WO2002083283A2

WO2002083283A2 - Catalyst and process for selective hydrogenation of sulfur-containing compounds

Info

Publication number: WO2002083283A2
Application number: PCT/US2002/012192
Authority: WO
Inventors: Yun-Feng Chang
Original assignee: Abb Lummus Global Inc.
Priority date: 2001-04-16
Filing date: 2002-04-16
Publication date: 2002-10-24
Also published as: AU2002309576A1; WO2002083283A3

Abstract

A calalyst for the selective hydrogenation of sulfur-containing compounds in an olefinic gasoline such as FCC cracked naphtha includes at least one noble metal on an acidic amorphous silica-alumina support containing not more than about 40 % silica.

Description

CATALYST AND PROCESS FOR SELECTIVE HYDROGENATION OF SULFUR-CONTAINING COMPOUNDS

CROSS REFERENCE TO RELATED APPLICATIONS

The present application is a continuation in part of copending U.S. application Serial No. 09/836,578 filed April 16, 2001.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to a catalyst and process for the hydrogenation of sulfur-containing compounds in an olefinic gasoline.

2. Description of the Prior Art

Gasoline stock may be produced by fluid catalytic cracking (FCC) or thermal cracking of higher boiling liquid hydrocarbon fractions. Catalytically cracked gasoline currently forms a major part of the gasoline product pool in the United States. The sulfur content of the gasoline is reduced, usually by hydrotreating, e.g., hydrodesulfurization, in order to meet product specifications or to ensure compliance with environmental regulations, both of which are expected to become more stringent in the future, possibly permitting no more than about 300 ppmw sulfur in motor gasoline. Newly enacted U.S. government regulations require sulfur in motor gasoline to be at or below 30 ppm by the year 2004. Low sulfur levels result in reduced emissions of CO, NO_x and hydrocarbons.

Naphthas and other light fractions such as heavy cracked gasoline may be hydrodesulfurized by passing the feed over a hydrodesulfurization catalyst at elevated temperature and somewhat elevated pressure in a hydrogen atmosphere. One suitable family of hydrodesulfurization catalysts which has been widely used for this operation is a combination of a Group VIII and a Group VI element, e.g., a mixture of cobalt and molybdenum, on a substrate such as alumina., silica- alumina, crystalline aluminosilicate (zeolite) , and the like. After the hydrodesulfurization operation is completed, the product may be fractionated, or simply flashed, to release by-product hydrogen sulfide and collect the now sweetened gasoline.

Cracked naphtha as it comes from the catalytic cracker and before any further processing (such as a purification operation) possesses a relatively high octane number due to its olefinic components. In some cases, this fraction may contribute up to half the gasoline in the refinery pool, thus making a significant contribution to product octane. Hydrotreating of any of the sulfur-containing fractions which boil in the gasoline boiling range causes a reduction in olefin content and consequently a reduction, in octane number. As the degree of desulfurization increases the octane number of the gasoline boiling range product decreases. Some of the hydrogen may also cause some hydrocracking as well as olefin saturation depending on the nature of the catalyst employed and the conditions of the hydrotreating operation.

Yet another alternative for the treatment of olefinic gasolines is the selective hydrogenation of sulfur-containing compounds, such as thiophene, and thiophene derivatives into less volatile compounds which are more easily removed by distillation.

There is yet a need for a selective hydrogenation catalyst with improved activity for hydrogenation of sulfur-containing compounds and higher selectivity for hydrogenation of sulfur-containing compounds over hydrogenation of olefinic hydrocarbon components of a gasoline feed.

SUMMARY OF THE INVENTION

In accordance with the present invention, a catalyst for the selective hydrogenation of sulfur- containing in an olefinic hydrocarbon feedstock in the gasoline boiling range is provided, the catalyst comprising a catalytically effective amount of at least one noble metal on an acidic amorphous silica-alumina support having a silica content of not more than about 40% by weight.

The catalyst of this invention exhibits increased activity for hydrogenation and increased selectivity for hydrogenation of sulfur-containing compounds as opposed to hydrogenation of olefinic components of a hydrocarbon feed, thereby providing improved conversion of sulfur-containing compounds in the gasoline stock with less reduction of octane number.

Without wishing to be bound, it is believed that by limiting the silica content of the silica- alumina support to no more than about 40% by weight, the affinity of the support for the noble metal components is increased with the result that the noble metal is more highly dispersed than would otherwise be the case. Higher dispersion of noble metal, in turn,' is believed to result in the improved hydrogenation results observed herein.

