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Title:
PROCESS FOR THE PRODUCTION OF GASOLINE FROM FUEL GAS AND CATALYTIC REFORMATE
Document Type and Number:
WIPO Patent Application WO/1989/007586
Kind Code:
A1
Abstract:
A process for the production of gasoline from a C4-fuel gas containing ethene and propene and catalytic reformate containing C6 to C8 aromatic hydrocarbons. The C4-fuel gas stream (28) is obtained by catalytic cracking of the heavy distillate fraction (7) obtained by distillation of a crude feed (1) in a distillation apparatus (2). The reformate C6 to C8 aromatic hydrocarbons in stream (15) is obtained by hydrotreating naphtha (5) in zone (9), followed by reforming in zone (13). The C4-fuel gas (28) and C6-C8 aromatic reformate (15) are mixed and contacted in a catalytic reactor (29) with a zeolite catalyst to convert ethene and propene in the fuel gas to C5+ aliphatic and aromatic hydrocarbon gasoline and to convert C6 to C8 aromatics in the reformate into C8 to C11 hydrocarbon gasoline. In a preferred embodiment, the process comprises maintaining a fluidized bed of zeolite catalyst particles in a turbulent reactor at a temperature of 316 to 399°C (600 to 750°F) and pressure of 790 to 825 kPa (100 to 250 psia).

Inventors:
BEECH JAMES HARDING JR (US)
HARANDI MOHSEN NADIMI (US)
KUSHNERICK JOHN DOUGLAS (US)
OWEN HARTLEY (US)
Application Number:
PCT/US1989/000626
Publication Date:
August 24, 1989
Filing Date:
February 14, 1989
Export Citation:
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Assignee:
MOBIL OIL CORP (US)
International Classes:
B01J8/00; B01J8/18; B01J8/26; C07C2/00; C10G29/20; C10G63/02; C10G69/00; (IPC1-7): C07C2/64; C07C2/66
Foreign References:
US3751506A1973-08-07
US3755483A1973-08-28
US3827968A1974-08-06
US4049737A1977-09-20
US4104319A1978-08-01
US4107224A1978-08-15
US4140622A1979-02-20
US4209383A1980-06-24
US4497968A1985-02-05
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Claims:
What Is Claimed Is:
1. A process for the production of gasoline which comprises contacting an olefin feed stream comprising C". olefin hydrocarbons with a hydrocarbon feed stream comprising Cfi to C aromatic hydrocarbons over a zeolite catalyst to convert the olefins to Cς hydrocarbon and to convert the olefins and CgC8 aromatics to Cy to C aromatic hydrocarbon gasoline products .
2. The process of claim 1 wherein the "contacting of the olfin feed stream and the hydrocarbon feed stream over the zeolite catalyst is at a temperature of 260 to 427°C and a pressure of 790 to 2860 kPa.
3. The process of claim 1 or 2 wherein the olefin feed comprises 1 to 90 wt. % ethene.
4. The process of claim 1 , 2 or 3 , wherein the zeolite catalyst comprises HZSM5 zeolite catalyst.
5. The process of any one of the preceding claims wherein C2 to C3 olefins , at WHSV 0.1 to 5 and CgC8 aromatics at WHSV 0.01 to 6.0 are contacted over the zeolite catalyst .
6. The process of any one of the preceding claims wherein the C^ hydrocarbon feed stream comprises 8 to 50 wt.% ethene and 3 to 55 wt.% propene and the aromatics feed stream comprises 10 to 95 wt.% Cg to Cg aromatics .
7. The process of any one of the preceding claims wherein the olefin feed stream comprises C". olefin hydrocarbons including ethene and propene , the hydrocarbon feed stream comprises Cg to Cg aromatic hydrocarbons over a zeolite catalyst at a temperature of 318 to 399°C and pressure of 793 to 1885 kPa to convert the ethene and propene to Cς hydrocarbon and to convert the ethene , propene and Cg to Cg aromatics to Cy. to C, , aromatic hydrocarbon gasoline products . —38— .
8. The process of any one of the preceding claims wherein the hydrocarbon feed stream comprises 2 to 40 wt.% benzene, 10 to 40 toluene and 4 to 50 wt.% Cg aromatics.
9. The process of any one of the preceding claims wherein the olefin feed stream comprises 8 to 35 wt.% ethene and 3 to 40 wt.% propene and the hydrocarbon feed stream comprises 10 to 95 wt.% Cg to Cg aromatics.
10. The process of any one of the preceding claims wherein the hydrocarbon feed stream is obtained by: feeding a first hydrocarbon naphtha distillate stream to a hydrotreating zone operated under hydrotreating conditions and recovering a hydrotreated effluent stream; feeding the hydrotreated effluent stream to a catalytic reforming zone operated under reforming conditions to recover a hydrocarbon stream comprising Cg to Cg aromatic hydrocarbons and Cg to Cy paraffinic hydrocarbons; and ° the olefin feed stream is obtained by: feeding a second hydrocarbon middle distillate stream into a fluidized catalytic cracking zone to recover an overhead olefin feed stream comprising C7 hydrocarbons including ethene and propene; at least a portion of the hydrocarbon feed stream being contacted with the olefin feed stream.
11. The process of claim 10 wherein the entire hydrocarbon feed stream is contacted with the olefin stream and the zeolite catalyst.
12. The process of claim 10 wherein the hydrocarbon stream is first fed to a fractionation column and separated into an aromatic hydrocarbon overhead stream comprising Cg to Cg aromatics and a bottom stream comprising in Cg to C,0 paraffinic hydrocarbons and contacting the Cg to Cg aromatic overhead stream with the olefin feed stream over the zeolite catalyst.
13. A fluidized bed catalytic process for conversion of C. light olefinic gas feedstock to C hydrocarbons and Cg to Cg aromatic feedstock is converted to Cy to ,ι aromatics by: maintaining a fluidiz«ed bed of zeolite catalyst particles in a turbulent reactor bed at a temperature of 260 to 427°C, the catalyst having an apparent particle density of 0.9 to 1.6 g/ml and a size range of 1 to 150 ^. , an average catalyst particle size of 20 to 100 ^ m containing 10 to 25 weight percent of fine particles having a particle size less than 32^ π; contacting the feedstocks and passing the feedstocks upwardly through the fluidized catalyst bed under turbulent flow conditions at reaction conditions sufficient to convert at least 70% 0 of the olefin feedstock and to convert at least 5% of the Cg to Cg aromatic feedstock; maintaining turbulent fluidized bed conditions through the reactor bed between transition velocity and transport velocity at a superficial fluid velocity of 0.3 to 2 meters per second ; and 5 recovering hydrocarbon product containing Cr hydrocarbons and Cy to C, , aromatic hydrocarbons .
14. The process according to claim 13 wherein the fluidized bed density is 100 to 500 kg/m , measured at the bottom of the bed, and wherein the catalyst comprises a siliceous metallosilicate acid zeolite having the structure of ZSM5 zeoli te .
15. The process of claim 13 or 14 wherein the C^ light olefin gas feed stock comprises to 50 wt . % ethene and the catalytic reformate comprises 10 to 95 wt.% Cg to Cg aromatics .
16. The process of any one of claims 1315 wherein the fluidized catalyst bed is maintained in a vertical reactor column having a turbulent reaction zone by passing feedstock gas upwardly through the reaction zone at a velocity greater than dense bed transition velocity to a turbulent regime and less than transport velocity for the average catalyst particle; and withdrawing a portion of coked catalyst from the reaction zone, oxidatively regenerating the withdrawn catalyst and returning regenerated catalyst to the reaction zone at a rate to control catalyst activity whereby the C7 light olefins are converted to Cς olefinic hydrocarbons and the Cg to Cg aromatic hydrocarbons in the catalytic reformate feedstock are converted to Cy to C,•, aromatic hydrocarbons.
17. The process of claim 16 wherein the superficial feedstock vapor velocity is 0.32 m/sec; the reaction temperature is 316 to 399°C; the weight hourly feedstock space velocity (based on olefin equivalent and total reactor catalyst inventory) is 0.1 to 5 and the weight hourly feedstock space velocity (based on Cg to Cg aromatics equivalent and total reactor catalyst inventor) is 001 to 6.0; and the average fluidized bed density measured at the reaction zone bottom is 300 to 500 kg/m .
18. 18 The process of claim 16 wherein the catalyst consists essentially of a medium pore pentasil zeolite having an apparent alpha value of 1 to 80, and average particle size of 20 to 100/z.m, the reactor catalyst inventory includes at least 10 weight percent fine particles having a particle size less than 32 c .
19. 19 The process of claim 18 wherein the catalyst particles comprise 5 to 95 weight percent ZSM5 zeolite having a crystal size of 0.022 _.m.
20. The process of claim 16 wherein the feedstocks comprise C7 light olefin hydrocarbons cracking gas containing 5 to 80 wt.% ethene and propene and catalytic reformate hydrocarbons containing 10 to 95 wt.% Cg to Cg aromatics.
21. The process of claim 16 wherein C7 hydrocarbon product is separated from the C, hydrocarbon product and is recycled back to the reactor.
22. The process of claim 16 wherein the heat of reaction removal and reactor temperature control are enhanced by controlling feed temperature by heat exchange with olefin gas feed and catalytic reformate feed.
23. A fluidized bed process according to claim 13 wherein the fluidized bed density is 100 to 500 kg/m measured at the bottom of the bed.
24. A fluidized bed process according to claims 13 or 16 integrated with a petroleum refinery unsaturated gas separation process and catalytic reforming process the integrated process obtaining light olefin containing feed from a FCC unit and catalytic reformate feed from a reformer unit.
Description:
PROCESS FORTHE PRODUCTION OF GASOLINE FROM FUEL GAS AND CATALYTIC REFORMATE

