Gasolines manufactured to contain a higher concentra- tion of aromatics such as benzene, toluene and xylenes

(BTX) can adequately meet the octane requirements of the marketplace for a high octane fuel. Aromatics, particularly benzene, are commonly produced in refinery processes such as catalytic reforming which have been a part of the conventional refinery complex for many years. Recent concerns about volatility and toxicity of hydrocarbon fuel and the resultant environment damage has prompted legislation that limits the content and composition of aromatic hydrocarbons in such fuels. Some of these limitations relate specifically to benzene which, due to its toxicity, will be substantially eliminated from the gasoline pool. Restrictions on the content of heavy aromatics will result from proposed end boiling point limits of gasoline fuels, referred to as T ₉₀ or (90 vol% temperature) . T ₉₀ limits curtail the presence of hydrocarbon components that boil above temperatures in a range of 177 to 221°C (350 to 430"F) .

When hydrocarbons boiling in the gasoline boiling range are reformed in the presence of a hydrogenation- dehydrogenation catalyst, a number of reactions take place which include dehydrogenation of naphthenes to form aromatics, dehydrocyclization of paraffins to form aromatics, isomerization reactions and hydrocracking reactions. The composition of the reformer effluent or reformate is shifted toward higher octane value product.

Catalytic reforming primarily increases the octane of motor

gasoline by aromatics formation but without increasing the yield of gasoline.

Reformates can be prepared by conventional techniques by contacting any suitable material such as a naphtha charge material boiling in the range of C5 or C ₆ up to 193*C (380"F) with hydrogen in contact with any conventional reforming catalyst.

Benzene content and T ₉₀ are two important specifications for reformulated gasoline (RFG) . It is known that reformate is the major benzene source in the gasoline pool. One option to reduce benzene content in the reformate is to remove the C6 cut, which is rich in benzene and its precursors, e.g. methylcyclopentanes and hexanes, from a pretreated naphtha using a dehexanizer, in order to divert these materials from the reformer. The C6 cut can be eventually combined with reformate in the gasoline pool. In order to meet the benzene specification in the gasoline pool, refineries need to convert benzene in this C6 cut, which can contain as much as 20 to 30% benzene. It would be advantageous to provide a way of treating naphtha reformer feedstock to produce gasoline of reduced benzene content and T ₉₀ specifications wherein a C6 cut is diverted from the reformer and combined with reformate, without requiring further treatment of the C6 cut to meet benzene specifications.

U.S. Pat. No. 4,927,521 to Chu discloses a process for pretreating naphtha prior to reforming, by contacting with a zeolite catalyst, e.g., zeolite beta, containing at least one noble metal and at least one alkali metal, for the purpose of producing higher yields of C4+ and C5+ gasolines.

The present invention relates to an integrated process for providing a gasoline boiling range reformate-containing product produced from naphtha. The process comprises: 1) pretreating a raw naphtha feedstream containing benzene by contacting with a hydrodesulfurization catalyst

thereby forming H2S and thereafter cascading said feedstream over a noble metal-containing porous crystalline inorganic oxide catalyst comprising pores having openings of 12-member rings under conditions sufficient to effect benzene saturation, thereby providing a pretreated effluent of reduced benzene content;

2) dehexanizing the pretreated effluent to provide a C6" hydrocarbon-containing overhead stream and a bottoms stream containing C7+ hydrocarbons; 3) reforming said bottoms stream to form a reformate stream; and

4) combining said reformate stream and said C6" hydrocarbon-containing overhead stream to provide a combined gasoline boiling range product. The present invention can be described more particularly as an integrated process for providing a gasoline boiling range reformate-containing product produced from naphtha comprising:

1) pretreating a raw naphtha feedstream by contacting with a hydrodesulfurization catalyst and thereafter cascading said feedstream over a noble metal-containing zeolite catalyst comprising pores having openings of 12- member rings thereby providing a pretreated effluent;