BRIEF DESCRIPTION OF THE DRAWINGS

Various embodiments are described below with reference to the drawings wherein:

FIG. 1 is a graph illustrating the hydrogenation performance of a PdPt/silica-alumina catalyst as a function of silica content;

FIG. 2 is a graph illustrating the strong acidity of silica-alumina as a function of silica loading;

FIG. 3 is a graph illustrating the metal dispersion of PdPt/silica-alumina catalyst as a function of silica content; and,

FIG. 4 is a graph illustrating the hydrogenation performance of PdPt/silica-alumina catalysts . DETAILED DESCRIPTION OF PREFERRED EMBODIMENT (S)

The catalyst described herein is useful for the selective hydrogenation of sulfur-containing compounds such as thiophene and thiophene derivatives (methylthiophene, ethylthiophene, and the like) into corresponding saturated compounds (e.g., tetrahydrothiophene) , which have higher boiling points. The higher boiling saturated sulfur-containing compounds are more easily removed from the olefinic gasoline by subsequent distillation, rendering a substantially sulfur free overhead product while concentrating the sulfur compounds in a bottom fraction that can subsequently be subjected to desulfurization. Use of the catalyst described herein enables a larger fraction of the original gasoline feedstock to be separated as a substantially sulfur free overhead product, while a smaller fraction needs further treatment for sulfur removal .

Selective hydrogenation of sulfur-containing compounds in an olefinic gasoline feed is improved by employing the catalyst described herein, which has a higher activity and selectivity for hydrogenation of sulfur-containing compounds than hydrogenation of olefinic hydrocarbon components of the gasoline. The catalyst can be in the form of extrudates, pellets, spheres, granules, etc. More particularly, the catalyst includes a catalytically effective amount of at least one noble metal on an acidic amorphous silica-alumina support containing not more than about 40% silica by weight .

The noble metal is selected from platinum, palladium, rhodium, ruthenium, osmium and iridium and their mixtures. While a single noble metal is effective, it is preferred to employ two or more noble metals, e.g., a mixture of palladium and platinum (PdPt) . Optionally, the catalyst may also contain one or more non-noble metal components, e.g., Cu, Zn, Mn, Re, Ag, Sn, Au, Fe, Co, Ni, K, Ca, B, P, Mo, W, Cr, or Tc.

The acidic amorphous silica-alumina catalyst support utilized herein contains not more than about 40, preferably not more than about 30, and more preferably not more than about 15, weight percent of silica. Within the foregoing limits, the support contains at least that amount of silica which will provide surface acid sites on the support. This amount of silica will ordinarily constitute not less than about 0.1, and preferably not less than about 0.5, weight percent of the silica-alumina support.

Limiting the amount of silica in the silica- alumina support in accordance with the invention has a beneficial influence on the degree of dispersion of the subsequently applied noble metal component (s) . Thus, the resulting higher levels of noble metal dispersion correlate with a greater degree of activity for the hydrogenation of sulfur-containing compounds. Without wishing to be bound, a possible explanation for this result is that it is a consequence of noble metal (s)- support interaction which affects the electronic properties of the noble metal (s).

Preferably, the silica-alumina support herein contains not more than about 0.2% carbon, not more than about 0.01% iron oxide and not more than about 0.01% sodium oxide. Preferably, the support possesses an average particle size of from about 40μ to about 60μ, a BET surface area of from about 10 m²/g to about 500 m²/g and a pore volume ranging from about 0.1 ml/g to about 0.90 ml/g.

Suitable acidic amorphous silica-alumina supports include the SIRALOX products available from Sasol GmbH, Hamburg, Germany, which are supplied in powder, extrudate, pellet, tablet and microsphere forms, all of which are useful herein. The silica in these products is substantially uniformly distributed throughout the silica-alumina matrix, an advantageous feature for the support employed in the catalyst of this invention.

The noble metal (s) can be loaded onto the support by any conventional method. For example, in the case of a PdPt combination, the support can be impregnated with aqueous solutions of platinumtetraamine nitrate, Pt (NH₃) ₄ (N0₃) ₂ and palladiumtetraamine nitrate, Pd (NH₃) ₄ (N0₃) ₂ and then dried and calcined to leave bimetallic PdPt deposited on the support. The total loading of noble metal can range from about 0.1% to about 5%, and preferably from about 0.5% to about 1.5% by weight based on total catalyst weight. When a mixture of noble metals is utilized, any ratio can be suitably employed. For example, in the case of a Pd/Pt combination, the Pd/Pt weight ratio can range from about 1/5 to about 5/1, more preferably from about 1/4 to about 4/1 and most preferably from about 2/1 to about 3.5/1. When employed, any of the optional non-noble metals referred to above can be present at levels, including their mixtures, of from about 0.05 to about 5, and preferably from about 0.1 to about 2, weight percent based on total catalyst weight.