The present invention relates to a petroleum refining process for the production of gasoline product. The present invention more specifically relates to the production of gasoline by contacting a CT fuel gas containing ethene and propene with a catalytic reformate containing C 6 to C g aro atics over a zeolite catalyst to convert the fuel gas to Cr hydrocarbon gasoline and to convert the C 6 to C„ aromatics to lower alkyl aromatic hydrocarbon gasoline. The process includes the catalytic reforming of naphtha to obtain the catalytic reformate feed and the fluid catalytic cracking of hydrocarbons to obtain the C.- fuel gas feed to the zeolite catalyst conversion zone.

The fluid catalytic cracking of hydrocarbons in modern refinery operations produces large amounts of C. - fuel gas of little or no gasoline product value and the catalytic reforming of hydrocarbons produces large amounts of Cg to C g aromatic hydrocarbons which though having value as gasoline blending stock are produced in excessive amounts. The present invention particularly relates to a catalytic technique for upgrading light olefin gas to heavier hydrocarbons and to alkylating C fi to C g aromatics to heavier lower alkyl aromatic hydrocarbons. In particular, it provides a continuous process for processing olefinic light gas feedstock, containing ethene and propene, or other lower alkenes, to produce C hydrocarbons, such as olefinic liquid fuels, isobutane, aromatics, e.g. benzene, and other useful products and at the same time alkylating Cg to Cg aromatics to produce C, to C. lower alkyl substituted aromatic hydrocarbons for use as gasoline blending stock. Ethene (ethylene, C-H.)-containing gases, such as petroleum cracking offgas, and catalytic reformate containing benzene, toluene, xylene and ethyl benzene are useful feedstocks for the process.

Developments in zeolite catalysis and hydrocarbon conversion processes have created interest in utilizing olefinic feedstocks for producing C ς gasoline, diesel fuel, etc. In addition to basic chemical reactions promoted by ZSM-5 type zeolite catalysts, a number of discoveries have contributed to the development of new industrial processes. These are safe, environmentally acceptable processes for utilizing feedstocks that contain lower olefins, especially £ -0. alkenes and feedstocks containing aromatic compounds, especially C, to C g aromatics. Chen USP 3,729,409 discloses improving the yield-octane number of a reformate by contacting the reformate in the presence of hydrogen over a zeolite catalyst. Garwood et al USP 4,150,062 discloses a process for the conversion of j to C. olefins to produce gasoline which comprises contacting the olefins with water over a zeolite catalyst. The Haag et al USP 4,016,218 and Burress - USP 3,751,506 disclose processes for the alkylation of benzene with olefins over a ZSM-5 type catalyst. The Reroute et al USP 4,209,383 discloses the catalytic alkylation of benzene in reformate with C,-C. olefins to produce gasoline. Conversion of C 2 -C 4 alkenes and alkanes to produce aromatics-rich liquid hydrocarbon products were found by Cattanach USP 3,760,024 and Yan et al USP 3,845,150 to be effective processes using the ZSM-5 type zeolite catalysts. In U.S. Patents 3,960,978 and 4,021,502, Plank, Rosinski and Givens disclose conversion of C 2 -C 5 °l e ^ ns > alone or in admixture with paraffinic components, into higher hydrocarbons over crystalline zeolites having controlled acidity. Garwood et al have also contributed to the understanding of catalytic olefin upgrading techniques and improved processes as in U.S. Patents 4,211,640 and 4,227,992. It has now been found that contacting a catalytic reformate feed comprising C g to C g aromatic hydrocarbons with a light olefin gas feed, comprising ethene or ethene and propene, over a zeolite catalyst that the C g to C g aromatics in the catalytic

reformate can be converted to lower alkyl aromatic hydrocarbons while at the same time converting the ethene and propene to Cr hydrocarbons both of which products are suitable for use as gasoline blending stocks.

The present invention is directed to a process for the production of gasoline which comprises the steps of fractionating a crude oil feed stream into a light first distillate and a heavy second distillate; passing the light first distillate through a catalytic hydrotreating zone and then through a catalytic reforming zone to obtain a catalytic reformate stream containing Cg to C g aromatic hydrocarbons; passing the heavy second distillate into a fluidized catalytic cracking zone which includes a fractionating column and producing an overhead C.- olefinic hydrocarbon fuel gas vapor stream; and contacting the catalytic reformate stream and the c7 fuel gas stream in a zeolite catalyst reaction zone under process conditions to produce Ce hydrocarbons from the C. fuel gas and lower alkyl aromatic hydrocarbons from the reformate stream. The C- hydrocarbons and the alkyl aromatic hydrocarbons are both suitable gasoline blending stocks. The present invention is more specifically directed to an improved process for the conversion of ethene-containing feedstocks and Cg to C g aromatics containing feedstocks to heavier hydrocarbon products of higher octane value wherein the feedstocks are contacted at elevated temperature and pressure with a fixed, moving or fluidized bed of zeolite catalyst under conversion conditions.