2) stripping off C4" hydrocarbons, hydrogen sulfide and ammonia from the pretreated effluent thereby providing a stripped effluent;

3) dehexanizing the stripped effluent to provide a C5 to C6 hydrocarbon-containing overhead stream and a bottoms stream containing C7+ hydrocarbons; 4) reforming said bottoms stream to form a reformate stream;

5) distilling said reformate stream to provide a Cl to C4 hydrocarbon-containing overhead and C5+ reformate bottoms fraction; and

6) combining said C5+ reformate bottoms fraction and said C5 to C6 hydrocarbon-containing overhead stream to provide a combined gasoline boiling range product.

The present invention relates to a process wherein a raw naphtha feed is pretreated to saturate benzene and to convert back-end materials (C9+) into lighter naphtha in an existing naphtha pretreater used for hydrodesulfurization. The process employs a noble metal promoted porous inorganic oxide catalyst downstream of the hydrodesulfurization cata- lyst. The noble metal-containing catalyst saturates benzene and reduces end-point in the presence of hydrogen sulfide produced during hydrodesulfurization. This approach eliminates or reduces the need for a secondary reactor for benzene conversion of C6 cut from dehexanizer overhead.

Inasmuch as noble metal promoted catalysts are generally sensitive to hydrogen sulfide poisoning which strongly inhibits hydrogenation activity of the noble metal, the ability of the noble metal-containing catalyst to retain its hydrogenation activity while contacting the H ₂ S-containing effluent from the dehyrosulfurization step is unexpected.

The Figure is a process flow diagram depicting a pre¬ ferred multistage embodiment of the present invention wherein raw naphtha is pretreated in two stages prior to stripping, removal of a C5 and C6 cut, reforming of the C7+ dehexanizer bottoms and combining the unreformed C5 and C6 cut with C5+ reformate. Feed The raw naphtha feedstream can comprise a mixture of aromatic and paraffin hydrocarbons having boiling points about 1.5 to 5.0 or higher mole percent benzene. It can also contain various C7 to CIO aromatic hydrocarbons including toluene and C8 to CIO aromatic hydrocarbons. The feedstream can also contain C4 to C6 paraffinic hydrocarbons including butane, isopentane, isohexane and

n-hexane which are normally present at a concentration above 5.0 mole percent. C7 to C9 paraffinic hydrocarbons such as isoheptane and isooctane can also be present. The exact composition of the raw naphtha feedstream will depend on its source. It may be formed by blending all or a portion of the effluent of several different petroleum processing units. Two such effluents are the bottoms product of the stripper column used in FCC gas concentration units and stabilized reformates which contain C6 to C9 aromatic hydrocarbons.

The raw naphtha contains sulfur. Products of catalytic cracking usually contain sulfur impurities which normally require removal, usually by hydrotreating, in order to comply with the relevant product specifications. These specifications are expected to become more stringent in the future, possibly permitting no more than 300 ppmw sulfur in motor gasolines. As a practical matter, the sulfur content will exceed 50 ppmw and usually will be in excess of 100 ppmw and in most cases in excess of 500 ppmw. For the fractions which have 95 percent points over 193°C (380'F) , the sulfur content may exceed 1,000 ppmw and may be as high as 4,000 or 5,000 ppmw or even higher, as shown below. The nitrogen content is not as characteristic of the feed as the sulfur content and is preferably not greater than 20 ppmw although higher nitrogen levels typically up to 50 ppmw may be found in certain higher boiling feeds with 95 percent points in excess of 193°C (380'F) . The nitrogen level will, however, usually not be greater than 250 or 300 ppmw. The raw naphtha feed to the process comprises a sulfur- and benzene-containing petroleum fraction which boils in the gasoline boiling range. Feeds of this type include light naphthas typically having a boiling range of

Co. to 166 ^β C (330'F), full range naphthas typically having a boiling range of C-. to 216°C (420 ^β F) , heavier naphtha fractions boiling in the range of 127°C to 211°C (260°F to

412'F), or heavy gasoline fractions boiling at, or at least within, the range of 166'C to 260'C (330'F to 500'F) , preferably 166°C to 211'C (330'F to 412'F).