The feed to the selective hydrogenation process of this invention is a sulfur-containing olefinic gasoline. Feeds of this type include light naphthas typically having a boiling range of about C₆ to 330°F, full range naphthas typically having a boiling range of about C₅ to 420°F, heavier naphtha fractions boiling in the range of about 260°F to 412°F, or heavy gasoline fractions boiling at, or at least within, the range of about 330°F to 500°F, preferably about 330°F to 412°F. While the most preferred feed appears at this time to be a heavy gasoline produced by catalytic cracking; or a light or full range gasoline boiling range fraction, the best results are obtained when, as described below, the process is operated with a gasoline boiling range fraction which has a 95 percent point (determined according to ASTM D 86) of at least about 325°F (163°C) and preferably at least about 350°F (177°C) , for example, 95 percent points of at least 380°F (about 193°C) or at least about 400°F (about 220°C) .

The selective hydrogenation process can be operated upon the entire gasoline fraction obtained from the catalytic cracking unit or with just a part of it, depending on the amount and the nature of the sulfur compounds present. If the front end of the cracked fraction contains relatively few sulfur components, it can be advantageous to separate the higher boiling fractions and treat them in accordance with the present process without processing the lower boiling cut. The cut point between the treated and untreated fractions can vary according to the sulfur compounds present, but usually a cut point in the range of from about 100°F (38°C) to about 300°F (150°C) , more usually in the range of about 200°F (93°C) to about 300°F (150°C), will be suitable. The exact cut point selected will depend on the sulfur specification for the gasoline product as well as on the type of sulfur compounds present: lower cut points will typically be necessary for lower product sulfur specifications.

The sulfur which is present in components boiling below about 150°F (65°C) is mostly in the form of mercaptans which can be removed by extractive type processes such as Merox but hydrotreating is appropriate for the removal of thiophene and other cyclic sulfur compounds present in higher boiling components, e.g., component fractions boiling above about 180°F (82°C) . Treatment of the lower boiling fraction in an extractive type process coupled with hydrotreating of the higher boiling component thus represents an embodiment of the present invention. Higher cut points will be preferred in order to minimize the amount of feed which is passed to the hydrotreater and the final selection of cut point together with other process options such as the extractive type desulfurization will therefore be made in accordance with the product specifications, feed constraints and other factors.

The sulfur content of these catalytically cracked fractions will depend on the sulfur content of the feed to the cracker as well as on the boiling range of the selected fraction used as the feed in the process. Lighter fractions, for example, will tend to have lower sulfur contents than the higher, boiling fractions. As a practical matter, the sulfur content of the olefinic gasoline feed herein will exceed 50 ppmw, will usually be in excess of 100 ppmw and in most cases will be in excess of about 500 ppmw. For the fractions which have 95 percent points over about 380°F (193°C) , the sulfur content may exceed about 1,000 ppmw and may be as high as 4000 or 5000 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 about 20 ppmw although higher nitrogen levels typically up to about 50 ppmw may be found in certain higher boiling feeds with 95 percent points in excess of about 380°F (193°C) . The nitrogen level will, however, usually not be greater than 250 or 300 ppmw. As a result of the cracking which has preceded the steps of the present process, the feed to the selective hydrogenation step will be olefinic, with an olefin content of at least 5 and more typically in the range of 10 to 20, e.g. 15-20, weight percent.

The selective hydrogenation of the sulfur- containing compounds in the feed is carried out with the aforedescribed hydrogenation catalyst under conditions which result in the conversion of at least some of the sulfur-containing compounds in the feed to less volatile saturated compounds to produce a product comprising a normally liquid fraction boiling in substantially the same boiling range as the feed to this step. Generally, the sulfur-containing compound to be removed is thiophene, which boils at about 84.4°C. The thiophene is converted by selective hydrogenation into tetrahydrothiophene, which boils at about 121°C. Thus, the sulfur content of the product fraction is associated with higher boiling compounds which are more easily removed by subsequent distillation.

