This invention relates to a process for hydrotreating a hydrocarbon feed. More particularly it relates to a hydroconversion process employing a catalyst with a specified pore size distribution which achieves improved levels of hydroconversion of feedstock components having a boiling point greater than 1000°F (538°C) to products having a boiling point less than 1000°F (538°C), improved hydrodesulphurisation, particularly improved sulphur removal from the unconverted 1000°F (538°C) products, and reduced sediment make and which allows operations at higher temperatures.

As is well known to those skilled in the art, it is desirable to convert heavy hydrocarbons, such as those having a boiling point above about 1000°F (538°C), into lighter hydrocarbons which are characterised by higher economic value. It is desirable to treat hydrocarbon feedstocks, particularly petroleum residue, to achieve other goals including hydrodesulphurisation (HDS), carbon residue reduction (CRR), and hydrodemetallation (HDM) - the latter particularly including removal of nickel compounds (HDNi) and vanadium compounds (HDV).

These processes typically employ hydrotreating catalysts with specified ranges of pores having relatively small diameters (i.e. micropores, herein defined as pores having diameters less than 25θΛ) and pores having relatively large diameters (i.e. macropores, herein defined as pores having diameters greater than 250A).

One approach to developing improved catalysts for petroleum resid processing has involved enlarging the micropore diameters of essentially monomodal catalysts (having no significant macroporosities) to overcome diffusion limitations. Early petroleum distillate hydrotreating catalysts were generally monomodal catalysts with very small micropore diameters (less than say 10θA) and rather broad pore size distributions. First generation petroleum resid hydrotreating catalysts were developed by introducing a large amount of

macroporosity into a distillate hydrotreating catalyst pore structure to overcome the diffusion resistance of large molecules. Such catalysts, which are considered fully bimodal HDS/HDM catalysts, are typified by United States Patents 4,746,419, 4,395,328, 4,395,329, and 4,089,774, discussed below. U.S. 4,746,419 (Peck et a/.) discloses an improved hydroconversion process for the hydroconversion of heavy hydrocarbon feedstocks containing asphaltenes, metals, and sulphur compounds, which process minimises the production of carbonaceous insoluble solids and catalyst attrition rates. The disclosed process employs a catalyst which has 0.1 to 0.3 cc/g of its pore volume in pores with diameters greater than 120θΛ and no more than 0.1 cc/g of its pore volume in pores having diameters greater than 4000A. The present invention is distinguished from this prior art because it discloses only features of macropore size distribution useful for minimising the production of carbonaceous insoluble solids and does not disclose a pore size distribution which would provide additional hydroconversion and hydrodesulphurisation activities By contrast, the catalysts of the present invention require a unique pore size distribution in order to provide additional hydroconversion of feedstock components having a boiling point greater than 1000°F (538 °C) to products having a boiling point less than 1000°F (538 °C) and additional hydrodesulphurisation. The present invention gives improved levels of hydroconversion of feedstock components having a boiling point greater than 1000°F (538 °C) to products having a boiling point less than 1000°F (538 °C), improved hydrodesulphurisation, particularly improved sulphur removal from the unconverted 1000°F+ (538°C + ) boiling point products, and reduced sediment make at the same operating conditions and allows operations at higher temperatures compared to operations with a commercial vacuum resid hydroconversion catalyst having a macropore size distribution which satisfies the requirements of this reference.

U.S. 4,395,328 (Hensley, Jr. et a/.) discloses process for the hydroconversion of a hydrocarbon stream containing asphaltenes and a substantial amount of metals, comprising contacting the stream (in the presence of hydrogen) with a catalyst present in one or more fixed or ebullating

beds, the catalyst comprising at least one metal which may be a Group VIB or Group VIII metal, an oxide of phosphorus, and an alumina support, where the alumina support material initially had at least 0.8 cc/g of TPV in pores having diameters of 0-1200A, at least 0.1 cc/g of TPV is in pores having diameters of 1200-50,OOθΛ, a surface area in the range of 140-190 m ² /g, and the support material was formed as a composite comprising alumina and one or more oxides of phosphorus into a shaped material and was thence heated with steam to increase the average pore diameter of the catalyst support material prior to impregnation with active metals. The present invention is distinguished from this reference because the the support of the present invention does not contain one or more oxides of phosphorus, is not heated with steam to increase the average pore diameter, and requires a higher surface area of about 205-275 m ² /g and there is a much more precise definition of pore volume distribution. U.S. 4,395,329 (Le Page et a/.) discloses a hydrorefining process of a high metal-containing feedstock employing a catalyst containing alumina, a metal from group VI and a metal from the iron group, the catalyst having a Total Surface Area of 120-200 m ² /g, a Total Pore Volume of 0.8-1.2 cc/g, and a Pore Diameter Distribution whereby 0-10% of the Total Pore Volume is present as micropores with diameters less than 100A, 35-60% of the Total Pore Volume is in pores with diameters of 100-600A, and 35-55% of the Total Pore Volume is present as macropores of diameter greater than 600A. The present invention is distinguished from this reference because the prior art requires 35-55% of the TPV in pores with a diameter > 600A and the catalysts of the present invention have only about 21-27% of the PV in pores greater than 600A.

U.S. 4,089,774 (Oleck et a/.) discloses a process for the demetallation and desulphurisation of a hydrocarbon oil comprising contacting the oil with hydrogen and a catalyst, the catalyst comprising a Group VIB metal and an iron group metal (i.e. iron, cobalt, or nickel) on a porous support, and having a surface area of 125-210 m ² /g and TPV of 0.4-0.65 cc/g with at least 10% TPV in pores having diameters less than 30A, at least 50% of pore volume

accessible to mercury being in pores having diameters of 30- 150A, and at least 16.6% of pores accessible to mercury being in pores having diameters greater than 30θA. The present invention is distinguished from this reference because the prior art requires a relatively low Total Pore Volume of only 0.4-0.65 cc/g, whereas the catalysts of the present invention require much higher Total Pore Volumes of 0.82-0.98 cc/g.

U.S. 5,221 ,656, to Clark et al. discloses a hydroprocessing catalyst comprising at least one hydrogenation metal selected from the group consisting of the Group VIB metals and Group VIII metals deposited on an inorganic oxide support, said catalyst characterised by a surface area of greater than about 220 m ² /g, a pore volume of 0.23-0.31 cc/g in pores with radii greater than about 600A (i.e. in pores with diameters greater than 120θΛ), an average pore radius of about 30-70A in pores with radii less than about 600A (i.e. an average pore diameter of about 60-140A in pores with diameters less than about 1200A), and an incremental pore volume curve with a maximum at about 20-50A radius (i.e. at about 40-100A diameter). In the present invention, pores having a diameter greater than 1200Λ are only about 0.15-0.20 cc/g and the incremental pore volume curve has a maximum (i.e. Pore Mode) at 1 10-130Λ. Also, reflective of the larger range of sizes of Pore Modes, the present catalysts have much lower surface areas of 175-205 m ² /g.