The raw naphtha may be obtained from straight run distillation or from a coker or FCC unit. Alternatively, pyrolysis gasoline may be used as well. However, diene- containing streams should be treated to reduce or remove sources of gumming, as necessary.

Process Configuration Referring to the Figure, the raw naphtha feedstream 1 containing sulfur compounds, nitrogen compounds and benzene is passed to a pretreater 2 where it is treated in a first pretreating (hydrodesulfurization) zone 3 by hydrotreating the feed by effective contact of the feed with a hydrotreating catalyst, which is suitably a conventional hydrotreating catalyst, such as a combination of a Group VI and a Group VIII metal on a suitable refractory support such as alumina, e.g., Co/Mo on alumina, under hydrotreating conditions, i.e., at elevated temperature and somewhat elevated pressure in the presence of a hydrogen atmosphere. More specifically, such conditions include temperatures ranging from 204'C to 399°C (400'F to 750°F), 1-10 LHSV, 790 kPa to 5500 kPa (100-800 psig) total pressure, and a hydrogen circulation rate of 89 to 890 nl/1 (500-5000 scf/b) .

Under these conditions, at least some of the sulfur is separated from the feed molecules and converted to hydrogen sulfide, to produce a hydrotreated product and hydrogen sulfide. One suitable family of catalysts which has been widely used for this service is a combination of a Group

VIII and a Group VI element, such as cobalt and molybdenum, on a suitable substrate, such as alumina. This hydrotreated product generally has a sulfur content of less than 5 ppm S, e.g., 1 ppm S.

The effluent is cascaded to a second pretreater zone 4 which contains a noble metal-containing porous inorganic oxide catalyst having pore openings of 12-member rings. Such porous inorganic oxides have pore windows framed by 12 tetrahedral members and include but are not limited to zeolites selected from the group consisting of zeolite beta, zeolite L, zeolite X, ZSM-12, ZSM-18, ZSM-20, mordenite and boggsite, zeolite beta being preferred. Faujasites such as Rare Earth Y (REY) , Dealuminized Y (DAY) , Ultrastable Y (USY) , Rare Earth Containing

Ultrastable Y (RE-USY) , Si-Enriched Dealuminized Zeolite Y (LZ-210) (disclosed in U.S. Patents 4,711,864, 4,711,770 and 4,503,023) are also suited to use in the present invention. In an alternative embodiment, the pretreater reactor zones may comprise separate reactors with an optional interstage separator therebetween.

The catalyst can contain from 0.1 to 1 wt %, preferably from 0.3 to 0.7 wt%, noble metal selected from the group consisting of platinum, palladium, iridium, rhodium and ruthenium. Platinum is preferred as well as combinations of platinum and palladium which are resistant to sulfur poisoning.

The noble metal component is preferably dispersed on the catalyst to provide a H/noble metal ratio of at least 0.8 as measured by hydrogen chemisorption, preferably at least 1.0 H/Pt metal ratio. The hydrogen chemisorption technique indicates the extent of noble metal agglomeration of a catalyst material. Details of the analytical technique may be found in Anderson, J.R., Structure of Metallic Catalyst, Chapter 6, p. 295, Academic Press (1975) . In general, hydrogen chemisorbs selectively on the metal so that a volumetric measurement of hydrogen capacity counts the number of metal adsorption sites. Preferably the noble metal-containing catalyst has an alpha value higher than 100 and is unsteamed. The high acidity permits

operation at lower temperatures so as to minimize thermodynamic constraints on benzene saturation. Alpha value, or alpha number, is a measure of zeolite acidic functionality and is more fully described together with details of its measurement in U.S. Patent No. 4,016,218, J. Catalysis. 6_, pp. 278-287 (1966) and J. Catalysis. 61. pp. 390-396 (1980).