In accordance with the present invention it has been found that ethene-rich olefinic light gas can be upgraded to liquid hydrocarbons rich in olefinic gasoline, isobutane and aromatics and that catalytic reformate containing Cg to C g aromatics can be upgraded to lower alkyl aromatic hydrocarbons of higher octane value by catalytic conversion in a turbulent fluidized bed of solid acid zeolite catalyst under reaction conditions in a single pass or with

recycle of gas product. This technique is particularly useful for upgrading FCC light gas, which usually contains significant amounts of ethene, propene, C -C4 paraffins and hydrogen produced in cracking heavy petroleum oils or the like and for upgrading catalytic reformate containing Cg to C g aromatics and C g to Cg paraffins. Ey upgrading the by-product light gas and the catalytic reformate, the gasoline yield of FCC units and catalytic reforming units can be significantly increased. Accordingly, it is a primary object of the present invention to provide a novel technique for upgrading ethene-rich light gas and Cg to C g rich catalytic reformate. improved process has been found for continuous conversion of ethene-containing and Cg to C g aromatic hydrocarbon containing feedstocks to heavier hydrocarbon products of higher octane value wherein the feedstock is contacted at elevated temperature with a fluidized bed of zeolite catalyst under conversion conditions. The improvement comprises maintaining the • fluidized catalyst bed in a vertical reactor column having a turbulent reaction zone by passing feedstock gas upwardly through the reaction zone at a velocity greater than dense bed transition velocity in a turbulent regime and less than transport- velocity for the average catalyst particle; and withdrawing a portion of coked catalyst from the reaction zone, oxidatively regenerating the withdrawn catalyst and returning regenerated catalyst to the reaction zone at a rate to control catalyst activity.

In accordance with the present invention in the same reaction zone an ethene-rich olefinic light gas can be upgraded to liquid hydrocarbons rich in olefinic gasoline and a catalytic reformate rich in Cg to C g aromatics can be upgraded to lower alkyl aromatic hydrocarbons of higher octane value by catalytic conversion in the turbulent regime of a fluidized bed of solid acid zeolite catalyst in a single pass or with recycle of light gas product.

Figure 1 of the drawings is a flow diagram of the petroleum refining process of the present invention for the production of gasoline.

Figure 2 is a graphic illustration showing the effect of ' 5 reaction zone temperature on the Cr hydrocarbon yield based on olefin feed.

Figure 3 is a graphic illustration showing the effect of reformate and olefin feed hourly space velocity and reaction zone temperature on the C r hydrocarbon product octane value. 10 Figure 4 is graphic illustration showing the effect of reaction zene temperature and pressure on the C_ hydrocarbon yield from reformate and olefin feed.

Figure 5 is a graphic illustration showing the effect of reaction zone temperature on the C_ hydrocarbon yield from " 15 reformate and olefin feed.

Figure 6 is a graphic illustration showing the effect of reaction zone temperature on the C r hydrocarbon yield from reformate and olefin feed.

Figure 7 illustrates an embodiment of this invention in 20 which the reaction is carried out in the turbulent zone of a fluidized bed and the regeneration and recycle of the catalyst.

The present invention utilizes conventional petroleum refining steps including fractionation, hydrotreating, catalytic reforming and fluidized catalytic cracking and a novel zeolite 25 catalyst process to upgrade the fuel gas and reformate process streams. A gasoline boiling range product is produced from the fuel gas stream from the fluidized catalytic cracking process step and the reformate stream from the catalytic reforming step.

In accordance with the present invention crude oil feed is 30 subjected to atmospheric distillation to separate several hydrocarbon streams including a light gas, a gasoline boiling range naphtha, a middle distillate, a heavy distillate and a bottoms or reduced crude stream.

The naphtha stream is hydrotreated to remove sulfur and 35 nitrogen compounds and then fed to a catalytic reforming zone wherein the octane value of this stream is increased, the

concentration of aromatic hydrocarbons is increased and hydrogen is produced as a by-product.

The middle distillate stream is hydrotreated to produce products such as kerosene and jet fuel. The heavy distillate is fed to a fluidized catalytic cracking (FCC) zone in which there is produced a light gasoline boiling range distillate, a fuel gas containing C-^ to C^ olefins and paraffins and a heavy distillate.

The reduced crude may be fed into a subatmospheric pressure or vacuum fractionation column. The reduced crude may also be subjected to processing steps such as propane deasphalting, hydrocracking, etc.

, The catalytic reformate containing Cg to C g aromatic hydrocarbons and the fuel gas stream containing C, to C . olefins and paraffins is then fed to the zeolite catalyst reaction zone.

The zeolite catalyst reactisn zone is operated under conditions such that ethene or ethene and propene in the fuel gas feed stream are converted to C^ olefinic gasoline product. The ethene or ethene and propene in the fuel gas feed stream also react with the Cg to C g aromatic hydrocarbons in the reformate feed stream to produce C- to C,, aromatic hydrocarbons such as toluene, xylene, ethyl benzene, methyl ethyl benzene, diethyl benzene, propyl benzene and methyl propyl benzene.

The effluent stream from the zeolite reaction zone is passed into a separator in which a Cg hydrocarbon stream is removed overhead and fed to an absorber in which the C hydrocarbons are absorbed and removed. The remaining C_ hydrocarbons are taken overhead and can be recycled to the zeolite catalyst reaction zone. The bottoms from the separator contain C-

.f. ' to C,, aromatic hydrocarbons and C- hydrocarbons and is fed to a debutanizer from whic an overhead C ~ gas stream is removed. A portion of the CT stream can be recycled to the zeolite catalyst reaction zone. The debutanized gasoline product is removed as a bottoms product and is fed to the gasoline product pool.

Description of Figure 1 of the Drawings A .crude oil feed is fed through line 1 to an atmospheric distillation column 2 and is separated into fractions having different boiling point ranges.* The C, to C . light hydrocarbons and any gases dissolved in the feed are removed overhead through line 3 and passed to a gas recovery zone. The light normally liquid hydrocarbons are removed as a naphtha stream through line 5, a middle distillate stream through line 6, and a heavy distillate stream through line 7. The remaining reduced crude is removed through line 8 for further processing.

The naphtha fraction and the middle distillate fraction are passed to hydrotreating zones 9 and 10, respectively. The middle distillate ' hydrotreated stream is removed in line 11 and is in the kerosene boiling range hydrocarbons. The naphtha hydrotreated stream is passed through line 12 to a reforming zone 13 wherein it is catalytically reformed to produce a reformate containing C fi to C g aromatic hydrocarbons and C, paraffinic hydrocarbons and hydrogen. The hydrogen is removed, as a by-product, overhead in line 14. The catalytic reformate is fed through line 15 to a fractionating column 16 in which a portion of the Cg paraffinic hydrocarbons can be removed through line 23 and fed to the gasoline product pool. The overhead line 17 contains C β to C g aromatic hydrocarbons, any remaining unseparated Cg paraffinic hydrocarbons and the Cg paraffinic hydrocarbons of the catalytic reformate. Alternatively, particularly with light reformate streams, the fractionating step can be omitted and the entire reformate effluent stream can be fed directly to the zeolite reaction zone 29.