A recent approach to developing improved catalysts for petroleum resid processing has involved the use of catalysts having micropore diameters intermediate between the above described monomodal HDS and HDM catalysts, as well as sufficient macroporosities to overcome the diffusion limitations for petroleum bottoms HDS (i.e. sulphur removal from hydrocarbon product of a hydrotreated petroleum resid having a boiling point greater than 1000°F (538°C)) but limited macroporosities to limit poisoning of the interiors of the catalyst particles. Catalysts with micropore diameters intermediate between the above described monomodal HDS and HDM catalysts with limited macroporosities include those of United States Patents 4,941 ,964, 5,047, 142 and 5,399,259 and copending United States Patent Application Serial

No. 08/425,971 , which is a divisional of United States Patent 5,435,908, discussed below.

U.S. 4,941 ,964 discloses a process for the hydrotreatment of a sulphur- and metal-containing feed which comprises contacting said feed with hydrogen and a catalyst in a manner such that the catalyst is maintained at isothermal conditions and is exposed to a uniform quality of feed, the catalyst comprising an oxide of a Group VIII metal, an oxide of a Group VIB metal and 0-2.0 weight % of an oxide of phosphorus on a porous alumina support, having a surface area of 150-210 m ² /g and a Total Pore Volume (TPV) of 0.50-0.75 cc/g such that 70-85% TPV is in pores having diameters of 100-160Λ and 5.5-22.0% TPV is in pores having diameters of greater than 25θΛ.

U.S. 5,047,142 discloses a catalyst composition useful in the hydroprocessing of a sulphur and metal-containing feedstock comprising an oxide of nickel or cobalt and an oxide of molybdenum on a porous alumina support in such a manner that the molybdenum gradient of the catalyst has a value of less than 6.0 and 15-30% of the nickel or cobalt is in an acid extractable form, having a surface area of 150-210 m ² /g, a Total Pore Volume (TPV) of 0.50-0.75 cc/g, and a pore size distribution such that less than 25% TPV is in pores having diameters less than 100A, 70.0-85.0% TPV is in pores having diameters of 100-160A and 1.0-15.0% TPV is in pores having diameters greater than 25θΛ.

U.S. 5,399,259 discloses a process for the hydrotreatment of a sulphur-, metals- and asphaltenes-containing feed which comprises contacting said feed with hydrogen and a catalyst in a manner such that the catalyst is maintained at isothermal conditions and is exposed to a uniform quality of feed, the catalyst comprising 3-6 wt % of an oxide of a Group VIII metal, 14.5-24 wt % of an oxide of a Group VIB metal and 0-6 wt % of an oxide of phosphorus on a porous alumina support, having a surface area of 165-230 m ² /g and a Total Pore Volume (TPV) of 0.5-0.8 cc/g such that less than 5% of TPV is in pores with diameters less than about 80A, at least 65% of the pore volume in pores with diameters less than 25θA is in pores with diameters within ± 20A of a Pore Mode of about 100-135Λ and 22-29% TPV is in pores having diameters

of greater than 250A. The present invention is distinguished from this reference because the prior art requires a relatively low Total Pore Volume of only 0.5-0.8 cc/g and a relatively low macroporosity of 22-29% TPV in pores having diameters of greater than 250A. By contrast, the catalysts of the present invention require much higher Total Pore Volumes of 0.82-0.98 cc/g and a much higher level of macroporosity of 29.6-33.0% TPV in pores having diameters of greater than 25θA.

In United States Patent Application Serial No. 08/425,971 there is disclosed a hydrotreating process employing, as catalyst, a porous alumina support with pellet diameters of 0.032-0.038 inches (0.81-0.96 mm) bearing 2.5-6 w % of a Group VIII non-noble metal oxide, 13-24 w % of a Group VIB metal oxide, less than or equal to 2.5 w % of silicon oxide, typically about 1.9- 2 w % of intentionally added silica oxide, and 0-2 w % of a phosphorus oxide, preferably less than 0.2 w % of a phosphorus oxide, with no phosphorus- containing components intentionally added during the catalyst preparation, said catalyst having a Total Surface Area of 165-210 m ² /g, a Total Pore Volume of 0.75-0.95 cc/g, and a Pore Diameter Distribution whereby 14-22% of the Total Pore Volume is present as macropores of diameter > 100θA, 22-32% of the Total Pore Volume is present as pores of diameter ≥25θA, 68-78% of the Total Pore Volume is present as pores of diameter <25θA, 26-35% of the Total Pore Volume is present as mesopores of diameters 20θA, 34-69% of the Total Pore Volume is present as secondary micropores of diameters 100- 20θA, 5-18% of the Total Pore Volume is present as primary micropores of diameter 10OA, and ≥ 57% of the micropore volume is present as micropores of diameter within ±2θA of a pore mode of 100-145A. By contrast, the present invention employs, as catalyst, a porous alumina support with pellet diameters of 0.032-0.044 inches (0.81-1.12 mm), preferably 0.039-0.044 inches (0.99-1.12 mm), bearing 2.2-6 w % of a Group VIII non-noble metal oxide, 7-24 w % of a Group VIB metal oxide, less than or equal to 2.5 w % of silicon oxide, preferably 1.3-2.5 w % of intentionally added silica oxide, and 0.3-2 w % of a phosphorus oxide, preferably 0.5-1.5 w % of a phosphorus oxide, with phosphorus-containing components intentionally added during the

catalyst preparation, said catalyst having a Total Surface Area of 175-205 m ² /g, a Total Pore Volume of 0.82-0.98 cc/g, and a Pore Diameter Distribution whereby 29.6-33.0% of the Total Pore Volume is present as macropores of diameter greater than 25θA, 67.0-70.4% of the Total Pore Volume is present as micropores of diameter less than 25θA, >65% of the micropore volume is present as micropores of diameter within ±2δA of a pore mode by volume of 110-130A, and less than or equal to 0.05 cc/g of micropore volume is present in micropores with diameters less than 8θA.

