RU2505515C1 - Method and device for reduction of benzol content in benzine in alkylation with diluted ethylene - Google Patents

Abstract

FIELD: chemistry.

SUBSTANCE: method includes supplying diluted flow of ethylene, which contains between 5 and 50 wt % of ethylene; supplying benzol flow, containing,at least 3 wt % of toluol and, at least, 20 wt % of paraffins; interaction of diluted ethylene flow and benzol flow with catalyst of alkylation, which contains UZM-8; and conversion of, at least, 20% of benzol in raw material flow into alkylbenzol. Invention also relates to device for alkylation of benzol with ethylene.

EFFECT: application of claimed invention makes it possible to eliminate application of reformate by fractionation column, which leads to essential exploitation and financial savings.

10 cl, 3 ex, 3 dwg

Description

FIELD OF THE INVENTION

The scope of the invention relates to a method and apparatus for the alkylation of benzene in a naphtha stream with diluted ethylene. Alkylated product can be used as motor fuel.

State of the art

Dried gas is the common name for the effluent gas stream from a cracked fluidized catalyst unit that contains all gases with ethane boiling points and below. The exhaust gas stream is compressed to remove as many C ₃ and C ₄ gases as possible. Sulfur is also largely absorbed from the off-gas stream in a scrubber that uses an amine absorbent. The remaining stream is known as dried gas FCC. A typical dried gas stream contains 5 to 50 wt.% Ethylene, 10 to 20 wt.% Ethane, 5 to 20 wt.% Hydrogen, 5 to 20 wt.% Nitrogen, 0.05 to 5.0 wt.% Carbon monoxide, 0.1 to 5.0 wt.% Carbon dioxide and less than 0.01 wt.% Hydrogen sulfide and ammonia, the rest is methane.

Currently, the dried FCC gas stream is directed to the burner as fuel gas. The FCC unit, which produces 7,949 kiloliters (50,000 barrels) daily, can burn 181,000 kg (200 tons) of dried gas containing, for example, 36,000 kg (40 tons) of ethylene as fuel, daily. Since there is a significant price difference between fuel gas and motor fuel products or pure ethylene, it seems economically viable to attempt to recover said ethylene. However, the dried gas stream contains impurities that can poison the catalysts and is so diluted that ethylene recovery is not economically feasible using gas recovery systems.

There is a need for the use of diluted ethylene in refinery streams.

Catalytic reforming is a long-established method for the conversion of hydrocarbons used in the oil refining industry to improve the octane quality of hydrocarbon feeds; the primary product of reforming is motor gasoline. In catalytic reforming, the naphtha feed stream is mixed with a recycle stream containing hydrogen and reacted with a catalyst in the reaction zone at temperatures of about 493 ° to 510 ° C (920 ° to 950 ° F) and an average pressure of about 1379 to 3792 kPa (200 up to 550 psi). A typical feed stream for catalytic reforming is an oil fraction known as naphtha and having a starting boiling point of 46 ° C (115 ° F) and an ending boiling point of 204 ° C (400 ° F).

The catalytic reforming process is particularly useful for treating straight-run gasoline fractions composed of relatively high concentrations of naphthenic and substantially unbranched paraffin hydrocarbons that are aromatized through dehydrogenation and / or cyclization reactions. The catalyst “reform” (converts) the molecular structures of the hydrocarbons incorporated into the crude naphtha by removing hydrogen and redistributing the molecular structure so as to improve the octane number of the naphtha. However, an increase in the octane number also lowers the liquid volume of naphtha, since the specific gravity is increased. Due to the variety of compounds in crude naphtha, the actual reactions that occur in catalytic reforming are numerous. However, some of the many products obtained are aryl or aromatic compounds, all of which exhibit high octane numbers. Aryl compounds were obtained depending on the starting compounds, which are controlled at the refinery using the boiling range of the naphtha used and the source of crude oil. A “reformed” product from a catalytic reforming process is generally referred to as reformate and is often separated into two fractions using traditional distillation methods, a light reformate having a boiling range of 46 ° to 121 ° C (115 ° to 250 ° F) and heavy reformate having a boiling range of 121 ° to 204 ° C (250 ° to 400 ° F). The aryl compounds in each fraction thus depend on their boiling points. Lower boiling or lighter aryl compounds, for example benzene, toluene and xylenes, are contained in light reform, and higher boiling aryl compounds are contained in heavy reform.

The concentration of benzene in gasoline is now regulated by the American government. Mobil Source Air Toxic regulation (MSAT II) requires that the average level of benzene in gasoline produced by a refinery specialist be lower than 0.62 vol.%, With a maximum of 1.3 vol.% In gasoline obtained by any one specialist in oil refining. In general, benzene is produced at higher levels than levels in reforming processes and FCC processes. Since reformate and naphtha flows from the FCC facility are the two most significant sources of gasoline in the refining industry, strategies to lower benzene levels should be used.

Currently, benzene is typically sent to the saturation unit to recover benzene to cyclohexane. However, this method uses at least three moles of hydrogen for every mole of converted benzene and there is a loss of octane associated with the conversion of benzene to cyclohexane. Methods are needed to lower the level of benzene in gasoline without losing octane or using hydrogen.

Alkylation of concentrated benzene streams with concentrated ethylene streams is known. Alkylation typically involves the use of pure ethylene streams, as the alkylation catalysts are susceptible to impurities in the feed. In addition, diluted ethylene is little used as an initial oligomerization stream, due to its much lower reactivity with respect to higher olefins. The benzene streams fed to the alkylation reactors are also concentrated because of the fear that heavier aromatic compounds will predominantly alkylate, therefore a significant excess of ethylene must be used before lowering the benzene concentration and producing undesired polyalkylated benzene.

Definitions

The following definitions should be applied through the cited document.

The term “bond” (compound system) means that the flow of a substance is functionally passed between the listed components.

The term “downward flow communication” means that at least a part of the substance flowing to the object connected by the downward flow can flow quickly from the object with which it is communicating.

The term “upstream connection” means that at least a portion of the substance from the object, connected by the upward flow, can functionally flow to the object with which it is communicating.

The term “column” means a distillation column or columns for separating one or more components of different volatilities based on the difference in boiling point. The column may have a boiler in the lower (still) part and a condenser in the upper part. Unless otherwise indicated, each column includes a condenser at the top of the column to condense and supply part of the top stream back to the top of the column as an irrigation and a boiler at the bottom of the column to evaporate and direct part of the bottom stream back to the top of the column . The feed stream supplied to the column may be preheated. The upper pressure is equal to the vapor pressure of the upper part at the [outlet of the column. The temperature of the lower part is equal to the temperature of the liquid at the outlet of the lower part. The lines of the upper part and the lines of the lower part relate to the scheme of lines from the downward (injection) column reflux stream or re-boiling in the column.

As used herein, the term “component rich stream” means that the rich stream leaving the tank has a higher concentration of the component than the feed to the tank.