Process conditions in the second reaction zone depend on zeolite catalyst activity and extent of benzene saturation and back-end bottoms required by product specifications. The total pressure and hydrogen partial pressure can be in the range of those used in conventional naphtha pretreating processes, e.g. 790-5500 kPa (100-800 psig) , preferably 1135-4240 kPA (150-600 psig) total pressure. Total pressure (or hydrogen partial pressure) can be higher if more benzene saturation is desired. More specifically such conditions for the second pretreating zone can be the same as those for the first pretreating zone as discussed above. The hydrodesulfurization catalyst and the inorganic oxide catalyst of the second reaction zone can be loaded either in the same reactor or in separate reactors operating in a cascade mode without interstage operation. The effluent 5 from the second pretreater zone 4 is passed to a stripper 6 wherein ammonia, hydrogen sulfide and Cl to C4 hydrocarbons are stripped off as overhead 7. The stripper bottoms 8 are passed to a dehexanizer 9 wherein a C5 and C6 fraction 10 is taken off and a C7+ bottoms fraction 11 is passed to a reformer 12. Reforming operating conditions include temperatures in the range of from 427'C (800'F) to 538°C (1000'F), preferably from 477'C up to 527'C (890°F up to 980'F) liquid hourly space velocity in the range of from 0.1 to 10, preferably from 0.5 to 5; a pressure in the range of from about atmospheric up to 4900 kPa (700 psig) and higher, preferably from 700 kPa to 4200 kPa (100 psig to

600 psig) and a hydrogen-hydrocarbon ratio in the charge in the range from 0.5 to 20 and preferably from 1 to 10.

The reformer effluent 13 is passed to a distillation unit 14 wherein Cl to C4 hydrocarbons 15 are taken off. The reformer effluent 16 is then combined with the C5 and C6 fraction 10 taken off the dehexanizer 9 to provide a combined gasoline boiling range product 17 which contains no greater than 1 wt% benzene, preferably no greater than 0.8 wt% benzene. The combined gasoline boiling range product can exhibit enhanced end boiling point limits of gasoline fuels, referred to as T ₉₀ or (90 vol% temperature) , reflecting the presence of fewer hydrocarbon components that boil above temperatures in a range of 177*C to 221'C (350"F to 430'F) . The combined gasoline boiling range product can have a T ₉₀ of no less than 166'C (330'F), preferably no less than 177°C (350°F) . This indicates that the noble metal-containing catalyst in the second pretreating zone selectively converts C9+ bottoms into isobutane and C5-C8 gasoline-range hydrocarbons.

The following example is provided to illustrate the invention. Example

Two zeolite beta-containing catalysts containing noble metal were evaluated under conditions compatible with conventional naphtha pretreating processes. The zeolite catalysts were evaluated in a hydrodesulfurization (HDS)/zeolite catalysts system using a commercial CoMo/A1203 catalyst as the desulfurization catalyst. The experiments were conducted in a fixed-bed, down-flow, dual reactor pilot unit. The commercial HDS catalyst was loaded in the first reactor and the zeolite-containing catalyst downstream in a second reactor in a 1/2 volumetric HDS/zeolite catalyst ratio. The pilot unit was operated in a cascade mode without interstage separation to remove zeolite catalyst poisoning ammonia and hydrogen sulfide

from the first reactor effluent. The normal operating conditions were: 4.0 HSV over HDS catalyst, 2.0 LHSV over zeolite catalyst, 712 nl/1 (4000 scf/bbl) of hydrogen circulation rate, and 3900 kPa (550 psig) total pressure. The HDS catalyst was kept at a constant 343'C (650*F) while the zeolite temperature was varied from 204"C-413'C (400°F- 775 ^β F) to obtain a wide range of conversion conditions. Table 1 lists properties of the naphtha used in the experi¬ ments. The pretreated product contains ammonia and hydrogen sulfide derived from organic nitrogen- and sulfur- containing compounds in the feed to the pretreater. Such ammonia and hydrogen can be cascaded to the zeolite cracking zone. Because impurities in the feed such as hydrogen sulfide and ammonia will deactivate the reforming catalyst, feed pretreating in the form of hydrotreating is usually employed to remove these materials.