Another recent approach to developing improved catalysts for the hydroconversion of feedstock components having a boiling point greater than 1000°F (538°C) to products having a boiling point less than 1000°F (538°C) has involved the use of catalysts having micropores intermediate between the above described monomodal HDS and HDM catalysts with pore volumes in the HDS type of range and with macroporosities sufficient to overcome the diffusion limitations for conversion of feedstock components having boiling points greater than 1000°F (538°C) into products having boiling points less than 1000°F (538 °C), but with limited macroporosities to limit poisoning of the interiors of the catalyst particles. Such catalysts are described in United States Patent 5,397,456 and copending United States Patent Application Serial No. 08/130,472 discussed below.

U.S. 5,397,456 discloses a catalyst composition useful in the hydroconversion of a sulphur- and metal-containing feedstock comprising an oxide of a Group VIII metal and an oxide of a Group V-IB metal and optionally phosphorus on a porous alumina support, the catalyst having a Total Surface Area of 240-310 m ² /g, a Total Pore Volume of 0.5-0.75 cc/g, and a Pore Diameter Distribution whereby 63-78% of the Total Pore Volume is present as micropores of diameter 55-11δA and 1 1-18% of the Total Pore Volume is present as macropores of diameter greater than 25θA. The heavy feedstocks are contacted with hydrogen and with the catalyst. The catalyst is maintained at isothermal conditions and is exposed to a uniform quality of feed. The process is particularly effective in achieving desired levels of hydroconversion of feedstock components having a boiling point greater than 1000°F (538°C)

to products having a boiling point less than 1000°F (538 °C). The present invention is distinguished from this reference because the prior art requires a catalyst with a Pore Diameter Distribution wherein 63-78% of the Total Pore Volume is present as micropores of diameter 55-115 A and 1 1 -18% of the Total Pore Volume is present as macropores of diameter greater than 25θA, whereas the catalysts employed in the present invention have only about 20-25% of the Total Pore Volume present as micropores of diameter 55-1 15A and 29.6- 33.0% of the Total Pore Volume is present as macropores of diameter greater than 25θA. In United States Patent Application Serial No. 08/130,472 there is disclosed a hydrotreating process and catalyst wherein 50-62.8% of the TPV is present in micropores of diameter 55-11 δA and 20-30.5% of the TPV is present as macropores of diameter greater than 25θA. In the instant case, the catalyst preferably has only about 20-25% of the TPV present in pores having diameter of 55-11 δA.

None of the above-identified catalyst types in the art have been found to be effective for achieving all of the desired improved process needs. Early catalysts in the art addressed the need for improved hydro-desulphurisation and/or hydrodemetallation as measured in the total liquid product. One recent line of catalyst development has been to develop improved catalysts for petroleum bottoms HDS (i.e. selective sulphur removal from the unconverted hydrocarbon product having a boiling point greater than 1000°F (538°C) from a hydroprocess operating with significant hydroconversion of feedstocks components [e.g. petroleum resids] having a boiling point greater than 1000°F (538°C) to products having a boiling point less than 1000°F (538°Q). More recent developments of petroleum bottoms HDS catalysts have been aimed at developing petroleum bottoms HDS catalysts with a degree of sediment control allowing the use of higher temperatures and reducing sediment make. However, none of the above-described petroleum bottoms HDS catalysts give improved levels of hydroconversion of feedstocks components having a boiling point greater than 1000°F (538°C) to products having a boiling point less than 1000°F (538°C) while, at the same time, reducing sediment make.

A further line of catalyst development has been to develop hydroconversion catalysts for the improved hydroconversion of feedstocks components having a boiling point greater than 1000°F (538°C) to products having a boiling point less than 1000°F (538°C). The most recent developments have led to hydroconversion catalysts with slightly improved bottoms HDS activities and some slight degree of sediment control allowing the use of some higher temperatures and reducing sediment make. Although the above-described hydroconversion catalysts give improved levels of hydro- conversion of feedstocks components having a boiling point greater than 1000°F (538°C) to products having a boiling point less than 1000°F (538°C), they do not give the desired levels of sulphur removal obtained from the above- described petroleum bottoms HDS catalysts and these hydroconversion catalysts still make some amount of sediment.

It would be desirable if a catalyst were available which provided improved hydroconversion, improved bottoms HDS, and no sediment make and which could also withstand operation at higher temperatures, so that it would be possible to attain an even higher level of hydroconversion without the undesirable formation of sediment. Undesirable low levels of hydroconversion represent a problem which is particularly acute for those refiners who operate vacuum resid hydroprocessing units at their maximum heat and/or temperature limits. Such limits often exist when refiners are operating at maximum charge rates.

It is an object of this invention to provide a process for hydroconverting a charge hydrocarbon feed, particularly, to hydroconvert feedstock components having boiling points greater than 1000°F (538°C) into products having boiling points less than 1000°F (538 °C) while simultaneously removing high amounts of sulphur from the unconverted 1000°F+ (538°C + ) product stream. It is also an object of this invention to provide improved conversion at low Existent IP Sediment values in the 650°F+ boiling point product (Discussed below under "Sediment Measurement"). It is also an object of this invention to allow the use of higher operating temperatures with reduced sediment make.

In accordance with certain of its aspects, this invention is directed to a process for hydroprocessing a charge hydrocarbon feed containing components boiling above 1000°F (538°C), and sulphur, metals, and carbon residue which process comprises: contacting said charge hydrocarbon feed with hydrogen at isothermal hydroprocessing conditions in the presence of, as catalyst, a porous alumina support containing <2.5 wt % of silica and bearing 2.2-6 wt % of a Group VIII metal oxide, 7-24 wt % of a Group VIB metal oxide, and 0.3-2 wt % of a phosphorus oxide, said catalyst having a Total Surface Area of 175-205 m ² /g, a Total Pore Volume of 0.82-0.98 cc/g, and a Pore Diameter Distribution whereby 29.6-33.0% of the Total Pore Volume is present as macropores of diameter greater than 25θA, 67.0-70.4% of the Total Pore Volume is present as micropores of diameter less than 25θA, >65% of the micropore volume is present as micropores of diameter within ±2δA of a pore mode by volume of 110-130A, less than 0.05 cc/g of micropore volume is present in micropores with diameters less than 8θA, thereby forming hydroprocessed product containing decreased content of components boiling above 1000°F (538°C) and sulphur, metals, and carbon residue; and recovering said hydroprocessed product containing decreased content of components boiling above 1000°F (538°C), and of sulphur, metals, and carbon residue, recovering said hydroprocessed product containing decreased content of sulphur in the portion of the hydroprocessed product boiling above 1000°F (538°C), and recovering said hydroprocessed product containing decreased content of sediment in the portion of the hydroprocessed product boiling above 650°F (343 °C).

The catalyst of the present invention allows operation at about + 10°F ( + 5.6°C) and about + 8 wt % 1000°F (538°C) conversion compared to operations with a first generation H-OIL catalyst. This constitutes a substantial economic advantage.