As used herein, the term “component depleted stream” means that the depleted stream exiting the tank has a lower component concentration than the tank feed.

The term “Cx” should be understood as referring to molecules having the number of carbon atoms represented by the subscript “x”. Similarly, the term “Cx-” refers to molecules that contain less than or equal to x, and preferably x, and less than carbon atoms. The term "Cx +" refers to hydrocarbons with greater than or equal to x, and preferably x or more carbon atoms.

Disclosure of invention

The inventors have found that diluted benzene in streams containing aromatic compounds such as reformate or light naphtha FCC can be alkylated over zeolite catalysts with ethylene in diluted ethylene streams such as a dried FCC gas stream. Heavier hydrocarbons can be separated and mixed in gasoline and diesel tanks (parks). The inventors have found that most zeolite catalysts that are suitable for the alkylation of benzene with light olefins are rapidly deactivated in dilute ethylene streams. Neither the diluted nature of ethylene nor the impurities present substantially affect the catalyst containing UZM-8. In addition, the conversion of benzene during alkylation with ethylene is as high as that of toluene and higher than that of higher aromatic hydrocarbons with the UZM-8 catalyst. Therefore, diluted benzene in a naphtha stream such as reformate can be alkylated with diluted ethylene in an FCC dried gas stream to provide a liquid fuel product that is reduced in benzene concentration and can easily be separated from unreacted gas stream. Unreacted gas can then be burned as fuel gas, but with more expensive ethylene recovered as the gasoline series of hydrocarbons.

Advantageously, the method and apparatus can enable the use of dilute ethylene in the stream and in the presence of impurities in the feed stream, which can be catalytic poisons.

Advantageously, the method and apparatus can enable a decrease in the diluted benzene concentration in the naphtha stream without using hydrogen or a decrease in the liquid volume or octane number of the naphtha stream even in the presence of heavier aromatic hydrocarbons.

In one embodiment, the invention relates to a method for alkylating benzene with ethylene, comprising providing a diluted ethylene stream containing between 5 and 50 wt.% Ethylene. A benzene stream containing at least 3 wt.% Toluene and at least 20 wt.% Paraffins is provided. The diluted ethylene stream and benzene stream are reacted with an alkylation catalyst containing UZM-8. Finally, at least 20% of the benzene in the feed stream is converted to alkylbenzene.

In an additional embodiment, the invention relates to a method for alkylating benzene with ethylene, comprising reacting a cracking catalyst with a hydrocarbon feed stream to crack hydrocarbons to hydrocarbon cracking products having a lower molecular weight and generating coke on the cracking catalyst, which promotes the formation of a cracking catalyst, coated with coke. The coke coated cracking catalyst is separated from the cracked products. Oxygen is added to the coke coated cracking catalyst, and coke on the coke coated cracking catalyst is burned with oxygen to regenerate the cracking catalyst. The cracked products are separated to obtain a dilute ethylene stream containing between 5 and 50 wt.% Ethylene. A benzene stream containing at least 3 wt.% Toluene and at least 20 wt.% Paraffins is provided. The diluted ethylene stream and benzene stream are reacted with an alkylation catalyst containing UZM-8. At least 20% of the benzene in the feed stream is converted to alkylbenzene.

In a further embodiment, the invention relates to a method for alkylating benzene with ethylene, comprising providing a diluted ethylene stream containing between 5 and 50 wt.% Ethylene. The naphtha stream is reacted with a reforming catalyst to provide a reformate benzene stream containing at least 3% by weight of toluene and at least 20% by weight of paraffins. The diluted ethylene stream and benzene stream are reacted with an alkylation catalyst containing UZM-8. At least 20% of the benzene in the feed stream is converted to alkylbenzene.

In one embodiment, the invention relates to a device for the alkylation of benzene with ethylene, containing a cracking reactor with a fluidized catalyst for reacting the cracking catalyst with a hydrocarbon feed stream to crack the hydrocarbon feed to cracking products having a lower molecular weight and to form coke on the cracking catalyst , which contributes to the formation of a cracking catalyst coated with coke. The device also includes a regenerator for burning coke from said coke-coated cracking catalyst by reacting with oxygen, and a separator, in combination with a fluidized catalyst cracking reactor, for separating C ₃ hydrocarbons from C ₂ hydrocarbons to provide a diluted ethylene stream. The device also comprises a reforming reactor for reacting the naphtha stream with the reforming catalyst to obtain a reformate stream. Finally, the apparatus comprises an alkylation reactor coupled (in connection) with a separator and a reforming reactor for alkylating benzene in a reformate stream with ethylene in a dilute ethylene stream on a stationary alkylation catalyst to heavier alkylbenzene hydrocarbons. A distillation column is absent in the communication system between the reforming reactor and the alkylation reactor. In an alternative embodiment, the distillation column is absent in the communication system between the separator and the alkylation reactor. In a further embodiment, the absorber, in conjunction with the outlet of the FCC reactor, provides an exhaust gas stream containing a diluted ethylene stream and there is no distillation column in the communication system between the reforming reactor and the alkylation reactor, and between the absorber and the alkylation reactor.

Additional features and advantages of the present invention will be apparent from the description of the invention, the figures and claims recited herein.

Brief Description of the Figures

Figure 1 is a schematic representation of an FCC unit, a reforming unit, and an alkylation unit.

Figure 2 is a graph of conversion and temperature versus time.

Figure 3 is a graph of conversion versus time.

The implementation of the invention

The present invention can be applied to any hydrocarbon stream containing ethylene and, preferably, a diluted portion of ethylene. A suitable diluted ethylene stream may typically contain between 5 and 50 wt.% Ethylene. The FCC dried gas stream is a suitable diluted ethylene stream. Other diluted ethylene streams, such as dried coke oven gas streams, may also be used in the present invention. Since the present invention is particularly suitable for dried FCC gas, the subject matter will be described with respect to the use of ethylene from FCC gas streams.

The present invention can also be applied to any hydrocarbon stream containing benzene and, preferably, a diluted portion of benzene. A suitable benzene stream may typically contain between 1 and 50 wt.% Benzene, at least 3 wt.% Toluene and at least 20 wt.% Paraffins. The reformate stream is a suitable benzene stream. Other benzene streams, such as FCC aromatic naphtha streams, may also be used in the present invention. Since the present invention is particularly suitable for a reformate stream, the subject matter of the application will be described with respect to the use of benzene from a reformate stream.

Now returning to FIG. 1, where identical numbers denote identical components, FIG. 1 illustrates a refinery complex 6, which typically includes an FCC unit 10 with a product recovery section 90, a reforming unit 200, and an alkylation unit 300.

The FCC 10 includes a reactor 12 and a catalyst regenerator 14. The process parameters typically include a cracking reaction temperature of 400 ° to 600 ° C and a catalyst regeneration temperature of 500 ° to 900 ° C. Both cracking and regeneration are conducted at absolute pressures below 506 kPa (72.5 psi).