The temperature of the hydrogenation step is suitably from about 300°F to 850°F (about 150°C to 454°C), preferably about 350°F to 800°F (about 180°C to 427°C) with the exact selection dependent on the hydrogenation desired for a given feed and catalyst. These temperatures are average temperatures and will, of course, vary according to the feed and other reaction paramenters including, for example, hydrogen pressure and catalyst activity.

The hydrogenation can be performed in any one of a variety of reactor systems such as catalytic distillation, fixed bed, ebulliated bed, fluidized bed, moving bed, slurry reactor, and the like. . In the case of a fixed bed, the reactor may contain more than one bed.

Since the feeds are usually treated without undue difficulty, low to moderate pressures may be used, typically from about 50 to 1500 psig (about 400 to 13000 kPa) , preferably about 200 to 800 psig (about 1700 to 7000 kPa) . Pressures are total system pressure. Pressure will normally be chosen to maintain the desired aging rate for the catalyst in use. The space velocity for the hydrogenation step overall is typically from about 0.1 to 50 LHSV (hr-¹) , preferably from about 0.2 to about 30, e.g., 3 to 20, LHSV(hr-¹), based on the total feed and the total catalyst volume. The hydrogen to hydrocarbon ratio in the feed is typically about 500 to 5000 SCF/Bbl (about 90 to 900 n.1.1^-1.), usually about 1000 to 2500 SCF/Bbl (about 180 to 445 n.l.l^'1.), again based on the total feed to hydrogen volumes. The extent of the hydrogenation will depend on the sulfur content of the olefinic gasoline feed and, of course, on the product sulfur specification, with the reaction parameters to be selected accordingly.

Features and advantages of the present invention are illustrated in the Examples given below. The following methods, materials and equipment were employed.

A formulated gasoline feed illustrative of an FCC olefinic gasoline was used for the measurement of catalyst activity. The formulated gasoline contained two olefins, i.e., octene-1 and 2, 4, 4-trimethylpentene-l (TMP), together with toluene, thiophene, pyridine and n- heptane. The olefins were chosen to represent olefin distribution of typical cracked naphtha gasolines, i.e., from about 10 to about 20% by weight terminal olefins and from about 80 to about 90% by weight branched olefins. Toluene was chosen to represent the aromatics contained in a typical olefinic gasoline. Thiophene was used to represent organic sulfur components contained in cracked naphtha gasoline. Pyridine was included to represent the basic components of an FCC gasoline. The amount of total olefins was varied in the range of from about 10 to about 40 weight %. The amount of aromatics was fixed to about 40 weight %. The amount of sulfur in the feed was varied in the range of from about 500 to about 2500 ppm wt . The amount of pyridine was varied in the range of from about 50 to about 250 ppm wt . Catalyst in the form of pellets or small particles mixed with a diluent, e.g., silicon carbide, was loaded into the reactor. The reactor was constructed of stainless steel (OD: " , wall thickness of 1/16", length: 8"). Typical diluent to catalyst ratio was 5-15 wt/wt.

The catalyst was positioned in between two quartz wool plugs to prevent the catalyst from being carried away.

The reactants (olefinic gasoline feed and hydrogen) were fed from the bottom. The liquid feed was delivered by an Eldex metering pump (Eldex Laboratories Inc., Napa, CA) . Hydrogen was controlled by a Brooks mass flow controller (Brooks Instrument, Hatfield, PA) .

Reactor pressure was controlled by a Mighty- Mite backpressure regulator (Grove Valve & Regulator Co., Oakland, CA) . Reaction products were analyzed by an on-line gas chromatography (GC) unit.

The feed was introduced once the reactor temperature was lowered to below 150°C when a pre- sulfidation step was used. Otherwise, the synthetic feed and hydrogen were first brought in to pre-wet the catalyst and the system pressurized to reaction pressure. The system was then heated at 0.2°C/min to the reaction temperature. The reactor temperature was raised to the desired temperature for hydrogenation. Typical hydrogenation conditions were: l-3g of catalyst, a feed rate of 0.1-2g/min, a hydrogen flow of 10-200 cc/min, a total pressure of 440 psig, and a temperature of 215°C

The activity for hydrogenation of thiophene in the present invention is expressed as first-order rate constant k_HDS=WHSV*ln [1/ (1-x) ] with unit of g-feed/g- cat/h, where "HDS" represents the selective hydrogenation of thiophene, and "x" is the factional conversion of thiophene measured from the on-line GC analysis .