DESCRIPTION OF THE INVENTION Feedstock

The hydrocarbon feed which may be charged to the process of this invention may include heavy, high boiling petroleum cuts typified by gas oils, vacuum gas oils, petroleum cokes, residual oils, vacuum resids, etc. The process of this invention is particularly useful to treat high boiling oils which contain components boiling above 1000°F (δ38°C) to convert them to products boiling below 1000°F (538°C). The charge may be a petroleum fraction having an initial boiling point of above 650°F (343°C) characterised by the presence of an undesirable high content of components boiling above 1000°F (538°C), and sulphur, carbon residue and metals; and such charge may be subjected to hydrodesulphurisation (HDS). In particular, the charge may be undiluted vacuum resid.

A typical charge which may be utilised is an Arabian Medium/Heavy Vacuum Resid having the properties shown in Table I below:

It is a particular feature of the process of this invention that it may permit treating of hydrocarbon charge, particularly those containing components boiling above about 1000°F (δ38°C), to form product which is characterised by an increased content of components boiling below 1000°F (538 °C) and by decreased content of undesirable components typified by sulphur, metals, and carbon residue. It is another feature of the process of the present invention that it provides improved sulphur removal from the unconverted 1000°F (δ38°C) products. It is another feature of the process of the present invention that it provides the above mentioned improvements with little or no sediment formation as measured by the Existent IP Sediment values of the 650°F+ (343°C + ) boiling point product. It is another feature of the process of the present invention that it allows operations at higher temperatures with consequent higher levels of 1000°F+ to 1000°F- (538°C + to 538 °C-) than may be achieved with the use of first generation catalysts.

Sediment Measurement

It is a particular feature of the catalyst of this invention that it permits operation to be carried out under conditions which yield a substantially decreased content of sediment in the product stream leaving hydrotreating.

The charge to a hydroconversion process is typically characterised by a very low sediment content of 0.01 % by weight (wt %) maximum. Sediment is typically measured by testing a sample using the Shell Hot Filtration Solids Test (SHFST) - See Jour. Inst. Pet. (1951 ) 37 pages 596-604 Van Kerknoort

et al. Typical hydroprocessing methods in the art commonly yield Shell Hot Filtration Solids of above about 0.17 wt % and as high as about 1 wt % in the 650°F+ (343°C + ) product recovered from the bottoms flash drum (BFD). Production of large amounts of sediment is undesirable in that it results in deposition in downstream units. In due course, the deposits must be removed. This of course requires that the unit be shut down for an undesirable long period of time. Sediment is also undesirable in the products because it deposits on and inside various pieces of equipment downstream of the hydroprocessing unit and interferes with proper functioning of pumps, heat exchangers, fractionating towers, etc.

Very high levels of sediment formation (e.g. 1 wt %), however, are not usually experienced by those refiners who operate vacuum resid hydroprocessing units at or near their maximum heat and feedstock charge rates. Such units are generally operating at moderate conversion levels of feedstock components having boiling points greater than 1000°F (538°C) into products having boiling points less than 1000°F (538°C) (say, 40-6δ volume percent - vol% - conversion) and at relatively low but still undesirable values of sediment (e.g. 0.17 wt %).

In the present invention the IP 375/86 test method for the determination of total sediment has been very useful. The test method is described in ASTM Designation D 4870-92. The IP 375/86 method was designed for the determination of total sediment in residual fuels and is very suitable for the determination of total sediment in the 650°F+ (343°C + ) boiling point product. The 650° F+ (343 °C + ) boiling point product can be directly tested for total sediment, which is designated as the "Existent IP Sediment value."

It has been found that the Existent IP Sediment Test gives essentially equivalent test results as the Shell Hot Filtration Solids Test described above.

However, it has been noted that even 650°F+ (343°C + ) boiling point products that give low Existent IP Sediment values may produce additional sediment upon storage. Thus, a more severe test for sediment has been developed. In this modified test, 60 grams of 6δ0°F+ (343°C + ) boiling point product are heated to about 90 °C and mixed with about 5 cm ³ of reagent

grade hexadecane. The mixture is aged for about one hour at about 100°C. The resultant sediment is then measured by the IP 375/86 test method. The values obtained from this modified test are designated the "Accelerated IP Sediment values." 5 As it is recommended that the IP 375/86 test method be restricted to samples containing less than or equal to about 0.4 to 0.5 wt % sediment, sample size is reduced when high sediment values are observed. This leads to fairly reproducible values for even those samples with very large sediment contents. 0 It will be noted that catalysts of this invention, characterised by (i) about

0.15-0.20 cc/g of pores in the > 120θA range, (ii) about 21-27% of TPV in pores in the >60θA range, (iii) 29.6-33.0% of the TPV in pores having a diameter of ≥ 25θA, (iv) 67.0-70.4% of the TPV in micropores of diameter less than 2δθA, (v) 65% of the micropore volume in micropores of diameter 5 within ± 25 A about a pore mode by volume of 110-130A, (vi) about 20-25% of the TPV in pores having a diameter of 55-1 1 δA, and (vii) less than 0.05 cc/g micropore volume in micropores with diameters less than 8θA, are particularly advantageous in that they permit attainment of product hydrocarbon streams containing the lowest content of sediment at highest conversion, while 0 producing product characterised by low sulphur, carbon residue and metals contents. It is a feature of the catalysts of this invention that they permit attainment of hydrotreated product with <0.15 wt % sediment, as measured by the Existent IP Sediment test in the portion of the hydroprocessed product boiling above 6δO°F (343°C), typically as low as 0.0-0.1 wt %, preferably 0.0- δ O.Oδ wt %, say 0.05 wt %.

Reaction Conditions

In the practice of the process of this invention (as typically conducted in a single-stage Robinson reactor in pilot plant operations), the charge 0 hydrocarbon feed is contacted with hydrogen at isothermal hydrotreating conditions in the presence of catalyst. Pressure of operation may be 1500-

10,000 psig (10.4-69 MPa), preferably 1800-2500 psig (12.4-17.3 MPa), say

1 δ

2260 psig (1 δ.δ MPa). Hydrogen is charged to the Robinson Reactor at a rate of 2000-10,000 SCFB (360-1800 m ³ /m ³ ), preferably 3000-8000 SCFB (640- 1440 m ³ /m ³ ), say 7000 SCFB (1260 m ³ /m ³ ). Liquid Hourly Space Velocity (LHSV) is typically 0.1 -1.5, say 0.56 volumes of oil per hour per volume of 5 liquid hold-up in the reactor. Temperature of operation is typically 700-900°F (371-482°C), preferably 750-87δ°F (399-468°C), say 760°F (404°C). Operation is essentially isothermal. The temperature may typically vary throughout the bed by less than about 20°F (11 °C).