The heavy distillate removed through line 7 is fed to a fluidized bed catalytic cracking zone 18 which includes a fractionating column 19. The overhead vapor stream from the

fractionating column 19 is removed in line 20 and cooled in condenser 21 and then fed to receiver 22. The condensate collected in receiver 22 is fed through line 24 to a primary absorber 25. The uncondensed gases in receiver 22 containing C 4 olefins are removed overhead through line 26 and fed to primary absorber 25. A bottom liquid stream is removed from the fluid catalytic cracker fractionating column 19 through line 27 and is fed to the top of the primary absorber 25. overhead gas stream including CT olefins is removed in line 28 and is fed to zeolite catalyst reaction zone 29 with the catalytic reformate including Cg to C g - aromatics fed through the line 17 and are contacted together and over a zeolite catalyst in reaction zone 29.

The bottom line 30 from the primary absorber 25 contains C,. gasoline product and is fed to gasoline product pool. The CT feed in line 28 is catalytically converted in zeolite catalyst reaction zone 29 to C- hydrocarbon gasoline product. The cT feed is contacted with the Cg to C g aromatics in the catalytic reformate in line 17 and is catalytically converted at the same time in the zeolite catalyst reaction zone 29 to C- to C,-, aromatic hydrocarbon gasoline product. The zeolite reaction zone 29 product is removed from the reaction zone via line 31 and passed to separator 32. The overhead vapor products are fed via line 38 to absorber 33 and contacted with a suitable absorber oil fed through line 34 to remove C_ hydrocarbons and absorber oil in line 35. The overhead line 36 contains Cl hydrocarbons which can be recycled via line 39 to the zeolite catalyst reaction zone 29. The absorber oil and C, hydrocarbons in line 35 are treated to separate the C- hydrocarbons and recycle the absorber oil. The bottoms from separator 32 is removed via line 37 and comprises the C hydrocarbon and C- to C,, aromatic hydrocarbon gasoline products and is fed to debutanizer 40 from whic an overhead CT gas stream is removed via line 41. A portion of the C. stream can be recycled via line 39 to the zeolite catalyst reaction zone 29.

The debutanized gasoline product is removed via line 42 and is fed to the gasoline product pool.

The fractionation column 16 when used functions to control the amount of Cg-C 8 paraffinic hydrocarbons and the amount. C ή " C 8 aromatic hydrocarbons that are fed to reaction zone 29. The bottom line 23 from separator 16 contains C g gasoline product.

The zeolite catalyst reaction zone 29 is maintained at conditions of temperature and pressure such that the C olefin " stream is converted to C_ hydrocarbons, including aliphatic and aromatic hydrocarbons, and the c7 olefin stream and the catalytic reformate stream containing Cg to C g aroϋiatics is converted to C- to C-,, alkyl aromatic hydrocarbons such as toluene, xylene, ethyl benzene, methyl ethyl benzene, diethyl benzene and propyl benzene.

For purposes of clarity in the above description of the invention various subsystems and apparatus normally associated with the operation of the process have not been shown. The omitted items include pump, temperature, pressure and flow control systems, reactor and fractionator internals, crude column stripers, separators, absorbers, reboilers, overhead condensing systems, etc. which may be of conventional design.

Description of the Zeolite Catalyst Recent developments in zeolite technology have provided a group of medium pore siliceous materials having similar pore geometry. Most prominent among these intermediate pore size zeolites is ZSM-5, which is usually synthesized with Eronsted acid active sites by incorporating a tetrahedrally coordinated metal, such as Al, Ga, B or Fe, within the zeolitic framework. These medium pore zeolites are favored for acid catalysis; however, the advantages of ZSM-5 structures may be utilized by employing highly

siliceous material or crystalline metallosilicate having one or more tetrahedral species having varying degrees of acidity. ZSM-5 crystalline structure is readily recognized by its X-ray diffraction pattern, which is described in USP 3 ,702,866 CArgauer, et al. ) , The zeolite catalysts preferred for use herein include the medium pore (i.e. , 5-7 x 10 mm) shape -selective crystalline aluminosilicate zeolites having a silica -to -alumina ratio of at least 12, a constraint index of 1 to 12 and acid cracking activity of 1-200. In an operating reactor the coked catalyst may have an apparent activity (alpha value " ) of 1 to 80 under the process conditions to achieve the required degree of reaction severity. Representative of the ZSM-5 type zeolites are ZSM-5 , ZSM-11 , ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38. ZSM-5 is disclosed in USP 3,702,886 and USP Re . 29,948. The ZSM-5 and ZSM-12 catalyst are preferred. Other suitable zeolites are disclosed in U.S. Patents

3 ,709 ,979; 3 ,832,449; 4,076 ,979; 3,832,449; 4,076 ,842; 4 ,016 ,245 and 4 ,046 ,839; 4,414,423; 4 ,417 ,086; 4 ,517 ,396 and 4 ,542,251. While suitable zeolites having a coordinated metal oxide to silica molar ratio of 20:1 to 200:1 or higher may be used , it is advantageous to employ a standard ZΣ -5 having a silica alumina molar ratio of 25 :1 to 70:1, suitably modified. A typical zeolite catalyst component having Eronsted acid sites may consist essentially of aluminosilicate ZSM-5 zeolite with 5 to 95 wt. % silica and/or alumina binder. Certain of the ZSM-5 type medium pore shape selective catalysts are sometimes known as pentasils. In addition to the preferred aluminosilicates , the borosilicate, ferrosilicate and "silicalite" materials may be employed. It is advantageous to employ a standard ZSM-5 having a silica: alumina molar ratio of 25 :1 to 70:1 with an apparent alpha value of 1-80 * to convert 60 to 100 percent, preferably at least 70%, of the olefins in the feedstock and to convert 1 to 50% preferably at least 5% of the Cg-C 8 aromatics in the feedstock.

ZSM-5 type pentasil zeolites are particularly useful in the process because of their regenerability, long life and stability under the extreme conditions of operation. Usually the zeolite crystals have a crystal size from 0.01 to over 2 microns or more, with 0.02-1 micron being preferred. The zeolite catalyst crystals are normally bound with a suitable inorganic oxide, such as silica, alumina, etc. to provide a zeolite concentration of 5 to 95 wt.%. A preferred catalyst, comprises 25% to 65% H-ZSM-5 catalyst contained within a silica-alumina matrix binder and having a fresh alpha value of less than 80. The process of the present invention can be carried out in a fixed bed, moving bed and fluidized bed reactor.

When employing a ZSM-5 type zeolite catalyst in fine powder form such a catalyst should comprise the zeolite suitably bound or impregnated on a suitable support with a solid density (weight of a representative individual particle divided by its apparent "outside" volume) in the range form 0.6-2 g/ml, preferably 0.9-1.6 g/ml. The catalyst particles can be in a wide range of particle sizes up to 250^m, with an average particle size between 20 and 100 m , preferably in the range of 10-150 y m and with the average particle size between 40 and 80 icτ ~ . When these solid particles are placed in a reactor bed where the superficial fluid velocity is 0.3-2, fluidized bed operation is obtained. The velocity specified here is for an operation at a total reactor pressure of 100 to 300 kPa (0 to 30 psig). Those skilled in the art will appreciate that at higher pressures, a lower gas velocity may be employed to ensure fluidized bed operation.