Similarly, the activity for hydrogenation of trimethylpentene-1 in the present invention is expressed as first-order rate constant k_HYD=WHSV*ln [1/ (1-y) ] , with unit of g-feed/g-cat/h, where "HYD" represents the hydrogenation of 2, 4 , 4-trimethylpentene-l, and "y" is the fractional conversion of 2, 4, 4-trimethylpentene-l to 2,2, 4-trimethylpentane, the hydrogenation product determined by the on-line GC analysis. The results represent steady state performance, typically after a time-on-stream of 6-20 hours depending on catalyst and feed rate.

EXAMPLE 1 A PdPt/silica-alumina catalyst of the present invention, designated PdPt/SA-1, was prepared by incipient wetness impregnation of a silica-alumina support from Condea Vista Company, designated SA-1, containing 5 weight % Si0₂. Silica-alumina SA-1, 20g, was impregnated by incipient wetness with a solution containing 0.1195g of platinumtetraamine nitrate (Alfa Aesar, 99.9% metal purity) and 0.5369g of palladiumtetraamine nitrate (Alfa Aesar, 9-9.9% metal purity) to give 0.3 wt% Pt and 0.9 wt% Pd. The mixture was dried at 110°C for two hours before being calcined at 500°C in air for 3 hours. This gave 1.2%PdPt/AS-l . Wafers of 1mm thick and 20-30 mm in diameter made by pressurizing the catalyst powder were crushed and sieved to have particles in the range of 0.6-1.2 mm. Catalyst particles in the size range of 0.6-1.2mm were used for the hydrogenation performance test. The catalyst was tested for activity in the reactor described above. Both hydrogenation activity and selectivity were determined. The results are set forth below in Table 1 below.

EXAMPLE 2 A PdPt/silica-alumina catalyst of the present invention, designated PdPt/SA-2, was prepared by incipient wetness impregnation of a low silica silica- alumina from Condea Vista Company, designated SA-2, containing 10% of Si0₂. The SA-2 support, 20g, was impregnated with a solution containing 0.1195g of platinumtetraamine nitrate (Alfa Aesar, 99.9% metal purity) and 0.5369g of palladiumtetraamine nitrate (Alfa Aesar, 99.9% metal purity) to give 0.3 wt% Pt and 0.9 wt% Pd. The mixture was dried at 110°C for two hours and thereafter calcined at 500°C in air for 3 hours. This gave 1.2%PdPt/SA-2. Wafers of 1mm thick and 20-30 mm in diameter made by pressurizing the catalyst powder were crushed and sieved to have particles in the range of 0.6-1.2 mm. Catalyst particles in the size range of 0.6-1.2 mm were used for the hydrogenation performance test. The catalyst was tested for activity in the reactor described above. Both hydrogenation activity and selectivity were determined. The results are set forth below in Table 1 below.

EXAMPLE 3 A PdPt/silica-alumina catalyst of the present invention, designated PdPt/SA-3, was prepared by incipient wetness impregnation of a medium silica silica-alumina from Sasol GmbH, Hamburg, Germany, designated SA-3, containing 30% of Si0₂. Support SA-3, 20g, was impregnated with a solution containing 0.1195g of platinumtetraamine nitrate (Alfa Aesar, 99.9% metal purity) and 0.5369g of palladiumtetraamine nitrate (Alfa Aesar, 99.9% metal purity) to give 0.3 wt% .Pt and 0.9 wt% Pd. The mixture was dried at 110°C for two hours and thereafter calcined at 500°C in air for 3 hours. This gave 1.2%PdPt/SA-3. Wafers of 1mm thick and 20-30 mm in diameter made by pressurizing the catalyst powder were crushed and sieved to have particles in the range of 0.6-1.2 mm. Catalyst particles in the size range of 0.6-1.2 mm were used for the hydrogenation performance test. The catalyst was tested for activity in the reactor described above. Both hydrogenation activity and selectivity were determined. The results are set forth below in Table 1 below.

Table 1 below summarizes the HDS hydrogenation activity of the catalysts exemplified above as measured by the hydrogenation of thiophene as compared with the HYD hydrogenation activity as measured by the hydrogenation of 2, 4, 4-trimethylpentene (TMP).

Table 1

HDS HYD

Activity Activity Selectivity Examples Catalyst (k_HDS)* (k„_YD)* (k_HDS/k_HYD)

Ex 1 1.2%PdPt/ 13.72 3.55 3.86

SA-1

Ex 2 1.2%PdPt/ 9.87 2.38 4.15

SA-2

Ex 3 1.2%PdPt/ 6.42 1.71 3.75

SA-3

* First order rate constant for HDS hydrogenation of thiophene (g-feed/g-cat/h) and HYD hydrogenation of 2,4,4- trimethylpentene-1 (g-feed/g-cat/h) at 215°C, 440 psig.