Typically, feedstock and reforming products or reformate have the following analysis set out in Table 1 below:

TABLE 1

API Gravity, 'API 54.2

Hydrogen, wt% 14.27

Sulfur, ppmw 500

Nitrogen, ppmw 10

Benzene, wt% 1.43

Distillation (D2887) ,'C ( ^• F)

IBP -18 (-i)

10% 57 (135)

50% 118 (244)

90% 206 (402)

EBP 279 (534)

The yields in the following Tables are based on raw naphtha feed.

CATALYSTS

Two catalysts were evaluated in the second pretreatment zone and their properties are set out in Table 2 below. Catalyst A was an unsteamed Pt/Beta/alumina catalyst of the present invention. Catalyst B was a steamed Mo/Beta/alumina catalyst used for comparative purposes.

TABLE 2 Catalyst A Catalyst B

Zeolite, wt% 65 65 Alumina, wt% 35 35 Platinum, wt% 0.5 Mo1ybdenum, wt% 3.6 Alpha ^* 350 110

Surface Area ^* , ² /g 459 422 n-C ₆ Sorption, cc/g 14.7 13.9 H/Pt Chemisorption 0.83 Prior to metal addition

BENZENE SATURATION The present invention serves to saturate benzene in the presence of hydrogen sulfide under naphtha pretreating conditions prior to dehexanizing. As shown in Table 3 below, the catalyst of the present invention achieved greater than 90% benzene saturation at temperatures above 288'C (550*F) while the Mo zeolite beta catalyst did not saturate the benzene at all. In fact, the Mo zeolite beta catalyst produced more benzene.

TABLE 3 Benzene Saturation Activity

Feed HDS Catalvst-A Catalvst-B

288'C 306'C 316'C 399'C Zeolite Temp. - - f550'F. t583 ' F . f600'F. C750'F.

Benzene Yld, wt.% 1.4 1.5 0.1 0.01 1.6 1.7

Benzene Conv. % -7 92 >99 -14 -21

Back-End Conversion The noble metal-containing catalyst of the present invention was found to be very active for the conversion of 149'C ⁺ (300'F ⁺ ) bottoms and for retaining products as C ₅ ⁺ gasoline-range hydrocarbons. Both zeolite-beta catalysts A and B also gave high yields of i-butane which can be alkylated with C ₃ -C ₅ olefins to high-octane alkylate. The back-end conversion comparison for the two catalysts are shown in Table 4 below.

TABLE 4

Back-End Conversion Activity

HDS onlv Catalvst-A Catalvst- -B

288'C 306°C 316'C 399'C

Zeolite Temp. - 550'F 583'F 600'F 750'F

149'C ⁺ (300'F ⁺ ) Conv., % 4.6 32 91 20 43 i-C ₄ Yield, wt.% 0.3 2.8 12.1 0.6 6.6

C_* Yield, wt.% 98.4 95.8 82.1 97.7 87.0

I

Driveability Index

T ₅₀ , T ₉₀ and driveability index (DI) of the pretreated naphtha by the HDS/zeolite-beta catalyst systems were reduced. The reformulated gasoline specification may limit T ₉₀ at 149 ^* C ⁺ (300°F ⁺ ) . The driveability index is defined as follows:

DI = 1.5 X T ₁₀ + 3.0 X T ₅₀ + T ₉₀ Some naphthas (or gasolines) have a DI of 1400 DI and lower DIs are desirable, e.g., 1300 or lower. As shown in the following table, the catalyst of the present invention is superior for the boiling point reduction and driveability index improvement:

TABLE 5 Driveability Index HDS only Catalvst-A Catalvst-B

288'C 306'C 316'C 399'C

Zeolite Temp. 550'F 583'F 600'F 750'F

149'C ⁺ (300'F ⁺ ) Conv., % 4.6 32 91 20 43

C« ^* Naphtha (based on D2887.