In another more preferred embodiment of the process of the present 0 invention, the liquid and gaseous effluent from the previously described first- stage Robinson reactor is routed to a second-stage Robinson reactor containing the same weight of catalyst as had been loaded to the first-stage Robinson reactor and which is operated at essentially the same temperature and pressure as the first-stage Robinson reactor. The difference in average temperature δ between the first- and second-stage reactors is 0-30°F (0-16.7°C), preferably 0-1 δ°F (0-8.3°C), say 0°F (0°C). No additional hydrogen is normally injected to the second-stage Robinson reactor. The liquid effluent passes through the second-stage Robinson reactor at a similar LHSV to that of the first-stage Robinson reactor. The liquid effluent from the first-stage Robinson reactor is 0 uniformly contacted with the hydrogen-containing gaseous effluent and the second loading of catalyst at isothermal conditions in the second-stage Robinson reactor. No attempt is made to maintain constant catalytic activity by periodic or continuous withdrawal of portions of used catalyst and replacement of the withdrawn material with fresh catalyst in the two-stage δ Robinson reactor system. The catalyst begins as fresh catalyst and accumulates catalyst age generally expressed in barrels per pound. The average temperature is defined as the average of the temperatures of the first- and second-stage reactors. Average temperature of operation is typically 700- 900°F (371-482°C), preferably 7δ0-87δ°F (399-468°C), say 760°F (404°C). 0 Overall, the hydrocarbon charge passes through the entire process system (i.e. the first- and second-stage Robinson reactors) at an overall LHSV of 0.0δ-0.7δ, say 0.28 volumes of oil per hour per volume of liquid hold-up in the reactor.

In general, reaction may be carried out in one or more continuously stirred tank reactors (CSTRs), such as Robinson reactors, in which the catalyst is exposed to a uniform quality of feed.

In one particularly preferred embodiment of the present invention, a δ sulphur-and metal-containing hydrocarbon feedstock is catalytically hydroprocessed using the H-OIL (TM) Process configuration. H-OIL is a proprietary ebullated bed process (co-owned by Hydrocarbon Research, Inc. and Texaco Development Corporation) for the catalytic hydrogenation of residua and heavy oils to produce upgraded distillate petroleum products and 0 an unconverted bottoms product particularly suited for blending to a low sulphur fuel oil. The ebullated bed system operates under essentially isothermal conditions and allows for exposure of catalyst particles to a uniform quality of feed.

In the H-OIL Process, a catalyst is contacted with hydrogen and a δ sulphur- and metal-containing hydrocarbon feedstock by means which insures that the catalyst is maintained at essentially isothermal conditions and exposed to a uniform quality of feed. Preferred means for achieving such contact include contacting the feed with hydrogen and the catalyst in a single ebullated bed reactor, or in a series of two to five ebullated bed reactors, a series of two 0 ebullated bed reactors being particularly preferred. This hydroprocessing method is particularly effective in achieving high levels of hydrodesulphurisation with vacuum residua feedstocks.

In the H-OIL Process, the hydrocarbon charge is admitted to the first- stage reactor of a two-stage ebullated bed H-OIL unit in the liquid phase at 6 6δ0-8δ0°F (343-4δ4°C), preferably 700-826 °F (371-441 °C) and 1000-3600 psia (6.9-24.2 MPa), preferably 1600-3000 psia (10.4-20.7 MPa). Hydrogen gas is admitted to the first-stage reactor of a two-stage ebullated bed H-OIL unit in amount of 2000-10,000 SCFB (360-1800 m ³ /m ³ ), preferably 3000-8000 SCFB (640-1440 m ³ /m ³ ). The hydrocarbon charge passes through the first- 0 stage ebullated bed reactor at a LHSV of 0.16-3.0 hr\ preferably 0.2-2 hr ^"1 . During operation, the catalyst bed is expanded to form an ebullated bed with a defined upper level. Operation is essentially isothermal with a typical

maximum temperature difference between the inlet and outlet of 0-60°F (0- 27.8°C), preferably 0-30°F (0-16.7°C). The liquid and gaseous effluent from the first-stage reactor is then routed to the second-stage reactor of the two- stage H-OIL unit which is operated at essentially the same temperature and pressure as the first-stage reactor. The difference in average temperature between the first- and second-stage reactors is 0-30°F (0-16.7°C), preferably 0-15 °F (0-8.3 °C). Some additional hydrogen may optionally be injected to the second-stage reactor to make up for the hydrogen consumed by reactions in the first-stage reactor. In the H-OIL process, constant catalytic activity is maintained by periodic or continuous withdrawal of portions of used catalyst and replacement of the withdrawn material with fresh catalyst. Fresh catalyst is typically added at the rate of 0.05-1.0 pounds per barrel of fresh feed, preferably 0.20-0.40 pounds per barrel of fresh feed. An equal volume of used catalyst is withdrawn and discarded to maintain a constant inventory of catalyst on the volume basis. The catalyst replacement is performed such that equal amounts of fresh catalyst are added to the first-stage reactor and the second-stage reactor of a two-stage H-OIL unit.

Catalyst Support

The catalyst support is alumina. Although the alumina may be alpha, beta, theta, or gamma alumina, gamma alumina is preferred.

The charge alumina which may be employed in practice of this invention may be available commercially from catalyst suppliers or it may be prepared by variety of processes typified by that wherein 85-90 parts of pseudoboehmite alumina is mixed with 10-15 parts of recycled fines. Silica (SiO ₂ ) may be incorporated in small amounts typically up to about 2.5 wt % on the finished catalyst basis, and preferably 1.3-2.5 wt % on the finished catalyst basis. Acid is added and the mixture is mulled and then extruded in an Auger type extruder through a die having cylindrical holes sized to yield a calcined substrate having a diameter of 0.032-0.044 inches (0.81-1.1 mm), preferably 0.039-0.044 inches (0.99-1.1 mm). Extrudate is air-dried to a final

temperature of typically 250-275°F (121-135°C) yielding extrudates with 20- 26% of ignited solids. The air-dried extrudate is then calcined in an indirect fired kiln for O.δ-4 hours in an atmosphere of air and steam at typically about 1000-1 150°F (538-621 °C).

Catalysts of the Present Invention - Pore Size Distribution

The catalyst which may be employed is characterised by Total Surface

Area (TSA), Total Pore Volume (TPV), and Pore Diameter Distribution (Pore Size

Distribution, PSD). The Total Surface Area is 176-205 m ² /g, preferably 175- 195 m ² /g, say 178 m ² /g. The total Pore Volume (TPV) may be 0.82-0.98, preferably 0.82-0.90, say 0.83 cc/g.