As shown in Table 1, the examples illustrated of the catalyst of this invention exhibit greater selectivity for hydrogenation of sulfur-containing compounds (e.g., thiophene) as compared with the hydrogenation of olefinic components (trimethylpentene) of the hydrocarbon feed. The k_HDS values of the catalysts of Examples 1 to 3 are charted in relationship to the respective silica content of the catalysts in FIG. 1, discussed below. The k_HDS values of the catalysts of Examples 1 to 3 are charted in relationship to their respective metal dispersion measurements in FIG. 4, discussed below.

EXAMPLE 4 Catalyst Samples 1-6 were prepared by impregnating various silica-alumina supports with an aqueous solution of platinum tetraamine nitrate and palladium tetraamine nitrate. Catalyst Samples 1-6 were then dried at 110°C for two hours and then calcined at 500°C in air for 3 hours^'. The amount of metal solutions were chosen to produce a Pd/Pt weight ratio of about 3/1 on each support and a metals loading of either 1.22% or 0.61% by weight. The characteristics of Catalyst Samples 1-6 are set forth in Table 2 below. Table 2 Catalyst Sample

Characteristic

Silica wt% 5 10 10 30 10 10

Surface 150 200 300 200 300 300 area,m₂/g

Moisture 2.25 2 1.5 2 2 2 content, wt%

Incipient 67 57 61 68 61 ^{' '} 61 wetness wt%)

Mass 20.46 20.42 20.30 20.41 20.41 20.41 support

(wet, g)

Mass 20.00 20.01 20.00 20.00 20.00 20.00 support

(dry, g) g of 0.12 0.12 0.12 0.12 0.12 0.12 platinum salt g of 4.65 4.63 4.64 4.63 4.62 2.33 palladium salt solution

Pt content 0.06 0.06 0.06 0.06 0.06 0.03

(g)

Pd content 0.18 0.18 0.18 0.18 0.18 0.18

(g) wt% Pt 0.30 0.30 0.30 0.30- 0.30 0.15 (based on total) wt% Pd 0.91 0.91 0.91 0.91 0.91 0.46 (based on total)

Total metal 1.22 1.21 1.22 1.22 1.21 0.61 ^"load (%)

Catalyst Samples 2-6 were evaluated for the extent of noble metal dispersion as a function of the silica content of the silica-alumina support. Noble metal dispersion was measured by means of chemisorption using a Micromeritics 2010 Chemisorption unit (Micromerticis Instrument, Norcross, GA) , . the samples being degassed at 50°C for 15 minutes, then reduced in hydrogen flow at 100°C for 15 minutes. The temperature was then increased at 10°C/min from 100°C to 350°C and held at 350°C for 2 hours. After reduction in hydrogen, the system^' was degassed at 360°C for 2 hours, then cooled down to 50°C in a vacuum and held at this temperature. Dispersion was calculated based on the amount of hydrogen absorbed and the amount of Pd and Pt in the catalyst, assuming an absorption stoichiometry of one hydrogen atom per Pd or Pt atom. The results are shown in FIG. 3, discussed below. EXAMPLE 5 In order to determine the number of strong acid and weak acid sites and their ratios in various catalyst support materials, isopropylamine (IPA) was used as the probe molecule to adsorb on acidic sites. This was carried out in a DuPont 2000 Thermogravimetric Analyzer (TGA) TA Instruments, New Castle, DE . Several samples of catalyst with various amounts of silica were prepared. Each sample was tested as follows. Catalyst material, 20-120mg, in the form of powder or pellets was placed in the platinum sample pan which was located in a quartz reactor chamber. The sample was treated under a flow of gas stream of either an inert gas, e.g., high purity nitrogen, or synthetic air, at 50-100 cc/min to clean up the catalyst surface at temperatures generally below 550°C for 2-10 hours. The sample was then cooled down to no more than 300°C in a nitrogen flow. IPA was introduced by flowing nitrogen through a vaporizer containing IPA which was maintained at room temperature. The catalyst materials picked up IPA which resulted in a weight gain. When it reached a constant weight (no more absorption) the system was purged in nitrogen at lOOcc/min at 100°C for at least 50 minutes. This led to a constant weight (no more desorption) . A temperature programmed. desorption (TPD) at 10°C/min was followed to measure the amount of IPA desorbed as a function of temperature .