Figure 1 shows a conventional FCC reactor 12 in which a heavy hydrocarbon feed or a crude oil stream 16 is brought into contact from a distributor with a regenerated cracking catalyst coming from a cylindrical tank of regenerated catalyst 18. Said interaction can be carried out in a narrow distribution tower 20 extending upward to the bottom of the reaction vessel 22. The interaction of the feedstock and the catalyst is fluidized by gas from the fluidization line 24. In one embodiment In an embodiment of the invention, the heat from the catalyst evaporates the hydrocarbon feedstock or oil, and then the hydrocarbon feedstock is cracked to hydrocarbon products with a lower molecular weight in the presence of a catalyst, which together with the stream are passed through separation column 20 to reaction vessel 22. Constant side reactions occur in separation column 20, which deposits coke deposits on the catalyst, which reduces catalytic activity. The light hydrocarbon cracking products are then separated from the coke-coated cracking catalyst using cyclone separators, which may include a first stage separator 26 and one or two-stage cyclones 28 in the reaction tank 22. Gaseous cracking products leave the reaction tank 22 through an outlet 31 through a pipe 32 to go to the downstream product recovery section 90. A spent or coke-coated catalyst requires regeneration for later use. The coke-coated cracking catalyst, after being separated from the gaseous hydrocarbon products, is fed to a product withdrawal section (evaporation section) 34, in which a stream is introduced through a nozzle to purify any remaining hydrocarbon vapor. After the product selection operation, the coke coated catalyst is sent to catalyst regenerator 14 through a cylindrical spent catalyst tank 36.

Figure 1 depicts a regenerator 14, known as a combustion chamber. However, other types of regenerators are suitable. In the catalyst regenerator 14, a stream of oxygen-containing gas, such as air, is introduced through an air control device 38 to interact with the coke-coated catalyst. Coke is burned on a coke coated catalyst to provide regenerated catalyst and flue gas. The catalyst regeneration method adds a significant amount of heat to the catalyst, providing energy for the shear of the endothermic cracking reactions occurring in the separation column of the reactor 20. The catalyst and air are directed upward along the combustion chamber of the separation column 40 located inside the catalyst regenerator 14 and, after regeneration, are initially separated , unloading through the unloading elevator 42. Additional extraction of the regenerated catalyst and flue gas leaving the unloading The riser 42 is achieved using one or two-stage cyclone separators 44, 46, respectively, inside the catalyst regenerator 14. The catalyst separated from the flue gas is distributed through the downpipes from the cyclones 44, 46, while the hot flue gas is relatively lighter in the catalyst in series leaves the cyclones 44, 46 and leaves the reservoir of the regenerator 14 through the outlet of the flue gas 47 through the line of the flue gas 48. The regenerated catalyst is returned back to the separation column 20 through cylin regenerated catalyst reservoir 18. As a result of coke burnout, flue gas vapors leaving the upper part of catalyst regenerator 14 via line 48 contain CO, CO ₂ , N ₂ and H ₂ O, together with smaller amounts of other types of constituents.

The product recovery section 90 is downstream to the product outlet 31. In the product recovery section 90, the FCC product gas is sent via line 32 to the bottom of the main distillation column FCC 92. The main column 92 is downstream to the product outlet 31. Several product fractions FCC can be separated and taken from the main column, including the stream of heavy suspension in the petroleum product from the bottom product through line 93, the heavy recycle gas oil stream through line 94, the light recycle gas oil through line 95 taken from the outlet 95a and a heavy naphtha stream along line 96 taken from the outlet 96a. Any or all (flows) of the lines (pipelines) 93-96 can be cooled and pumped back to the main column 92 to cool the main column, usually at a higher location. Gasoline and gaseous light hydrocarbons are removed through the overhead of line 97 from the main column 92 and condensed before being fed to the head column collector 99. The main column collector 99 is connected downstream to the product outlet 31 and the main column 92 is connected upstream to the main column collector 99 .

The water stream is removed from the sump in collection tank 99. In addition, the condensed light naphtha stream is removed through line 101, while the overhead stream is removed through line 102. The overhead stream from line 102 contains gaseous light hydrocarbons, which may contain a dilute ethylene stream . A portion of the condensed light naphtha stream through line 101 can be returned as irrigation to the main distillation column 92 via line 105 leaving the final condensed stream of light naphtha through line 103. The streams along lines 101 and 102 may enter the vapor collection section 120 of the recovery section product 90.

The vapor storage section 120 is presented as an absorption system, but any vapor recovery system can be used, including a refrigeration system. In order to obtain sufficient separation of the light gaseous components, the gaseous stream through line 102 is compressed in the compressor 104. More than a single stage compressor can be used and, as a rule, two-stage compression is used to compress the gaseous stream through line 102 to between 1.2 MPa to 2 , 1 MPa (excess) (180 to 300 psi).

Three-stage compression may be advantageous to provide additional pressure of at least 3.4 MPa (gauge) (500 psi).

To the compressed light hydrocarbon stream through line 106, a C ₃ + hydrocarbon liquid bottoms product stream of the first stage absorber is connected via line 107 and C ₂ is a hydrocarbon stream along line 108 of the upper part of the stripping column, cooled and directed to a high-pressure collector 110. The water stream from the collector 110 may be directed to the main column 99 collector. The gaseous first FCC product stream through line 112, containing the diluted ethylene stream, is sent to a unit that performs separation between C ₃ + and C ₂ hydrocarbons, which in yes This embodiment of the invention is a first stage absorber 114. In a first stage absorber 114, diluted ethylene, a stream of a first (first stream) FCC product is reacted with a stream of a second (second stream) FCC product containing unstable gasoline from a main column tank 99 line 103 to perform a separation between C ₃ + and C ₂ - hydrocarbons. A separator for separating C ₃ from C ₂ hydrocarbons, which may be the first stage absorber 114, is in the downstream connection to the main column collector 99. The liquid C ₃ + stream through line 107 is returned to line 106 before cooling. The overhead line of the first stage absorber 114, containing dried gas, mainly C ₂ hydrocarbons with hydrogen sulfide, ammonia, carbon oxides and hydrogen, is removed in the primary off-gas stream via line 116 to the content of the diluted ethylene stream. However, in order to further concentrate the ethylene stream and recover the heavier components of line 116, optionally, they can be directed to a second section, which performs the separation between C ₃ + and C ₂ - hydrocarbons, which in the above embodiment is an absorber of the second stage 118. In the absorber of the second stage, the circulating stream of light recycle gas oil through line 121, diverted from line 95, absorbs most of the remaining components of C ₅ + and some components of C ₃ -C ₄ in the primary stream leaving its gas. The second stage absorber 118 is connected downstream to the first stage absorber 114. A light recycle gas oil from the lower part of the second stage absorber 118 through line 119 enriched with a C ₃ + component is returned to the main column 92 through circulation for line 95. The overhead head of the second stage absorber 118 containing dried gas, mainly C ₂ - hydrocarbons with hydrogen sulfide, ammonia, carbon oxides and hydrogen, are removed in the stream of the second stage of the exhaust gas through line 122, until the content of the diluted stream of ethylene. Both the absorption column 114 and the absorption column 118 do not have a condenser or boiler, but circulating cooling schemes can be used.