Less than 0.05 cc/g of micropore volume is present in micropores with diameters less than 8θA.

Micropores of diameter in the range of less than 25θA are present in an amount of about 67.0-70.4% of the Total Pore Volume, preferably 67.0-69.1

%TPV, say 67.0 %TPV. Preferably 6δ% of the micropore volume is present as micropores of diameter within ±2δA of a pore mode by volume of 110-

130A.

The amount of Total Pore Volume in the range of δδ-1 1 δA is only about 20-26% and preferably 20.8%.

The Pore Size Distribution is such that 29.6-33% of the Total Pore Volume, and preferably about 33.0% is present as macropores of diameter greater than 2δθA.

The amount of Total Pore Volume in pores with a diameter greater than 60θA is only about 21-27% and preferably 26.6 %TPV.

The amount of Total Pore Volume in pores having a diameter greater than 120θA is only about 0.16-0.20 cc/g and preferably 0.20 cc/g.

It should be noted that the percentages of the pores in the finished catalyst are essentially the same as in the charge gamma alumina substrate from which it is prepared, although the Total Surface Area of the finished catalyst may be 75-85%, say 80% of the charge gamma alumina substrate from which it is prepared (i.e. 75-86% of a support surface area of 205-

276 m ² /g, say 221 m ² /g). It should also be noted that the Median Pore Diameter by Surface Area from mercury porosimetry of the finished catalyst is essentially the same as that of the charge gamma alumina substrate from which it is prepared, It is also noted that the Pore Size Distribution (percent of total) in the finished catalyst may be essentially the same as in the charge alumina from which it is prepared (unless the majority of the pore volume distribution in a given range is near a "break-point" - e.g. 8θA or 2δθA, in which case a small change in the amount of pores of a stated size could modify the reported value of the pore volume falling in a reported range). The Total Pore Volume of the finished catalyst may be 75%-98%, say 80% of the charge alumina from which it is prepared.

Catalysts of the Present Invention - Metals Loadings The alumina charge extrudates may be loaded with metals to yield a product catalyst containing a Group VIII non-noble metal oxide in an amount of 2.2-6 wt %, preferably 3.0-3.6 wt %, say 3.3 wt % and a Group VIB metal oxide in an amount of 7-24 wt %, preferably 12.6-15.6 wt %, say 14.4 wt %.

The Group VIII metal may be a non-noble metal such as iron, cobalt, or nickel. This metal may be loaded onto the alumina typically from a 10%-30%, say 15% aqueous solution of a water-soluble salt (e.g. a nitrate, acetate, oxalate etc.). The preferred metal is nickel, employed as about a 11.3 wt % aqueous solution of nickel nitrate hexahydrate Ni(NO ₃ ) ₂ .6H ₂ O.

The Group VIB metal can be chromium, molybdenum or tungsten. This metal may be loaded onto the alumina typically from a 10%-40%, say 20% aqueous solution of a water-soluble salt. The preferred metal is molybdenum, employed as about a 15.5 wt % aqueous solution of ammonium molybdate tetrahydrate (NH ₄ ) ₆ Mo ₇ O ₂₄ .4H ₂ O.

It is a feature of the catalyst of the invention that it contains about 0.3-2 wt % of P ₂ 0 ₅ and preferably about 0.5-1.6 wt %. This level of phosphorus oxide loading is very small representing only 0.13-0.87 wt % of elemental phosphorus and preferably 0.22-0.87 wt % of elemental phosphorus. The

phosphorus component may be loaded onto the alumina as a 0-4 wt %, say 1.1 wt % aqueous solution of 85 wt % phosphoric acid H ₃ PO ₄ in water.

As described above, silica SiO ₂ may be incorporated into the catalyst supports prior to impregnation and may therefore be present in small amounts, typically up to about 2.δ wt % and preferably 1.3-2.6 wt %, although the benefits of the invention may be attained without addition of silica.

These catalyst metals and phosphorus may be loaded onto the alumina support by impregnating the latter with a solution of the former. Although it is preferred to load the metals simultaneously, it is possible to load each separately. Small amounts of H ₂ O ₂ may be added to stabilise the impregnating solution. It is preferred that the catalyst be impregnated by filling 90-106%, preferably 97-98%, say 97% of the substrate pore volume with the solution containing the requisite amounts of metals and phosphorus. Loading is followed by drying and calcining at 900-12δO°F (482-677°C), preferably 1150-1210°F (621 -654°C), say 1180°F (638°C) for 0.5-δ hours, say 1 hour.

Another feature of the catalyst composition of the present invention is that the ratio of the measured hydrodesulphurisation (HDS) microactivity rate constant k of the catalyst of the present invention to the measured HDS microactivity rate constant k of a standard hydroprocessing catalyst (namely, Criterion HDS-1443B, a commercially available, state-of-the-art catalyst for use in hydroprocessing resid oils), has a value of O.δ-1.0, preferably 0.6-0.86. As used in this description, the phrase "HDS microactivity" means the intrinsic hydrodesulphurisation activity of a catalyst in the absence of diffusion, as measured according to the HDS Microactivity (HDS-MAT) Test, described as follows. In the HDS-MAT Test, a given catalyst is ground to a 30-60 mesh (0.0071-0.013 mm) fraction and presulphided at 750°F (399°C) with a gas stream containing 10% H2S/90% H2. The catalyst is then exposed to a sulphur-containing feed, namely benzothiophene, which acts as a model sulphur compound probe, at reaction temperature and with flowing hydrogen for approximately 4 hours. Samples are taken periodically and analysed by gas chromatography for the conversion of benzothiophene to ethylbenzene, thereby indicating the hydrodesulphurisation properties of the catalyst being tested.

The activity is calculated on both a catalyst weight and catalyst volume basis to account for any density differences between catalysts. The conditions for a typical HDS-MAT Test are as follows:

Temperature: about 560°F (about 288°C)

Pressure: about atmospheric

Feedstock: about 0.857 molar Benzothiophene in reagent grade normal heptane Space Velocity: 4 hr ¹ Catalyst Charge: 0.5 gram

The kinetics of the reactor used in the HDS-MAT Test are first order, plug flow. At the above-stated temperature and space velocity, the rate constant, k, may be expressed as:

k = In (1/1-HDS)

where HDS is the fractional hydrodesulphurisation value obtained for a given catalyst at the above stated conditions. A commercially available, state-of-the- art catalyst sold for use in hydroprocessing resid oils (Criterion HDS-1443B catalyst) was evaluated with the HDS-MAT Test under the above stated conditions and was found to have a %HDS value of 73% on a weight basis and a corresponding rate constant k value of 1.3093. The catalysts of the present invention require that the ratios of the measured HDS-MAT rate constant k, relative to that obtained with Criterion HDS-1443B, have values of 0.5-1.0, preferably 0.6-0.85, whereas catalysts of the prior art such as disclosed in United States Patent 5,047,142 are required to have values > 1.0, preferably > 1.5.