^■ ιo» C (°F) 63 (146) 56 (132) 29 (85) 59 (138) 48 (118) 50» ^• C (°F) 123 (253) 117 (243) 93 (199) 118 (244) 112 (233)

10 90» C (°F) 206 (402) 178 (353) 135 (275) 194 (381) 178 (353)

Driveability Index 1380 1280 1000 1320 1229

Inteαration With Reformer

To minimize benzene production in the conventional reforming complex, only C ₇ ⁺ naphtha will be upgraded. Consequently, a dehexanizer before the reforming reactors to remove benzene and its precursors (methylcyclopentane and cyclohexane) is used. The impact of reforming C ₇ * naphtha was examined using a kinetic model that simulates commercial reforming performance. The simulations were set at conditions to produce C ₅ ⁺ reformates with an octane of 100 R+O. The C ₅ -C ₆ cut produced from the HDS/Beta pretreating stage, which was removed by dehexanizer and not reformed, was then blended back into the C ₅ ⁺ reformate as "combined C ₅ ⁺ product." The following table illustrates the comparison between the two zeolite beta-containing catalysts.

TABLE 6 HDS/Zeolite/Dehexanizer/Reforming Integration

HDS only Catalvst-A Catalvst-B

Pretreater Conditions

288'C 306" 316'C 399'C

Zeolite Temp. 550'F 583' 600'F 750'F

149°C ⁺ (300°F ⁺ ) Conv., % 4.6 32 91 20 43 Integrated Process Performance

Total i-C ₄ , wt.% 1.2 3.8 12.8 1.5 7.4

10 H ₂ Product, nl/1 (scf/b) 90 (505) 91 (510) 28 (160) 93 (520) 50 (280)

Combined C, ⁺ Product

Octane, R+O 94.5 94.0 92.1 94.0 93.7

Yield, vol.% 86.5 83.4 75.7 85.8 79.2

Benzene, vol.% 1.9 0.8 0.6 2.1 2.0

15 T ₁₀ , -C CF) 33 (92) 31 (88) 28 (82) 32 (90) 28 (82)

T ₅₀ , 'C (-F) 106 (223) 104 (219) 86 (186) 104 (219) 96 (205)

T ₉₀ , 'C (-F) 159 (319) 157 (315) 133 (272) 158 (317) 156 (313)

Driveability Index 1126 1104 953 1109 1051

As shown in Table 6, the benefits found in the pretreating stage also cascade to the overall process integration. For example, benzene content, end-point, and driveability index of the combined C ₅ ⁺ product were reduced as compared to those for conventional HDS alone in the pretreater.

Incorporating a sulfur-tolerant noble metal promoted porous inorganic oxide, e.g., zeolite, catalyst into the bottom of a naphtha pretreating reactor can lower combined C ₅ * product (reformate and C ₅ -C ₆ dehexanizer overhead) end- point and benzene content for refiners with dehexanizers. Reformulated gasoline specifications limit benzene content to 1.0 % and T ₉₀ is much lower than that for conventional gasoline. Cascade HDS/zeolite followed by a dehexanizer can allow refineries to achieve benzene reduction and driveability index improvement by using a low-capital cost catalyst replacement approach. The present invention is especially attractive to refineries without C ₅ /C ₆ isomerization units. In a preferred embodiment the invention provides a process integrated into the reformer section of a refinery for the manufacture of high octane gasoline. The invention can improve the economics of meeting the benzene specifica¬ tion of the gasoline pool, preferably reducing the pool benzene content below 1% or 0.8 %.