The liquid from the high-pressure collector 110 is sent via line 124 to the distillation column 126. Most of C ₂ is removed in the overhead of the distillation column 126 and returned to the stream of line 106 through the overhead of line 108. The distillation column 126 does not contain a condenser, but receives a cooled liquid food on line 124. The stream of liquid bottoms product from the distillation column 126 is sent to the butane fraction 130 distillation column through line 128. The overhead stream (distillate stream) along line 132 from the butane fraction distillation column contains C ₃ -C ₄ olefins the second product, while the bottoms product via line 134 containing stabilized gasoline can be further processed and sent to the gasoline collection tank. In one embodiment of the invention, the bottoms stream via line 134 may be directed to the naphtha distillation column 140. Light naphtha may be recovered in line overhead via line 142 from the naphtha distillation column 140 with a heavy naphtha stream recovered as bottoms from line 144. Aromatic naphtha containing benzene can be selected as the middle fraction along line 146, in which case the naphtha distillation column may be a separation partition column. One or both of the light and heavy naphtha streams along lines 142 and 144 can be directed to the gasoline distribution system 148. Figure 1 shows only the direction of the heavy naphtha flow along line 144 to the gasoline distribution system 148, since the light naphtha 142 can be directed on technological separation from heavy naphtha.

The diluted ethylene stream in the present invention may contain a dried FCC gas stream containing between 5 and 50 wt.% Ethylene, and preferably 10 to 30 wt.% Ethylene. The predominant component will typically be methane in a dilute ethylene stream at a concentration between 25 and 55 wt.%, With ethane present essentially, typically between 5 and 45 wt.%. Between 1 and 25 wt.% And typically 5 to 20 wt.% Of hydrogen and nitrogen each may be present in a dilute ethylene stream. Saturation levels with water may also be present in the diluted ethylene stream. The diluted ethylene stream in the overhead line 116 can contain no more than 3 wt.% And, accordingly, no more than 1 wt.% Propylene and, as a rule, no more than 25 wt.% And, accordingly, no more than 15 wt.% C ₃ + components. If a second stage absorber 118 is used, the diluted ethylene stream in overhead 122 may contain no more than 5 wt.% C ₃ +, typically with a propylene content of less than 0.5 wt.%. In addition to hydrogen, other impurities, such as hydrogen sulfide, ammonia, carbon oxides and acetylene, may also be present in the diluted ethylene stream.

Impurities in the dried gas ethylene stream may poison the alkylation catalyst. Carbon dioxide and ammonia can attack the acid sites of the catalyst. It is known that hydrogen sulfide periodically deactivates the zeolite catalyst. Acetylene can polymerize and clog the catalyst or equipment.

A first stage off-gas stream through line 116 or a second-stage off-gas stream through line 122 containing a diluted ethylene stream can be introduced into an optional amine absorber unit 150 to remove hydrogen sulfide to a lower concentration. An aqueous amine solution, such as monoethanolamine or diethanolamine, is introduced through lines 152 into the absorber 150 and reacted with a flow of exhaust gas to absorb hydrogen sulfide, and the aqueous amine absorption solution enriched in hydrogen sulfide is removed from the absorption zone 150 via line 154 and removed, and possibly further processed.

Optionally, a hydrogen sulfide-depleted, amine-treated diluted ethylene stream through line 156 can be introduced into an optional water-wash unit 160 to remove residual amine carried away from the amine absorber 150 and lower the concentration of ammonia and carbon dioxide in the diluted ethylene stream through line 156. Water is fed to rinsing with water along line 162. Water along line 162 is typically slightly acidified to enhance the uptake of basic molecules such as amine. The water stream through line 164, enriched with amine and potentially ammonia and carbon dioxide, leaves the water wash unit 160 and can be further processed.

Optionally treated with an amine, optionally washed with water, the diluted ethylene stream through line 166 can then be treated in an optional adsorber to protect from impurities 170 to remove one or more of such impurities such as carbon monoxide, hydrogen sulfide and ammonia, lowering to an even lower level concentration. The anti-impurity adsorber 170 may contain an adsorbent to adsorb impurities, such as ammonia, which can poison the alkylation catalyst. The anti-impurity adsorber 170 may comprise a plurality of adsorbents for adsorbing more than one type of impurity. A typical adsorbent for adsorption of hydrogen sulfide is ADS-12, for adsorption of carbon monoxide is ADS-106, and for adsorption of ammonia is TOP MOLSIV 3A, all available from UOP, LLC. Adsorbents can be mixed in one layer or can be distributed in successive layers.

Optionally treated with an amine, optionally washed with water and optionally through an adsorption step, line 172 can be dried in a dryer 174 to reduce the amount of water to below 500 ppm. Water can adversely affect the alkylation catalyst.

The diluted ethylene stream through line 176, optionally treated with an amine, optionally washed with water, optionally undergone an adsorption step and optionally dried, typically contains at least one of the following impurities in a concentration of 0.05 mass% and up to 5, 0 wt.% Carbon monoxide, and / or 0.1 wt.% And up to 5.0 wt.% Carbon dioxide, and / or at least 1 part per million and up to 500 parts per million of hydrogen sulfide, and / or at least 1 and up to 500 ppm ammonia, and / or at least 5 and up to 20 wt.% hydrogen a. The types of impurities present and their concentration may vary depending on the treatment and the origin of the diluted ethylene stream.

Line 176 carries the diluted ethylene stream to compressor 180 if necessary to compress to the pressure of the alkylation reactor. Compressor 180 is connected downstream to main column 92, product recovery section 90 and product outlet 31. Compressor 180 may comprise one or two stages with interstage cooling. A heater may be required to bring the compressed stream to the reaction temperature. The compressed diluted ethylene gas stream is led through line 182 to the alkylation unit 300. Line 182 leads the diluted ethylene stream to the installation of the alkylation reactor 300. The installation of the alkylation reactor 300 may be downstream to compressor 180 and / or the first or second separators for separating C ₃ hydrocarbons from C ₂ hydrocarbons, which may be absorbers of the first or second stage 114 and 118, respectively. In one embodiment of the invention, no distillation unit is downstream between the first or second separator for separating C ₃ hydrocarbons from C ₂ hydrocarbons and the alkylation reactor 320. Therefore, in one embodiment of the invention, no distillation unit is connected between the absorber of the first stage 114 or a second stage absorber 118 and an alkylation reactor 320. In the above embodiment, the diluted ethylene stream may be separated m on the basis of adsorption or absorption, but no fractionation based on the difference in boiling points is carried out for a dilute ethylene stream, which may contain exhaust gas of the first stage or second stage between the absorber of the first stage 114 and / or the absorber of the second stage 118 and the alkylation reactor 320 The above embodiment of the invention is in contrast with the traditional conviction, which believes that the diluted ethylene stream of dried gas, which needs to be divided into fractions, for example p in the demethanizer column to remove lighter components before ethylene is unlikely to survive the alkylation reaction with benzene. The elimination of the methane stripping column leads to substantial operational and capital savings.