The results of this test are graphically presented in FIG. 2 below. The number of strong acid sites (IPA desorbed at higher than 300°C) increases with the amount of silica present up to 10 weight % and then declines at a silica loading of 30 weight % when the total surface of the material is from 150 m²/g to 200 m²/g.

EXAMPLE 6 Several catalyst supports were measured for acid sites in accordance with the method of Example 5. Catalyst Support Samples 1 to 7 represent those in accordance with the invention herein. Catalyst Support Sample 8 at 90% silica is outside the scope of the invention as are Catalyst Support Sample 9 (the support employed for the catalyst of Comparative Example III) and Catalyst Support Sample 10 which is a mixture of 70% ceria and 30% alumina. Catalyst Support Sample 1 (SA-0) is entirely alumina, i.e., it contains no silica; Catalyst Support Sample 5 (SA-5/320) contains 5 weight percent silica and has a surface area of 320m²/g; Catalyst Support Sample 6 (SA-10/360) contains 10 weight percent silica and has a surface area of 360m²/g; Catalyst Support Sample 7 (SA-30/480) contains 30 weight percent silica and has a surface area of 480 m²/g; and, Catalyst Support Sample 8 (SA-90/380) contains 90 weight percent silica and has a surface area of 380 m²/g. The results are shown below in Table 3.

Table 3

Ratio of Strong Acid Weak Acid Strong to Weak Sample Catalyst Sites Sites Acid Sites No. Support (mmol/g)¹ (mmol/g)² (SWAR)

1 SA-0 0.144 0.239 0.607

2 SA-1 0.154 0.238 0.667

3 SA-2 0.176 0.254 0.693

4 SA-3 0.102 0.152 0.667

5 SA-5/320 0.296 0.558 0.530

6 SA-10/360 0.367 0.491 0.748

7 SA-30/480 0.342 0.592 0.577

8 SA-90/380 0.159 0.930 0.171

9 20%TF-Al₂O₃ 0.272 0.251 1.07

10 (Ce0₂+Al₂0₃) 0.302 0.202 1.50

¹ Strong acid sites: corresponding to IPA desorbed at temperature over 300°C during TPD.

² Weak acid sites: corresponding to IPA desorbed at temperature below 300°C during TPD. These results show that silica-alumina has a lower strong to weak acidity ratio (SWAR) than γ-alumina or a mixed oxide of ceria and alumina. Moreover, higher silica loadings tend to lower the SWAR value. In general, the SWAR of the silica-alumina support can advantageously range from about 0.1 to about 1.0, preferably from about 0.2 to about 0.8 and more preferably from about 0.25 to about 0.75.

Features of the catalyst of the present invention exemplified by the Examples and its relationship to hydrogenation activity and selectivity are illustrated in the drawings. Referring now to FIG. 1, this graph illustrates the hydrogenation performance of bimetallic PdPt/silica-alumina catalysts (exemplified in Examples 1 to 3) as a function of silica content. As can be seen from FIG. 1, the rate constant for hydrogenation of the sulfur compound, k_HDS, decreases as the silica ^' content increases.

FIG. 2 illustrates the strong acidity of silica-alumina as a function of silica loading. The graph of Brδnsted acidity of silica-alumina as measured by isopropanolamine- temperature programmed desorption (described above) shows an increase (i.e., an increase in the number of acid sites) with the amount of silica present up to about 10%, then a relatively steep decline to the point at which silica loading is 30%, followed by a more gradual decline. While not wishing to be bound by any theory, it is suggested that this behavior may be the result of the physical limitation of the amount of silica that can be dispersed on alumina in a mono-layer fashion before forming an over-layer of silica.

Referring now to FIG. 3, which illustrates the metal dispersion of several PdPt on silica-alumina catalysts (exemplified in Example 4), as can be seen, metal dispersion decreases as silica content increases.

Referring now to FIG. 4, the relationship between hydrogenation performance and metal dispersion is illustrated. As can be seen, the chart of hydrogenation rate constants k_HDS for the catalysts of Examples 1 to 3 below shows that hydrogenation performance improves as metal dispersion increases.

While the above description contains many specifics, these specifics should not be construed as limitations on the scope of the invention, but merely as exemplifications of preferred embodiments thereof. Those skilled in the art will envision many other possibilities within the scope and spirit of the invention as defined by the claims appended hereto.