Particle size distribution can be a significant factor in achieving overall homogeneity in turbulent regime fluidization. It is desired to operate the present process with particles that will mix well throughout the bed. Large particles having a particle size greater than 250 _<m should be avoided and it is advantageous to employ a particle size range consisting essentially of 1 to 150 ι<_ m. Average particle size is usually 20 to 100 _ , preferably 40

to 80 ^ m. Particle distribution may be enhanced by having a mixture of larger and smaller particles within the operative range, and it is particularly desirable to have a significant amount of fines. Close control of distribution can be maintained to keep 10 to 25 wt % of the total catalyst in the reaction zone in the size range less than 32 ^m. This class of fluidizable particles is classified as Geldart Group A. Accordingly, the turbulent fluidization regime is controlled to assure operation between the transition velocity and transport velocity. Fluidization conditions are substantially different from those found in non-turbulent dense beds or transport beds

The light paraffin production and alkyl aromatic production is promoted by the zeolite catalysts having a high concentration of Bronsted acid reaction sites. Accordingly, an important criterion is selecting and maintaining the catalyst to provide either fresh catalyst having acid activity or by controlling catalyst deactivation and regeneration rates to provide an apparent average alpha value of 1 to 50.

Hydrocarbon Feed Streams To Zeolite Catalyst Reaction Zone Light Olefin Gas

The preferred light olefin gas feedstock contains C- to C 4 alkenes (mono-olefins) including at least 2 moles % ethene, wherein the total C 2 -C, alkenes are in the range of 10 to 40 wt %. Non-deleterious components, such as methane, Z- to C. paraffins and inert gases, may be present. Some of the paraffins will be converted to C. hydrocarbons depending on the reaction conditions and catalyst employed. A particularly useful feedstock is a light gas by-product of FCC gas oil cracking units containing typically 10-40 mol % C~-C- olefins and 5-35 ol % H 2 with varying amounts of C, to C, paraffins and inert gas , such as N~ . The feedstock can contain primarily ethene or ethene and propene.

The light olefin feed gas is described in more detail in the Table 1 below.

Table 1

The catalytic reformate feedstock contains Cg to C g aromatic hydrocarbons and Cr to Cg paraffinic hydrocarbons. The Cg to C g aromatic hydrocarbons include benzene, toluene, xylene and ethyl benzene. The xylene and ethyl benzene are herein considered together as C g aromatic hydrocarbon. Though the catalytic reformate is a preferred feedstock, hydrocarbon process streams containing essentially the same hydrocarbon components can also be used.

The catalytic reformate feedstock is described in more detail below Table 2.

Table 2

Specific Gravity Boiling Range , °F Mole % Benzene Toluene

C g Aromatic CD .

Weigit .

Benzene

Toluene

C g Aromatic Cg-C 8 Aromatics

(1) Xylene and ethyl benzene component.

Hydrocarbon Products

The contacting of the light olefin gas feed with the catalytic reformate feed over the zeolite catalyst in accordance with the present invention produces the ollowing products.

The ethene and propene components of the light olefin gas feed react to produce primarily C ζ to C g olefinic, Cr to C Q paraffinic and Cg to C g aromatic gasoline products which have a higher product value than the ethene and propene in the feed. The principle product is Cr to C g olefinic gasoline product, i.e. the C ς olefinic hydrocarbons.

The ethene and propene components of the light olefin gas feed in addition react with the Cg to C g aromatics in the catalytic reformate feed to produce primarily Cy to C,, aromatics which may themselves rearrange and transalkylate over the zeolite catalyst.

The C to C,-, aromatic hydrocarbon product obtained includes C, to C. lower alkyl substituted aromatic hydrocarbons such as methyl, ethyl, propyl and butyl benzene compounds. The C to C,-, aromatic hydrocarbon product contains one or more of the foregoing lower alkyl substituents, providing however that the total numbers of carbon atoms in the substituents does not exceed 5. Typical, C_ to C,, aromatic hydrocarbons include toluene, ethyl benzene, methyl ethyl benzene, propyl benzene, methyl propyl benzene, butyl benzene, methyl butyl benzene and diethyl benzene. The incorporation of the Cr hydrocarbon component, e.g. the Cr olefinic hydrocarbons, into the Cy-C,-, aromatic hydrocarbon component enriches the overal octane quality of the gasoline product obtained.

The zeolite catalyst process conditions of temperature and pressure are closely controlled to minimize cracking of C 4 to Cy paraffin hydrocarbons in the feed and is an important feature of the present invention.

Unreacted ethene and propene, and butene formed in the reaction can be recycled to the zeolite catalyst reactor. The ethene and propene in the light olefin feed are converted in an amount of 20 to 100, preferably 60 to 100 and more preferably 80 to 100 wt.% of the feed.

The Cg to Co aromatics in the catalytic reformate feed, including benzene, toluene and C g aromatics, are converted in an amount of 5 to 60 and preferably 8 to 40 wt.% of the feed.

Process Conditions

The process of the present invention using a ZSM-5 type zeolite catalyst is carried out at temperatures of 204 to 427°C (400 to 800°F), for example 260 to 427°C (500 to 800°F), preferably 260 to 399°C.(500 to 750°F) and more preferably 316 to 399°C (600 to

750°F). -J The pressure at which the reaction is carried out is an

- important parameter of the invention. The process can be carried ' out at pressures of 445 to 3550 kPa (50 to 500 psig), preferably 790 0 to 2860 kPa (100 to 400 psig) and more preferably 790 to 825 kPa (100-250 psig).

The weight hourly space velocity (WHSV) of the light olefin feed and the catalytic reformate feed are also important parameters of the process. 5 The principal reactants in the process are the ethene or ethene and propene constituents of the light olefin gas and the Cg to C g aromatic constituent of the catalytic reformate and the WHSV " are given in terms of these components. -

The ethene and propene WHSV can be 0.1 to 5.0, preferably 0 0.1 to 2 and more preferably 0.5 to 1.5.

The Cg to C g aromatics WHSV can be 0.01 to 6.0, preferably 0.1 to 4.0 and more preferably 0.1 to 2.0.

The C ς hydrocarbon production and alkyl aromatic production is promoted by those zeolite catalysts having a high 5 concentration of Bronsted acid reaction sites. Accordingly, an important criterion is selecting and maintaining catalyst inventory to provide either fresh catalyst having acid activity or by controlling catalyst deactivation and regeneration rates to provide an apparent average alpha value of 1 to 80.

The Reactor

The process can be carried out in a conventional fixed bed, moving bed or fluidized bed reactor.

In a preferred embodiment, the use of the turbulent regime fluidized bed catalyst process permits the conversion system to be operated at low pressure drop. An important advantage of the process is the close temperature control that is made possible by turbulent regime operation, wherein the uniformity of conversion temperature can be maintained within close tolerances, often less than 25°C. Except for a small zone adjacent the bottom gas inlet, the midpoint measurement is representative of the entire bed, due to the thorough mixing achieved.

_ In a typical process, the ethene-ri -ch C Δ- olefinic feedstock and Cg to C g rich feedstock are converted in a catalytic reactor under 600 to 750°F (260 to 399°C) temperature and moderate pressure 100 to 250 psig (i .e . 790 to 1825 kPa) to produce * a predominantly liquid product consisting essentially of C- aliphatic hydrocarbons rich in gasoline-range olefins and Cy to C- j , alkyl aromatic hydrocarbons . Referring to the Figure 7 of the drawings , a pressurized feed gas rich in C 2 -C, olefins is fed through line 112 and heated in heat exchanger 115 and then fed to line 113. A pressurized reformate feed rich in Cg-C 8 aromatic hydrocarbons is fed through line 110 and heated in heat exchanger 111 and then fed to line 113 wherein it is contacted and mixed with heated olefin feed gas. A major portion of the olefin feed gas is mixed in line 113 with the reformate feed and fed through line 113 to the bottom inlet of reactor vessel 120 for distribution through grid plate 122 into fluidization zone 124. Here the mixed olefin and Cg to C g aromatic hydrocarbon feed contact the turbulent bed of finely divided catalyst particles . The remainder of the heated olefin feed gas is fed through line 114 to catalyst return riser conduit 150 in which it functions as a lift gas for the regenerated catalyst.