Returning to the reformer 200, the naphtha feed stream through line 202 is mixed with the hydrogen containing stream from line 204, heated, and reacted with a catalyst in reforming reactor 210 to obtain reformate. Preferably, the reforming reactor 210 is a moving-bed reactor that receives the regenerated catalyst via line 220 and discharges the spent catalyst through line 222 to a regeneration zone 230 provided with a fluidized inert gas. The catalyst flows from the upper part to the lower part of the reactor, composed of several zones 210, first passing through the reduction zone 224, in which the hydrogen enriched gas from line 226 contacts and reduces oxidized catalytic particles. From there, the catalyst flows through multiple reaction zones in which the starting naphtha interacts with the catalyst particles. Reformation reactor 210 may comprise a reactor configuration composed of several zones, which may include multiple reaction zones. Each reaction zone may contain a catalyst layer in a folded reactor 210 to allow a continuous or batch flow of catalyst particles from the upper reaction zone 212 through the second and third reaction zones 214 and 216, respectively, to the last zone 218. The effluent from the first through the third reaction zone 212 -216 may be withdrawn, heated, and returned to subsequent reaction zones 214-218, respectively. More or less catalytic beds may be used. The reformate product stream may be withdrawn via line 232 from the last reaction zone 218. The lower storage chamber 234 at the bottom, composed of several zones, of the reactor 210 receives spent catalyst. A washing liquid, preferably containing hydrogen, enters the lower storage chamber 234 via line 236, at a speed that flushes hydrocarbons from the catalyst particles in the lower storage chamber 234.

A typical feed stream for catalytic reforming is an oil fraction known as naphtha and having an initial boiling point of 46 ° C (115 ° F) and an end point of 204 ° C (400 ° F). The catalytic reforming process is particularly applicable to the processing of straight-run naphtha, consisting of relatively high concentrations of naphthenic and essentially unbranched paraffin hydrocarbons, which are aromatized through dehydrogenation and / or cyclization reactions. During reforming, dehydrogenation of cyclohexanes and dehydroisomerization of alkylcyclopentanes gives aromatics, dehydrogenation of paraffins gives olefins, dehydrocyclization of paraffins and olefins gives aromatics, isomerization of n-paraffins and alkylcycloparaffins gives cyclohexanes, substituted aromatics are isomerized products and are hydromerized.

The catalytic reforming reaction is usually effective in the presence of a portion of the catalyst consisting of one or more noble metals of group VIII, such as platinum, iridium, rhodium, palladium, and halogen combined with a porous carrier, such as a refractory inorganic oxide. Halogen is usually chloride. Alumina is a widely used carrier. Preferred alumina compounds are known as gamma, this and theta alumina, with gamma and this alumina giving the best results. An important property related to the working embodiment of the catalyst is the surface area of the carrier. Preferably, the carrier should have a surface area of from 100 to 500 m ² / g. Particles are usually spheroidal and have a diameter of 1/16th to 1/8th of an inch (1.5-3.1 mm), although particles can be as large as 1/4 of an inch (6.35 mm) . During the reforming reaction period, the catalyst particles become deactivated as a result of mechanisms, such as the deposition of coke on the particles, that is, after a period of time in use, the ability of the catalytic particles to promote the reforming reaction decreases to the point at which the catalyst is no longer useful. The catalyst must be regenerated or regenerated before it can be reused in the reforming process.

Regeneration zone 230 regenerates the catalyst and recycles through line 220. Spent catalyst particles containing coke deposits flow from the bottom storage chamber 234 of the folded reactor 210 through a discharge pipe 222 to a drain tank 240 in a regeneration zone 230. The drain tank 240 may comprise two sections. In the upper section 242, the decantation liquid enters the upper section 242 at a speed that separates broken or chipped catalyst particles and small catalyst particles from the whole catalyst particles that exit the bottom of the outlet tank 240. Chipped catalyst particles and small particles leave the upper section 242 with a decantation liquid, which can be filtered and recycled to the upper section 242. In the lower section 244, the decanted catalyst is reacted with cooled flue gas Zoom from line 246 to adsorb chlorine and hydrogen chloride on the catalyst. The flue gas on line 246 can be heat exchanged with dechlorinated flue gas on line 248 from the bottom section 244, until it is cooled under adsorption before being fed to the bottom section 244. The catalyst passes from the bottom section 244 to the catalyst regenerator 250 through line 252.

The catalyst regenerator 250 comprises a combustion section 254 and a section for adjusting to the required parameters 256 (conditioning section). In the combustion section, the catalyst moves down into the inner annular chamber 258, which may contain coaxial meshes (filters) impervious to the passage of the catalyst. The dechlorinated flue gas returned via line 248 is heated by heat exchange with the flue gas via line 246 and is supplied to an external annular chamber 260. Recycled flue gas passes from an external annular chamber 260 through an external mesh to an internal annular chamber 258 to heat the catalyst therein . The gases used in section 256, containing oxygen and chlorine, rising from the conditioning section also pass into the inner annular chamber 258 to burn coke deposits from the catalyst. Hot combustion products enter the inner coaxial grid in the inner pipe 245 and leave the combustion section 254 along line 246. The calcined catalyst passes into the conditioning section 256 in the central part 262 with an annular partition inside. Air, which can be enriched with oxygen and mixed with chlorine, is passed through line 264 to the lower annular chamber 266 outside the annular partition, and passes into the central part 262, in contact with the catalyst, until the metal disperses on the catalyst. The regenerated catalyst exits conditioning section 256 and is sent via line 220 back to reforming reactor 210 provided with fluidized gas.

The reformate stripping column 270 may be downstream of the reforming reactor 210 and connected between the reforming reactor 210 and the alkylation reactor via line 232. In the illustrated embodiment, the control valve on line 278 is at least partially closed and the control valve on line 272 at least partially open, so line 272 provides the flow of reformate product through line 232 to the distillate column of reformat 270. The separation into fractions produces an easy stream of reformate in the head e line 274, which then feeds the benzene stream through an open control valve on it through line 304 and the alkylation reactor 320. The light reformate benzene stream may have 20 to 50 wt.% benzene, with at least 20 wt.% paraffins and the rest, at least 3 wt.% toluene. In one embodiment of the invention, the light stream of reformate benzene has a higher concentration of paraffins than benzene. The benzene stream preferably contains at least 4.0 wt.% Benzene. The heavy reformate bottoms product stream through line 276, leaving the lower half of the reformate 270 stripping column, can bypass (bypass) the alkylation reactor 320 and flow to the gasoline distribution system 148, without being connected to the alkylation reactor 320.

