that the dehydrogenation cyclic dimerization catalyst comprises a crystalline aluminosilicate zeolite component, a phosphorous-containing alumina component, and a component selected from Groups IIB to IVB of the Periodic Table of Elements. the method of.