It is another feature of the catalyst composition of the present invention that the oxide of molybdenum, preferably MoO ₃ , is distributed on the above described porous alumina support in such a manner that the molybdenum gradient is about 1.0. As used in this description, the phrase "molybdenum

gradient" means the atomic ratio of molybdenum/aluminium observed on the exterior surfaces of catalyst pellets relative to the molybdenum/aluminium atomic ratio observed on surfaces of a sample of the same catalyst which has been ground to a fine powder, the atomic ratios being measured by X-Ray photoelectron spectroscopy (XPS), sometimes referred to as Electron Spectroscopy for Chemical Analysis (ESCA). The molybdenum gradient is thought to be strongly affected by the impregnation of molybdenum on the catalyst support and subsequent drying of the catalyst during its preparation. ESCA data were obtained on an ESCALAB MKII instrument available from V.G. Scientific Ltd., which uses a 1263.6 eV magnesium X-Ray source.

Generally, the finished catalysts of this invention will be characterised by the properties set forth in Table II wherein the columns show the following:

(a) The first column lists the broad ranges for the catalysts of this invention and the second column lists the preferred ranges for the catalysts of this invention, including: Total Pore Volume in cc/g; Pore Volume occupied by pores falling in designated ranges - as a volume percentage of Total Pore Volume (%TPV) or as a volume percentage of the Pore Volume in pores with diameters less than 2δθA (i.e. % of Pore Volume in the micropores) or in cc of Pore Volume per gram of catalyst; Pore Mode by volume from mercury porosimetry (dV/dD); Pore Volume falling with ± 2δA from the dV/dD peak in the less than 2δθA region; and Surface Area in m ² /g.

(b) The third column lists specific properties of the best known mode catalyst, Example I. The fourth column lists specific properties of a second sample, Example II, made by the same formula as Example I. (c) The remaining columns list properties for other hydroprocessing catalysts in the art.

The catalyst may be evaluated in a two-stage Robinson Reactor, a

Continuously Stirred Tank Reactor (CSTR) which evaluates catalyst deactivation at conditions simulating those of a two-stage H-OIL ebullated bed Unit. The feedstock is an Arabian Medium/Heavy Vacuum Resid of the type set forth above. Evaluation is carried out for four or more weeks to a catalyst age of 1.86 or more barrels per pound.

Table II Analyses of Catalyst Samples ^* ee ve t

, rom , flofPV <25θA ≥65 ≥65 72.6 65.0

U.S. Patent No./ Application Serial N Reference

PV. cc/g 55-115A, % of TPV 5,397,456 08/130,472

PV. cc/g > 100θA, % of TPV -20-25 20.9-24.4 24.4 20.9 14-22 5,435,908; 08/425,9

5,047,142

5,435,908; 08/425,9

* Values in parentheses obtained at Cytec Industries Stamford Research Laboratories. ** Contact angle *= 130"; surface tension = 484 dynes/cm. * ^♦♦ As described in U.S. Patent No. 5,047,142.

26

Preferred Embodiment

In practice of the process of this invention, the catalyst, preferably in the form of extruded cylinders of 0.039-0.044 inch (0.99-1.1 mm) diameter and about 0.15 inch (3.8 mm) length may be placed within the first- and second- stage reactors of a two-stage H-OIL Unit. The hydrocarbon charge is admitted to the lower portion of the first-stage reactor bed in the liquid phase at 650- 850°F (343-454°C), preferably 700-825 °F (371-441 °C) and 1000-3500 psia (6.9-24.2 MPa), preferably 1500-3000 psia (10.4-20.7 MPa). Hydrogen gas is admitted to the first-stage reactor of the two-stage ebullated bed H-OIL unit in an amount of 2000-10,000 SCFB (360-1800 m ³ /m ³ ), preferably 3000-8000 SCFB (540-1440 m ³ /m ³ ). The hydrocarbon charge passes through the first- stage ebullated bed reactor at a LHSV of 0.16-3.0 hr\ preferably 0.2-2 hr ^"1 . During operation, the first-stage reactor catalyst bed is expanded to form an ebullated bed with a defined upper level. Operation is essentially isothermal with a typical maximum temperature difference between the inlet and outlet of 0-50°F (0-27.8°C), preferably 0-30°F (0-16.7°C). The liquid and gaseous effluent from the first-stage reactor is admitted to the lower portion of the second-stage reactor of the two-stage H-OIL unit which is operated at essentially the same temperature and pressure as the first-stage reactor. The difference in average temperature between the first- and second-stage reactors is 0-30°F (0-16.7°C), preferably 0-15°F (0-8.3°C). Some additional hydrogen may optionally be injected to the second-stage reactor to make up for the hydrogen consumed by reactions in the first-stage reactor. During operation, the second-stage reactor catalyst bed is also expanded to form an ebullated bed with a defined upper level. Constant catalytic activity is maintained by periodic or continuous withdrawal of portions of used catalyst and replacement of the withdrawn material with fresh catalyst. Fresh catalyst is typically added at the rate of 0.05-1.0 pounds per barrel of fresh feed, preferably 0.20-0.40 pounds per barrel of fresh feed. An equal volume of used catalyst is withdrawn and discarded to maintain a constant inventory of catalyst on the volume basis. The catalyst replacement is performed such that equal amounts

of fresh catalyst are added to the first-stage reactor and the second-stage reactor of a two-stage H-OIL unit.

In a less preferred embodiment, the reaction may be carried out in one or more continuously stirred tank reactors (CSTR) which also provides essentially isothermal conditions.

During passage through the reactor, preferably containing an ebullated bed, the hydrocarbon feedstock is converted to lower boiling products by the hydrotreating/hydrocracking reaction.

Practice of the Present Invention

In a typical embodiment, employing a two-stage Robinson reactor pilot unit, a charge containing 60-95 wt %, say 88.5 wt % boiling above 1000°F (538°C) may be converted to a hydrotreated product containing only 28- 45 wt %, say 42 wt % boiling above 1000°F (538°C). The sulphur of the original charge is 3-7 wt %, typically 5.1 wt %; the sulphur content of the unconverted 1000°F+ (538°C + ) component in the product is 0.5-3.5 wt %, typically 1.6 wt %.