The reaction heat can be partially or completely removed by using cold or only partially preheated olefin feed gas and catalytic reformate feed. Baffles may be added to the reactor vessel to control radial and axial mixing. Heat released from the reaction can be controlled by adjusting feed temperature in a known manner. Catalyst outlet means 128 is provided for withdrawing catalyst from bed 124 and passed for catalyst regeneration in vessel 130 via control valve 129. The outlet means 128 may include a steam stripping section, not shown, in which useful hydrocarbons are ' removed from the catalyst prior to regeneration of the catalyst. The partially deactivated catalyst is oxidatively regenerated by controlled contact with air or other regeneration gas-,,at elevated temperature in a fluidized regeneration zone 130 to remove carbonaceous deposits and restore catalyst acitivity. The catalyst particles are entrained in a lift gas provided via line 147 and transported via riser tube 132 to a top portion of vessel 130. Air is distributed at the bottom of the bed via line 144 to effect fluidization, with oxidation byproducts being carried out of the regeneration zone through cyclone separator 134, which returns any entrained solids to the bed. Flue gas is withdrawn via top conduit 136 for disposal; however, a portion of the flue gas may be recirculated via heat exchanger 138, separator 140, and compressor 142 for return to the vessel through line 147 with fresh oxidation gas fed via line 144 and as fluidizing gas for the regenerator 130 and as lift gas for the catalyst in riser 132.

Regenerated catalyst is passed to the main reactor 120 throug conduit 146 provided with flow control valve 148. The regenerated catalyst may be lifted to the catalyst bed through return riser conduit 150 with pressurized olefin feed gas fed through line 114 to catalyst return riser conduit 150. Since the amount of regenerated catalyst passed to the reactor is relatively small, the temperature of the regenerated catalyst does not upset the temperature constraints of the reactor operations in significant

amount. A series of sequentially connected cyclone separators 152, 154 are provided with diplegs 152A, 154A to return any entrained catalyst fines to the lower bed. These separators are positioned in an upper portion of the reactor vessel comprising dispersed ' catalyst phase. Filters, such as sintered metal plate filters, can be used alone or in conjunction with cyclones.

The hydrocarbon product effluent separated from catalyst particles in the cyclone separating system is then withdrawn from the reactor vessel 120 through top gas outlet means 156. The recovered hydrocarbon product comprising C ς olefins, aromatics, paraffins, alkyl aromatics and naphthenes is thereafter processed as required to provide the desired gasoline product.

Under optimized process conditions the turbulent bed has a .. superficial vapor velocity of 0.3 to 2 meters per second (m/sec). At higher velocities entrain ent of fine particles may become excessive and beyond.3 m/sec the entire bed may be transported out of the reaction zone. At lower velocities, the formation of large bubbles or gas voids can be detrimental to conversion. Even fine particles cannot be maintained effectively in a turbulent bed below 0.1 m/sec.

A convenient measure of turbulent fluidization is the bed density. A typical turbulent bed has an operating density of 100 to

3 ~

500 kg/m , preferably 300 to 500 kg/m , measured at the bottom of the reaction zone, becoming less dense toward the top of the reaction zone, due to pressure drop and particle size differentiation. This density is generally between the catalyst concentration employed in dense beds and the dispersed transport systems. Pressure differential between two vertically spaced points in the reactor column can be measured to obtain the average bed density at such portion of the reaction zone. For instance, in a fluidized bed system employing ZSM-5 particles having an apparent

- - packed density of 750 kg/m and real density of 2430 kg/m , an

- average fluidized bed density of 300 to 500 kg/m" is satisfactory.

By virtue of the turbulence experienced in the turbulent regime, gas-solid contact in the catalytic reactor is improved, providing substantially complete conversion, enhanced selectivity and temperature uniformity. One main advantage of this technique is the inherent control of bubble size and characteristic bubble lifetime. Bubbles of the gaseous reaction mixture are small, random and short-lived, thus resulting in good contact between the gaseous reactants and the solid catalyst particles.

A significant difference between the process of this invention and conversion processes of the prior art is that operation in the turbulent fluidization regime is optimized to produce high octane C_ aliphatic hydrocarbon liquid in good yield from the CT fuel gas feed and to produce high octane Cy to C,, aromatic hydrocarbon product in good yield from the catalytic reformate feed. The zeolite catalyst process conditions, including temperature and pressure, in the turbulent regime of the fluidized bed are closely controlled to minimize cracking of C, to Cg paraffin hydrocarbons in the feed and is an important feature of the present invention. The weight hourly space velocity and uniform contact provides a close control of contact time between vapor or vapor and liquid and solid phases, typically 3 to 25 seconds. Another advantage of operating in such a mode is the control of bubble size and life span, thus avoiding large scale gas by-passing in the reactor. As the superficial gas velocity is increased in the dense bed, eventually slugging conditions occur and with a further increase in the superficial gas velocity the slug flow breaks down into a turbulent regime. The transition velocity at which this turbulent regime occurs appears to decrease with particle size. The turbulent regime extends from the transition velocity to the so-called transport velocity, as described by Avidan et al in USP 4,547,616. As the transport velocity is approached, there is a sharp increase in the rate of particle carryover, and in the absence of solid recycle, the bed could empty quickly.

Several useful parameters contribute to fluidization in the

_ turbulent regime in accordance with the process of present invention. When employing a ZSM-5 type zeolite catalyst in fine powder form such a catalyst should comprise the zeolite suitably bound or impregnated on a suitable support with a solid density (weight of a representative individual particle divided by its apparent "outside" volume) in the range from 0.6-2 g/ml, preferably 0.9-1.6 g/ml. The catalyst particles can be in a wide range of particle sizes up to 250,__m, with an average particle size between 20 and 100 Mπ, preferably in the range of 10-150^ and with the average particle size between 40 and 80 ___m. When these solid particles are placed in a fluidized bed where the superficial fluid velocity is 0.3-2 m/sec, operation in the turbulent regime is obtained. The velocity specified here is for an operation at a total reactor pressure of 100 to 300 kPa (0 to 30 psig). Those skilled in the art will appreciate that at higher pressures, a lower gas velocity may be employed to ensure operation in the turbulent fluidization regime.

The reactor can assume any technically feasible configuration, but several important criteria should be considered. The bed of catalyst in the reactor can be at least 5-20 meters in height. Fine particles may be included in the bed, especially due to attrition, and the fines may be entrained in the product gas stream. A typical turbulent bed may have a catalyst carryover rate up to 1.5 times the reaction zone inventory per hour. If the fraction of fines becomes large, a portion of the carryover can be removed from the system and replaced by larger particles. It is feasible to have a fine particle separator, such as a cyclone and/or filter means, disposed within or outside the reactor shell to recover catalyst carryover and return this fraction continuously to the bottom of the reaction zone for recirculation at a rate of one catalyst inventory per hour. Optionally, fine particles carried from the reactor vessel entrained with effluent gas can be recovered by a high operating temperature sintered metal filter.