Alternatively, an entire reformate stream through line 232 may bypass reformat 270 or the reformer 200 may not include reformat 270 at all. In the above embodiment, the control valve on lines 272 and 274 is closed and the control valve on line 278 is at least partially open, so line 232 is capable of supplying the entire reformate product stream on line 304 and alkylation reactor 320. In the above In an embodiment of the invention, the whole reformate stream may be directed to the alkylation unit 300 via line 278. The reformate may not be completely turned off. In this embodiment, there is no distillation column in the connection between the reforming reactor 210 and the alkylation reactor 320. In this embodiment, the whole reformate stream can be separated based on adsorption or absorption, however, there is no fractionation based on the difference in boiling point carried out in a reformate stream between the reforming reactor 210 and the alkylation reactor 320. The above embodiment of the invention is in contrast to traditional CB belief, which states that a reformate stream that needs separation into fractions such as reformate stripping column for removal of heavy aromatics to benzene hardly transfer alkylation reaction olefins. The elimination of the reformate distillation column leads to substantial operational and capital savings.

The whole reformate stream may contain 1 to 10 wt.% Benzene, 3 to 30 wt.% Toluene, the rest is at least 20 wt.% Paraffins and heavier aromatic compounds. The whole reformate stream contains a higher concentration of aromatic compounds with a molecular weight greater than benzene than a concentration of benzene. Especially, the whole reformate stream may have aromatic compounds with eight carbon atoms in a higher concentration than the concentration of benzene. In addition, the whole reformate stream may have a higher concentration of paraffins than benzene.

Line 304 feeds the benzene stream to the alkylation unit 300. The alkylation unit 300 may be downstream of reformate 270. The alkylation reactor 320 preferably comprises a fixed catalyst bed 342 and may contain multiple catalyst beds 342-346. The catalyst is preferably a UZM-8 zeolite bound to alumina.

In one embodiment of the invention, diolefins in a bottoms stream in a gaseous diluted ethylene stream via line 182 may optionally be delivered to first react with a selective hydrogenation catalyst in a selective hydrogenation zone 310 to selectively convert to saturated diolefins without completely saturating them to paraffins. Suitable conditions for performing the selective hydrogenation process include the passage of a dilute ethylene stream through line 302 in the gas phase and hydrogen through line 308, with a molar ratio of 0.5 to 5 moles of hydrogen for each mole of diolefin with respect to the catalyst containing at least one metal, selected from the group formed by nickel, palladium and platinum, deposited on a support such as alumina at a temperature of from 20 ° to 200 ° C (68 ° to 392 ° F), a pressure of 689 to 3447 kPa (excess) (100) up to 500 psi), and surround speed awns from 0.5 to 10 h ^-1 . Two or more reaction zones may be used, although only one is shown. Each reaction zone may use recycling (not shown) of the reactor effluent into the reactor inlet at a ratio of recycling to ethylene feed stream ranging from 0 to 20. The residual diolefin content in such a process can be in the range of 1 to 100 ppm, depending from the rigidity of the execution conditions.

A diluted ethylene stream from the selective hydrogenation reactor 310 via line 312 is introduced into the alkylation reactor 320. In one aspect, a drying chamber (not shown) on line 312 can be used to remove water to low concentrations that are unlikely to affect the alkylation catalyst. Other protective layers are also contemplated to remove catalytic poisons, such as removing ammonia or amines up to 1 to 500 ppm. One of the catalyst layers can serve as a protective layer to remove water and catalytic poisons. If a drying chamber or a protective layer is used on the ethylene supply side of the alkylation unit 300, then one, part or both of the drying chamber 176 and the adsorbent layer 170 can be excluded from the product recovery section 90. The diluted ethylene stream through line 312 has essentially that the same composition as on line 114 or 122, with the exception of removed impurities. The diluted ethylene stream through line 312, optionally treated with an amine, optionally washed with water, optionally through an adsorption step, and optionally dried, can typically have at least one of the following impurities in concentrations: 0.05 wt.% And up to 5.0 wt.% carbon monoxide, and / or 0.1 wt.% and up to 5.0 wt.% carbon dioxide, and / or at least 1 part per million and up to 500 parts per million of hydrogen sulfide and / or at least at least 1 and up to 500 ppm of ammonia and amines, and / or at least 5 and up to 20 wt. .% Hydrogen.

Although transalkylation reactions may occur in alkylation reactor 320, alkylation reactions are preferred. An alkylation reactor is shown as an upflow reactor, however a downflow reactor may also be suitable. The diluted ethylene stream through line 312 is introduced into the alkylation reactor 320 via several lines 322, 324 and 326 to the preceding catalyst beds of spaces 332, 334 and 336 before entering the catalyst beds 342, 344 and 346, respectively. Catalyst layers 342, 344 and 346 contain an alkylation catalyst to alkylate benzene with ethylene to produce ethylbenzene and toluene to produce ethyl toluene. Other alkylation reactions occur to produce alkylbenzenes and alkyl aromatic compounds. The benzene liquid stream through line 304 is supplied to the alkylation reactor 320, where the stream first absorbs the diluted ethylene stream from line 326 in the preceding layer 336 and together enters the catalytic layer 346. The aromatic reformate feed stream through line 304 can also receive an FCC aromatic stream from line 146 before entering the alkylation reactor 320. Alternatively, if the naphtha is distilled off with only two streams, one of the streams may be supplied by aromatic naphtha line 304, which preferably is a light naphtha stream on line 142. The gaseous ethylene dissolved in the liquid reformate stream to alkylate aromatic rings.

The effluent from the catalytic layer 346 is mixed with fresh dried gas containing diluted ethylene from line 324 in the preceding layer space 334 and is introduced together into the catalytic layer 344. The effluent from catalytic layer 344 is mixed with fresh dried gas containing diluted ethylene from line 322 in the previous layer space 332 and are introduced together into the catalytic bed 342. The process is repeated for a number of layers in the alkylation reactor 300. Although three catalyst layers are shown in the alkylation reactor 300, more or less is suitable for Oev and additional reactors. The alkylation effluent from alkylation reactor 320 is transported to effluent line 352. Heat exchanger 354 can cool the Effluent on line 352 to a desired temperature. The effluent pressure of the alkylation reactor can be reduced by passing through a pressure control valve or increased by passing through a pump, none of which are shown.

A preferred alkylation catalyst of the present invention is described below. A preferred alkylation catalyst contains zeolite UZM-8. One of the components of the catalyst carrier used in the present invention is alumina. The alumina source may be any of various aqueous aluminas or alumina gels, such as boehmite or pseudoboehmite alpha alumina monohydrate, gibbsite structure alpha alumina trihydrate, beyerite structure beta alumina trihydrate, and the like. A particularly preferred alumina is available from the Sasol North America Alumina Product Group under the trademark Catapal. This material is an extremely high pure alpha alumina monohydrate (pseudoboehmite), after calcining at high temperature it has been shown to produce highly pure gamma alumina. The zeolite component of the catalyst is UZM-8 described in US 6756030.