Claims

WHAT IS CLAIMED IS:

1. A catalyst for the selective hydrogenation of sulfur-containing compounds in an olefinic gasoline which comprises a catalytically effective amount of at least one noble metal on an acidic amorphous silica- alumina support having a silica content of not more than about 40% by weight.

2. The catalyst of Claim 1 having a silica content of not more than about 30% by weight.

3. The catalyst of Claim 1 having a silica content of not more than about 15% by weight.

4. The catalyst of Claim 1 wherein the noble metal is selected from the group consisting of platinum, palladium, ^■ rhodium, ruthenium, osmium, iridium and mixtures thereof.

5. The catalyst of Claim 1 wherein the total noble metal content of the catalyst ranges from about 0.1% to about 5% of the total weight of the catalyst.

6. The catalyst of Claim 1 wherein the total noble metal content of the catalyst ranges from about 0.5% to about 1.5% of the total weight of the catalyst.

7. The catalyst of Claim 1 comprising at least two noble metals.

8. The catalyst of Claim 7 comprising palladium and platinum.

9. The catalyst of claim 8 wherein the weight ratio of palladium to platinum ranges from about 1/5 to about 5/1.

10. The catalyst of Claim 8 wherein the weight ratio of palladium to platinum ranges from about 1/4 to about 4/1.

11. The catalyst of Claim 1 further including at least one non-noble metal component selected from the group consisting of Cu, Zn, Mn, Re, Ag, Sn, Au, Fe, Co, Ni, K, Ca, B, P, Mo, W, Cr and Tc.

12. The catalyst of Claim 1 wherein the silica in the silica-alumina support is substantially uniformly distributed therein.

13. The catalyst of Claim 1 wherein the silica-alumina support has a strong to weak acidity ratio of from about 0.1 to about 1.0.

14. The catalyst of Claim 1 wherein the silica-alumina support has a strong to weak acidity ratio of from about 0.2 to about 0.8.

15. The catalyst of Claim 1 wherein the silica-alumina support has a strong to weak acidity ratio of from about 0.25 to about 0.75.

16. A process for the selective hydrogenation of sulfur-containing compounds in an olefinic gasoline which comprises contacting a sulfur-containing olefinic gasoline feed under hydrogenation conditions with a hydrogenation catalyst comprising at least one noble metal on an acidic amorphous silica-alumina support having a silica content of up to about 40% by weight.

17. The process of Claim 16 wherein the gasoline stock has an olefin content of at least about 5% by weight and a sulfur content of at least about 300 ppm wt .

18. The process of Claim 16 wherein the hydrogenation conditions include a temperature of from about 300°F to about 850°F,, a pressure of from about 50 psig to about 1500 psig, a LHSV of from about 0.1 hr^"1 to about 50 LHSV hr^"1 and a hydrogen to hydrocarbon ratio of from about 500 SCF/Bbl to about 5,000 SCF/Bbl.

19. The process of Claim 16 wherein the silica-alumina support has a silica content of not more than about 30% by weight.

20. The process of Claim 16 wherein the silica-alumina support has a silica content of not more than about 15% by weight.

21. The process of Claim 16 wherein the noble metal is selected from the group consisting of platinum, palladium, rhodium, ruthenium, osmium, iridium and mixtures thereof.

22. The process of Claim 16 wherein the total noble metal content of the catalyst ranges from about 0.1% to about 5% of the total weight of the catalyst.

23. The process of Claim 16 wherein the total noble metal content of the catalyst ranges from about 0.5% to about 1.5% of the total weight of the catalyst.

24. The process of Claim 16 wherein the catalyst contains at least two noble metals.

25. The process of Claim 24 wherein the catalyst contains palladium and platinum.

26. The process of Claim 25 wherein the weight ratio of the palladium to the platinum ranges from about 1/5 to about 5/1.

27. The process of Claim 25 wherein the weight ratio of the palladium to the platinum ranges from about 1/4 to about 4/1.

28. The process of Claim 16 wherein the catalyst further includes at least one non-noble metal component selected from the group consisting of Fe, Co, Ni, Mo, W, Cr, Re, Mn and Tc.

29. The process of Claim 16 wherein the silica in the silica-alumina support is substantially uniformly distributed therein.

30. The process of Claim 16 wherein the silica-alumina support has a strong to weak acidity ratio of from about 0.1 to about 1.0.

31. The process of Claim 16 wherein the silica-alumina support has a strong to weak acidity ratio of from about 0.2 to about 0.8.

32. The process of Claim 16 wherein the silica-alumina support has a strong to weak acidity ratio of from about 0.25 to about 0.75.