This ' process can be used with process streams which contains sufficient amounts of light olefins and Cg to C g aromatics. For example, it can be used to process FCC by-product fuel gas , which typically contains 10 to 40 wt.% total ethene and propene and catalytic reformate which contains 2 to 40 wt.% Cg to

C g aromatics .

Reactor Operation

A typical reactor unit employs a temperature-controlled catalyst zone with indirect heat exchange and/or adjustable gas quench, whereby the reaction temperature can be carefully controlled within an operating range of 204 to 427°C (500 to 800 °F) , preferably at average reactor temperature of 316 to 399°C (600 to 750°F) . The reaction temperature can be in part controlled by exchanging hot reactor effluent with feedstock and/or recycle streams. Optional heat exchangers may recover heat from the effluent stream prior to fractionation. Part or all of the reaction heat can be removed from the reactor by using cold feed, whereby reactor temperature can be controlled by adjusting feed temperature. The reactor is operated at moderate pressure of 50 to 500 psig (445 to 3550 kPa) , preferably 100 to 250 psig (790 to 1825 kPa) .

The weight hourly space velocity (WHSV) , based on olefins in the fresh feedstock is 0.1=5 WHSV and the weight hourly space velocity (WHSV) based on Cg-C 8 aromatics 0.01 to 6.0 WHSV. Typical product fractionation systems that can be used are described in USP 4 ,456,779 and USP 4 ,504 ,693 (Owen et al ) .

The present invention is exemplified by the following Example. The process was carried out in a turbulent fluidized bed reactor using a HZSM-5 catalyst comprising a weight ratio of catalyst to silica-alumina binder of 25/75.

EXAMPLE 1

The process is carried out in a fluidized bed reactor using an HZSM-5 zeolite catalyst having an alpha value of 40. The reactor bed temperature is maintained at 316 °C (600 °F) and at a pressure of 790 kPa (100 psig) . The olefin feed is fed at a WHSV of 0.5 , based on ethene and propene. The reformate is fed at a WHSV of 0.7 , based on Cg to Cg aromatics (0.1 based on benzene) in the reformate feed . The reaction is carried out without recycle of light olefins . The components of the olefin gas feed stream and of the reformate feed stream and the components of the total hydrocarbon feed as well as the components of the hydrocarbon product are given below.

—24-

Individual Feed Stream Com ositions

The above Example shows substantial conversion of C- and C_ olefins to Cr olefins and substantial conversion of Cg-C 8 aromatic hydrocarbons to C Q aromatic hydrocarbons.

—25—

The Examples 2, 3 and 4 were carried out using a fixed bed tubular reactor and an HZS_-5 zeolite catalyst.

EXAMPLE 2

The process was carried out in a fixed bed reactor using an HZSM-5 zeolite catalyst having an alpha value of 40. The catalyst had a silica to alumina ratio of 70/1. The catalyst was incorporated in a silica-alumina binder at a ratio of HZSM-5/Binder of 65/35. The reaction was carried out at temperatures ©f 149 to 371°C (300 to 700 °F) and at a pressure of 1825 kPa (250 psig) . The components of the olefin gas feed stream and the reformate feed stream are given below. Individual Feed Stream Compositions Olefin Gas Feed (Average)

N 2 H 2

C 2

S

Reformate Feed

Benzene Toluene

C g Aromatic

Hexanes

Heptanes

Octanes c;

C 5

Reformate Feed Properties

S.G 16°C( 60°F) 0.7755

C* R+O 95.1

The olefin feed was fed to the reactor at 1 WHSV (ethene and propene basis) and the reformate was fed to the reactor at amounts varying from 0.6 to 4.6 moles of reformate (total reformate basis) per mole of olefin. This is equal to 1.7 to 12.8 WFSV on reformate basis (0.6 to 4.6 WHSV on Cg to C g aromatics basis).

A recycle gas stream obtained by flashing total reactor effluent at reactor pressure and ambient temperature was recirculated to the reactor inlet at 2 moles/mole of olefin feed. The reactor was maintained isothermal (+5°F). The net C- hydrocarbon yields are calculated on the basis of olefin feed to the reactor. For example the Cr yield on olefin feed is found by:

(Cf product) - (Ct reformate feed)

100 x ~ . M ,.. = _ - ζ * yield on olefin fee Olefin Feed s

The data obtained is graphically presented in Figures 2 and 3 of the drawings.

The Figure 2 of the drawings shows Cr yield on olefin feed, the effect of weight hourly space velocity (WHSV) of the ' reformate feed and the effect of varying the temperature between 149 to 371°C (300-700°F). The data show that at a given WHSV of reformate feed, increasing the temperature increases the Cr hydrocarbon yield based on olefin feed.

The Figure 3 of the drawing shows that the octane of the total C_ gasoline product increases with the reactor temperature.

The product properties are reported on the basis of the total liquid product to demonstrate the octane upgrading of the reformate feed. The data illustrated in Figures 2 and 3 are in part reported below in Table 3. The detailed process conditions and hydrocarbon feed and hydrocarbon product material balances are given in Table 3.

Table 3

Fuel Gas/Reformate Reaction Fresh 40 Alpha HZSM-5B

Hours on Stream 106.0 158.0 119.0 164. Pressure, psig _

Reactor Temperature, F 5

Gas Recycle, Mbl/Mol Fuel Gas WHSV on Reformate WHSV on Olefin Individual Feed Stream Compositions

Olefin Feed Wt .%

N,

H.

Reformate Feed Wt .%

Benzene

Toluene

Xylene

Hexane

Heptane

Octane

Raw Reformate S. G. 16 °C ( 60F)

RtO C5+ Reformate S. G. 16 °C ( 60F)

R-f

—28—

Total Hydrocarbon Feed (Olefin plus Reformate) Composition, Wt.% (H-/N 2 Free) C-i -C- Paraffins

V s c _

Benzene

Toluene

Xylene Product Properties Raw Product

SG 16°C (60F)

R- ) C5+ Product

SG 16°C (60F)

UtO

Process Yields

C5+ Yield on Reformate, Wt.

Vol%

C5+ Yield on Olefin, Wt..

C5= IC5 NC5

C6+ Total 80.9 80.6 85 .0 85.

Paraffins (Cg to C g ) Olefins (Cg to C Q ) ISO C10+ P+O+N Naphthenes

Aromatics , Total C6A C7A C8A C9A C10+A

The Table 3 data show at the preferred reaction temperatures of 600 to 700°F substantial conversion of C 2 to C, olefins to C c olefins and substantial conversion of C_-C_ 5 6 8 aromatic hydrocarbons to C g aromatic hydrocarbons and a significant increase in the octane value of the Cr hydrocarbons are obtained. The data also show that C ς paraffins are not cracked to form lighter products as the olefin conversion increases,

EXAMPLE 3

The process was carried out in a fixed bed reactor using a fresh HZSM-5 zeolite catalyst having an alpha value of 40 and a silica to alumina ratio of 70/1. The catalyst was incorporated in a silica -alumina binder at a ratio of HZSM-5/Binder of 65/35. The reaction was carried out at temperatures of 316 to 371°C (600 to 700°F) and at pressures of 790 to 2860 kPa (100 to 400 psig) .