A suitable alkylation catalyst is obtained by mixing proportional volumes of UZM-8 and alumina to achieve the desired zeolite to alumina ratio. In one embodiment, 70 wt.% UZM-8 and 30 wt.% Alumina powder provides a suitable carrier. In one embodiment of the invention, other weight ratios than 70 to 30 UZM-8 to alumina, ranging from 90 wt.% UZM-8 in content to 20 wt.% UZM-8 in content, the rest alumina, may be suitable.

A monobasic acid, such as nitric acid or formic acid, can be added to the mixture in an aqueous solution to plasticize the alumina in the binder. Additional water may be added to the mixture to provide sufficient moisture to form a pasty mass with sufficient consistency, which is formed by extrusion or spray dried.

The paste or pasty mass can be prepared in the form of particles with a certain shape, with a preferred method in which the pasty mass is extruded through a preform having holes in it of the desired size and shape, after which the extruded material is separated into extrudates of the desired length and dried. An additional calcination step may be used to add additional length to the extrudate. Basically, the calcination is carried out in a stream of dry air at a temperature of from 260 ° C (500 ° F) to 815 ° C (1500 ° F).

The extruded particles can have any suitable cross-sectional shape, namely, symmetric or asymmetric, but often the majority has a symmetrical cross-sectional shape, preferably spherical, cylindrical or "polylobal" -shaped. The diameter of the cross section of the particles can be up to 40 microns; however, typically the diameter is 0.635 mm (0.25 inches) to 12.7 mm (0.5 inches), preferably 0.79 mm (1/32 inches) to 6.35 mm (0.25 inches), and most preferably 0.06 mm (1/24 in.) To 4.23 mm (1/6 in.). Among the preferred catalyst configurations, cross-sectional shapes similar to a three-leaf clover are presented, as shown, for example, in US Pat. having a diameter between 0.51 mm (0.02 inches) and 1.27 mm (0.05 inches). Other preferred particles are particles having “quadralobab-shaped cross-sections including asymmetric shapes and symmetrical shapes, such as in US Pat. No. 4,028,227 in FIG. 10.

Since the reaction is carried out, at least under partial conditions of the liquid phase, the reaction pressure can be adjusted to maintain ethylene, at least partially in the liquid phase. Ethylene in the gas phase may also be suitable. The pressures can vary within a wide range from 101 to 13172 kPa (gauge) (1 to 1900 psi). From a practical point of view, the pressure is usually in the range between 1379 and 6985 kPa (excess) (200 to 1000 psi), but in most cases, it is in the range between 2069 and 4137 kPa (300 and 600 psi inch). A temperature range suitable for the alkylation of benzene with ethylene is between 100 ° and 300 ° C. The ratio of aromatic compounds to ethylene should be between 1:10 and up to 10: 1, with a preferred ratio of 0.5: 1.0.

The diluted ethylene stream can be fed to the alkylation reactor 320 in the gas phase, at a temperature between 100 ° and 300 ° C. The reaction takes place mainly in the liquid phase at a space velocity (WHSV) of 0.01 to 10 h ^-1 based on ethylene. Ethylene gas can be absorbed in the liquid benzene stream so that alkylation occurs. The inventors unexpectedly found that despite the presence of impurities in the feed stream that poison the catalyst and diluted ethylene in the feed stream, that at least 40 wt.% And up to 75 wt.% Ethylene in the feed stream alkylates aromatic rings to convert to heavier alkyl aromatic hydrocarbons.

The diluted benzene stream can be fed to the alkylation reactor 320 in the liquid phase, at a temperature between 100 ° and 300 ° C. The reaction takes place mainly in the liquid phase at a space velocity (WHSV) of 0.01 to 40 h ^-1 based on benzene. The inventors unexpectedly found that despite the presence of heavier aromatic compounds and paraffins in the feed stream, that the dilution of benzene in the feed stream is at least 20 wt.%, Respectively, at least 50 wt.% And up to 100 wt.% benzene in the feed stream is converted into heavier alkylbenzenes. In addition, the benzene conversion is at least 80%, preferably at least 90%, and most preferably at least 95% of the toluene conversion. Even though a dilute stream of benzene may have a higher concentration of aromatic compounds having eight carbon atoms than a concentration of benzene, benzene undergoes a higher conversion than aromatic compounds having eight carbon atoms. The ratio of benzene to olefin may be between 0.2 and 4.0.

The catalyst remains stable despite impurities in the cheese, but it can be regenerated by deactivation. Suitable regeneration conditions include exposing the catalyst, for example, in situ, to heating with hot air at 500 ° C. for 3 hours. The activity and selectivity of the regenerated catalyst is comparable to that of a fresh catalyst.

The alkylation product stream from the alkylation reactor via line 352 can be transported to the distillation column of the alkylation product 360, which can be a simple low pressure separator, but preferably is a distillation column to separate the gaseous stream from the liquid stream. The distillation column of the alkylation product 360 is connected downstream to the alkylation reactor 320. The gaseous product stream from the overhead via line 362 containing light gases such as hydrogen, methane, ethane, unreacted olefins and light impurities can be directed to the installation of the combustion chamber 370, to generate flow on line 372. Alternatively, the gaseous product from the overhead on line 362 can be burned to start a heater (not shown) and / or provide a source of flue gas to turn Use a gas turbine to generate energy. The overhead line of line 362 is in the upstream connection with the installation of the combustion chamber 370. The liquid bottoms product stream containing heavier hydrocarbons via line 364 from the distillation column of alkylation product 360 can be discharged over the valve and recycled back to product separation section 90 through irrigation LCO 95. Accordingly, the main column 92 is in the downstream and upstream connection to the alkylation reactor 320. The bottoms product stream through line 364 is preferably recycled to the main column 92 at a location between the outlet of the heavy naphtha 96a and the outlet of the light recycle gas oil 95a. Recycling line 364 is in downstream connection with the bottom of the distillation column of the alkylation product 360. Alternatively, recycle line 364 delivers the light recycle gas oil by circulation line 95 or by heavy naphtha circulation through line 96. Recycling line 364 is connected downward to alkylation reactor 320 and is connected upstream with a main column 92. Alternatively, the alkylation product via lines 364 can be transferred to a gasoline distribution system 148, without recycle to the product separation zone 90.

The naphtha stream ranks alkylbenzene with a lower benzene concentration than in line 304, which may be from the side fraction of outlet 366, can be recovered and routed to gasoline distribution system 148. Gasoline distribution system 148 may include a pipe system for the outlet, a dispenser for filling a transportation tank or a gasoline storage tank. The gasoline distribution system 148 is in downstream connection to an alkylation reactor 320. In one embodiment of the invention, the side fraction stream along line 366 has a higher flow rate than the bottom product stream along line 364, which may simply be a resistance stream. EXAMPLES

The usefulness of the present invention will be illustrated by the following examples.