The olefin feed was fed to the reactor at 1 WHSV (ethene/propene basis) and the reformate was fed to the reactor at 7.5 WHSV based on reformate (2.7 WHSV Cg to C g aromatics basis). A recycle gas stream obtained by flashing total reactor effluent at reactor pressure and ambient temperature was recirculated to the reactor inlet at 2 moles/mole of olefin feed. The reactor was maintained isothermal 2.8°C (+5°F).

The Cr hydrocarbon yields, as in Example 2, were calculated on the basis of olefin feed to the reactor. The data obtained is reported in Figure 4 and in Table 4.

The Figure 4 of the drawings graphically reports the data obtained and shows C r yield on olefin feed, and the effect of varying the pressure between 790 to 2860 kPa (100 and 400 psig), varying the temperature between 316 to 371°C (600 to 700°F) and feeding C or a mixture of Cl and CI in the olefin fuel gas feed stream. The data show that at a given operating pressure, increasing the temperature increases the r hydrocarbon yield. The product properties, in Table 4 are reported on the basis of the total liquid product to demonstrate the octane * upgrading of the reformate feed.

The detailed process conditions and hydrocarbon feed and product material balances are given below in Table ..

—31—

Table 4

F el Gas / Reformate Reaction

Fresh 40 Alpha HZSM-5

65/35 HZSM- 5 /Binder, 70/1 Silica/Alumina

Hours on Stream 119.0 164.0 221.0 226.0 216.5 2

Fuel Gas Feed 1 1 2 2 ' 2

Pressure, psig 250 250 100 100 400 4

Reactor Temperature, F 597 701 600 701 600 7

Gas Recycle, Mol/Mol Fuel Gas 2 2 2 2 2

WHSV on Reformate 7.5 7.5 7.5 7.5 7.5

WHSV on Olefin 0.9 1.0 0.5 0.5 0.5

Individual Feed Stream Compositions

Olefin Feed Composition ' •' Olefin Feed Wt .%

N 2 H2

C

Reformate Feed Composition Wt .% Benzene Toluene Xylene Hexane Heptane Octane C£

Raw Reformate S. G. 16 °C (60F) 0.77550.7755 0.7692 0.7692 0.7692 2

R+O 95.1 95.1 95.3 95.3 95.3 C5+ Reformate S.G. 16 °C (60F) 0.7812 0.7812 0.78290.78290.7829 9

R+0 95.1 95.1 95.0 95.0 95.0

Total Feed Composition, Wt .% (H2/N2 Free) C1-C3 Paraffins 0.1 0.1 0.9 0.9 0.9

—32-

C2= C3= C4's C5+

(1) The Feed 1 contains 53.6 mol% N2, 25 .9 mol% H 2 , 12.6 mol% C and The Feed 2 contains 61.5 mol% 2, 25.9 mol% H2 and

Product Properties Raw Product SG 16°C (60F) - 0.7745 0.7753 0.7839 0 .7835 0.7800 0.

R+O 96.2 97.0 95 .8 96.0 95.4 96.

C5+ Product

SG (60F) * . 0.7809 0.7822 0.78410.7833 0.7844 0.

R+O- 96.0 96.8 95.4 95.6 95.2 96.

Process Yields

EXAMPLE 4

This Example is described with reference to Figures 5 and 6 of the drawings.

The process was carried out in a fixed bed reactor using HZSM-5 zeolite catalyst having a silica to alumina ratio of 70/1 which was incorporated in a silica -alumina binder at a ratio of HZSM-5/Binder of 65/35. The reaction was carried out at temperatures of 329 to 427°C (625 to 800 °F) and at a pressure of 2860 kPa (400 psig) .

The olefin feed and reformate feeds are described below. Individual feed Stream Compositions Olefin Feed

N 2 H 2

C 2 C = L 3

C 3 ' Reformate Feed

Benzene 3.6

Toluene 12.6 Xylene 19.5

Hexane 16.6

Heptane 9.3

Octane 3.2

C g 27.0 The process was carried out at 625 to 800 °F (329 to 427°C) .

There were two runs carried out.

In the first run the olefin feed was fed at 1.0 WHSV olefin basis and the reformate at 1.0 WHSV, reformate basis (0.36 WHSV

Cg-Cg aromatics basis ). The data obtained is reported graphically in Figure 5 of the drawings . In the second run the olefin feed was fed at 0.5 WHSV olefin basis and the reformate at

0.5 WHSV, reformate basis (0.18 WHSV Cg-C 8 aromatics basis). The data obtained is reported graphically in Figure 6 of the drawings.

The Figure 5 of the drawings shows that the Cr hydrocarbon yield increases with temperature up to 371 to 399°C

(700-750°F) after which the yield decreases due to cracking which is undesireable.

The Figure 6 of the drawings shows that the Cr hydrocarbon yield reaches a maximum at 343°C (650°F) after which the yield decreases.

While the invention has been shown by describing preferred embodiments of the process, there is no intent to limit the inventive concept, except as set forth in the following claims.

EXAMPLE 5

The process is carried out in a fluidized bed reactor using an HZSM-5 zeolite catalyst having an acid value of.40. The reactor bed temperature is maintained at 316°C (600°F) and at a pressure of 790 kPa (100 psig). The olefin feed is fed at a WHSV of 0.5, based on ethene and propene. The reformate is fed at a WHSV of 0.8, based on Cg to Cg aromatics (0.1 based on benzene) in the reformate feed. The reaction is carried out without recycle of light olefins. The components of the olefin gas feed stream and of the reformate feed stream and the components of the total hydrocarbon feed as well as the components of the hydrocarbon product are given below.

—35-

Individual Feed Stream Compositions Olefin Gas Wt.%

Hydrogen

Ethene

Propene Reformate Wt.%

Benzene Toluene C g aromatics

Olefins Total Hydrocarbon Distribution And Yield, Wt.%

Methane . Ethene Ethane Propene 1 Propane Isobutane n-Butane Eutenes

Cr Paraffinic Hydrocarbons Cr Olefinic Hydrocarbons Aromatics Hydrocarbons

Benzene

Toluene

C g Aromatics

C g Aromatics

Product Properties

R+O Octane 95.2 97.7

Specific Gravity .775 .778

The above Example indicates substantial conversion of C 2 and C, olefins to Cr olefins and substantial conversion of Cg-Cg aromatic hydrocarbons to C g aromatic hydrocarbons .

The maximum yield C ς plus hydrocarbons and alkyl " aromatics can be achieved at a conversion temperature between 316 to

399°C (600-750°F). The flexibility of the turbulent regime fluid bed for controlling the reactor temperature under exothermic reaction conditions allows an easy adjustment for achieving the optimal yield structure. The proposed fuel gas-catalytic reformate

_ conversion unit can fit into an existing FCC gas and catalytic reforming plant refinery.

The use of a fluid bed (turbulent zone) reactor in this process offers several advantages over a fixed bed reactor. EUe to continuous catalyst regeneration, fluid bed reactor operation will not be adversely affected by oxygenate, sulfur and/or nitrogen containing contaminants presented in FCC fuel gas.

The reaction temperature can be controlled by adjusting the feed temperature so that the enthalphy change balances the heat of reaction. The feed temperature can be adjusted by a feed preheater, heat exchange between the feed and the product, or a combination of both. Cnce the feed and product compositions are determined using, for example, an on-line gas chromatograph, the feed temperature needed to maintain the desired reactor temperature, and consequent olefin and Cg to C g aromatic conversion, can be easily calculated from a heat balance of the system. In a commercial unit this can be done automatically by state-of-the-art control techniques.