EXAMPLE 1

The extruded UZM-8 catalyst was synthesized by combining UZM-8 powder with Si / Al at a ratio of 12 and pseudoboehmite, sold under the trademark Catapal. Pseudoboehmite was plasticized with nitric acid before mixing with amorphous aluminosilicate. The doughy mass of the catalyst was extruded through a 1.59 mm hole in a cylindrical stamped board and separated into pieces before calcining at 550 ° C. The resulting catalyst consisted of 70 wt.% UZM-8 and 30 wt.% Alumina and had a surface area of 368 m ² / g.

EXAMPLE 2

The catalyst of Example 1 was tested for the alkylation of benzene at 205 ° C., 3447 kPa (500 psi), 0.4 OWHSV (olefin volumetric rate in mass parts per hour) and 4.9 WHSV of liquid feedstock, when conducting the experiment in a stationary layer at 10 ml of catalyst. The gaseous feed consisted of 23 mol% C ₂ H ₄ and 77 mol% CH ₄ . The liquid feed consisted of 2.6 wt.% Pentene, 6.4 wt.% C ₆ H ₆ , 30 wt.% N-heptane, 25 wt.% Toluene, 18 wt.% Aromatic compounds having eight carbon atoms (C ₈ aromatic compounds) and 19 wt.% Aromatic compounds having nine carbon atoms (C ₉ aromatic compounds).

83% of ethylene was converted and 51% of benzene was converted. The toluene conversion was 49%, although the conversion of C ₈ and C ₉ aromatics was less than 20%. The selectivity for ethylbenzene was 5%, for ethyltoluenes it was 25%, for diethylbenzenes it was 6%, for a number of other gasoline compounds it was 37% and for compounds boiling at more than 225 ° C it was 27%.

After 13 hours, the temperature was raised to 236 ° C. At around 26 hours, the WHSV (space velocity) of the liquid feed was increased to 12.7. Under the above conditions, the conversion of ethylene was 83%, with the conversion of benzene 32%, the conversion of toluene 27% and the conversion of C ₈ and C ₉ aromatic compounds less than 15%. The selectivity was 4% for ethylbenzene, 17% for ethyltoluenes, 2.5% for diethylbenzenes, 53.5% for a number of other gasoline compounds, and 23% for compounds boiling at more than 225 ° C.

At 46 hours of operation, the gaseous feed was changed to a mixture representing an FCC dried gas composition containing 34 mol% CH ₄ , 23 mol% C ₂ H ₄ , 14 mol% C ₂ H ₆ , 13 mol% H ₂ , 13 mol.% N ₂ , 2 mol.% CO ₂ , 1 mol.% CO and 1 part per million H ₂ S. Conversions and selectivity did not change. The experiment is presented in figure 2.

EXAMPLE 3

The catalyst of Example 1 was tested for the alkylation of benzene at 205 ° C, 3447 kPa (500 psi), 0.8 OWHSV and 7 WHSV liquid feedstock, in an experiment in a stationary layer at 12 ml of catalyst. The gaseous feed consisted of 34 mol% CH ₄ , 23 mol% C ₂ H ₄ , 14 mol% C ₂ H ₆ , 13 mol% H ₂ , 13 mol% N ₂ , 2 mol% CO ₂ , 1 mol.% CO and 1 part per million H ₂ S, representing raw drained gas. The liquid feed consisted of 39 wt.% Benzene, 49 wt.% N-heptane and 12 wt.% Toluene. Ethylene conversion changed from 77% to 58% at 50 hours of operation, with a constant conversion of benzene 44%. Toluene conversion was 47%. The selectivity was 50-53% with respect to ethylbenzene, 14-16% with respect to ethyltoluenes, 17-19% with respect to diethylbenzenes, 11-14% with respect to a number of other gasoline compounds and 3% with respect to compounds boiling at more than 225 ° C. The experiment is presented in figure 3.

Without further elaboration, it is believed that one skilled in the art will, using the preceding description, be able to use the present invention to its maximum extent. The preceding preferred specific embodiments of the invention are therefore to be construed only as illustrative and not limiting the remainder of the description in any way.

In the foregoing, all temperatures are given in degrees Celsius and all parts and percentages are by weight unless otherwise indicated.

From the above description, a qualified specialist in the art can easily establish the essential features of the present invention and, without going beyond the essence and scope of the invention, can make various changes and modifications of the invention to adapt it to various applications and conditions.

Claims (10)

1. The method of alkylation of benzene with ethylene, including:
providing a diluted ethylene stream containing between 5 and 50 wt.% ethylene;
providing a benzene stream containing at least 3 wt.% toluene and at least 20 wt.% paraffins;
reacting a diluted ethylene stream and a benzene stream with an alkylation catalyst containing UZM-8; and
conversion of at least 20% of the benzene in the feed stream to alkylbenzene. 2. The method according to claim 1, where the conversion of benzene is at least 90% of the conversion of toluene. 3. The method according to claim 1, where the benzene stream has a higher concentration of aromatic compounds with a molecular weight greater than benzene than the concentration of benzene. 4. The method according to claim 1, where the benzene stream contains a higher concentration of aromatic compounds having eight carbon atoms than benzene, however, benzene undergoes a higher conversion than aromatic compounds having eight carbon atoms. 5. The method according to claim 1, where the specified stream of benzene has a more significant concentration of paraffins than benzene. 6. The method according to claim 1, where the specified stage of interaction is performed with a fixed catalyst bed. 7. The method according to claim 1, where the specified diluted stream of ethylene includes at least one impurity selected from the group containing at least 0.05 wt.% Carbon monoxide, at least 1 wt.h. per million hydrogen sulfide, at least 5 wt.% hydrogen and at least 0.1 wt.% carbon dioxide. 8. The method according to claim 1, wherein said benzene stream is supplied as a reformate stream from a reforming reactor. 9. The method of claim 8, wherein said reformate stream is not shared between said reforming reactor and said alkylation reactor. 10. A device for the alkylation of benzene with ethylene, containing:
a fluidized catalyst cracking reactor for reacting the cracking catalyst with a hydrocarbon feed stream to crack the hydrocarbon feed to cracking products having a lower molecular weight and to form coke on the cracking catalyst, thereby promoting the formation of a coke coated cracking catalyst;
a regenerator for burning coke from said cracking catalyst coated with coke by reacting with oxygen;
a separator, in conjunction with said fluidized catalyst cracking reactor, for separating C ₃ hydrocarbons from C ₂ hydrocarbons to provide a diluted ethylene stream;
a reforming reactor for reacting a naphtha stream with a reforming catalyst to obtain a reformate stream; and
an alkylation reactor coupled to said separator and said reforming reactor for alkylating benzene in said reformate stream with ethylene in said diluted ethylene stream above a fixed bed of alkylation catalyst to heavier alkyl benzene hydrocarbons, without distillation column being present in the connection system between said reforming reactor and said alkylation reactor.

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