In another embodiment, employing a two-stage Robinson reactor pilot Unit operating at + 10°F ( + 5.6°C) over normal operating temperatures and at a larger value of catalyst age, a charge containing 60-95 wt %, say 88.5 wt % boiling above 1000°F (538°C) may be converted to a hydrotreated product containing only 24-38 wt %, say 35.4 wt % boiling above 1000°F (538°C). The sulphur content of the unconverted 1000°F + (538°C + ) component in the product is 0.5-3.5 wt %, typically 2.2 wt %. In both embodiments, the Existent IP sediment values of the 650°F +

(343°C + ) product leaving the reactor are extremely small, viz r≤O.05 wt %.

ADVANTAGES OF THE INVENTION

It will be apparent to those skilled in the art that this invention is characterised by advantages including the following:

(a) It permits attainment of increased yield of hydrocarbon products boiling below 1000°F (538°C);

(b) It permits the attainment of the above mentioned increased yield with little or no sediment as measured by the Existent IP Sediment values of the 650°F+ (343°C + ) boiling point product;

(c) It permits an improved level of sulphur removal as seen in the observed hydrodesulphurisation (HDS) of the total liquid product and the substantially improved, lower level of sulphur in the unconverted 1000°F (538°C) stream; and,

(d) It permits improved levels of carbon residue reduction and nickel and vanadium removal.

Practice of the process of this invention will be apparent to those skilled in the art from the following wherein all parts are parts by weight unless otherwise stated.

DESCRIPTION OF SPECIFIC EMBODIMENTS Best Known Mode Reactor Data

Equal amounts of Example I catalyst are placed within the reaction vessels (the first-stage and second-stage Robinson Reactors) . The hydrocarbon charge (i.e. the undiluted Arabian Medium/Heavy vacuum resid, described in Table I) is admitted in liquid phase to the first-stage Robinson reactor at 760°F (404°C) and 2250 psig (15.5 MPa). Hydrogen gas is admitted to the first- stage Robinson reactor in the amount of 7000 SCFB (1260 m ³ /m ³ ). The hydrocarbon charge passes through the first-stage Robinson reactor at a Liquid Hourly Space Velocity (LHSV) of 0.56 volumes of oil per hour per volume of liquid hold up. This is equivalent to a Catalyst Space Velocity (CSV) of 0.130 barrels of hydrocarbon charge per pound of catalyst per day. The hydrocarbon feed is uniformly contacted with hydrogen and catalyst at isothermal conditions in the first-stage Robinson reactor. The liquid and gaseous effluent from the first-stage Robinson reactor is then routed to the second-stage Robinson reactor which is operated at essentially the same temperature and pressure as the first-stage Robinson reactor. The difference in average temperature between the first- and second-stage reactors is nominally 0°F (0°C). No

additional hydrogen is injected to the second-stage Robinson reactor. The liquid effluent passes through the second-stage Robinson reactor at a Liquid Hourly Space Velocity (LHSV) of 0.56 volumes of liquid effluent per hour per volume of liquid hold up. This is equivalent to a Catalyst Space Velocity (CSV) of 0.130 barrels of liquid effluent per pound of catalyst per day. The liquid effluent from the first-stage Robinson reactor is uniformly contacted with the hydrogen-containing gaseous effluent and the second loading of catalyst at isothermal conditions in the second-stage Robinson reactor. No attempt is made to maintain constant catalytic activity by periodic or continuous withdrawal of portions of used catalyst and replacement of the withdrawn material with fresh catalyst in the two-stage Robinson reactor system. The catalyst begins as fresh catalyst and accumulates catalyst age generally expressed in barrels per pound. The average temperature is defined as the average of the temperatures of the first- and second-stage reactors. Overall, the hydrocarbon charge passes through the entire process system (i.e. the first- and second-stage Robinson reactors) at an overall LHSV of 0.28 volumes of oil per hour per volume of liquid hold up. This is equivalent to an overall CSV of 0.065 barrels of hydrocarbon charge per pound of catalyst per day. As will be discussed below, the temperatures of the first- and second-stage reactors may be raised to higher levels with the catalyst of the present invention.

Product is collected and analysed over a range of catalyst age from 0.195 to 1.08 barrels per pound (corresponding approximately to the 3rd to 16th days of the evaluation) to yield the following averaged data:

From the above Table III, it is apparent that the process of the present invention permits increasing the conversion of materials boiling above 1000°F (538°C) by 52.6 wt %; and sulphur, carbon residue, and metals are removed.

Upon distillation to recover (1 ) a first cut from the initial boiling point to 650°F (343°C), (2) a second cut from 650-1000°F (343-538°C), and (3) a third cut above 1000°F (538°C), the following is noted:

From the above Table IV, it is apparent that the sulphur content is decreased in all of the product fractions (from 5.1 wt % in the feed).

Upon distillation to recover (4) a cut which boils at temperatures of about 650°F (343°C) and higher, the following is noted:

From the above Table, it is apparent that the process of the present invention can operate at about 52.6 wt % conversion of feed components with boiling points greater than 1000°F (538°C) to products with boiling points less than 1000°F (538°C) without making any sediment.

EXAMPLE A

COMPARISON TO FIRST GENERATION CATALYST

Comparative data between the Example I catalyst of the present invention and a first generation nickel/molybdenum H-OIL catalyst (Criterion HDS-1443B), collected under virtually identical reactor conditions, are given in Table VI. The process of the present invention is superior in that it gives:

(a) No sediment versus an undesirable level with a commercially available first generation nickel/molybdenum H-OIL catalyst;

(b) An improved level of 1000°F+ to 1000°F- (538°C+ to 538°C-) wt % conversion once the data from both catalysts are kinetically adjusted to the target CSV and temperature;

(c) An improved level of sulphur removal as seen in the observed hydrodesulphurisation (HDS) of the total liquid product and the substantially improved, lower level of sulphur in the unconverted 1000°F (538°C) stream; and,

(d) Improved levels of carbon residue reduction and nickel and vanadium removal.

* Criterion HDS-1443B H-OIL catalyst.

* * 1st order CSTR kinetics (assuming equal rate constants for the 1 st- and

2nd- stage reactors); Activation Energy = 65 kcal/mole.

EXAMPLE B

DATA AT HIGHER TEMPERATURES

In the evaluation of the Example I catalyst of the present invention, reactor temperatures were raised about 10°F (5.6°C) over a period of 2.5 days to a final temperature of approximately 770°F (410°C) (i.e. the first-stage, second-stage, and average temperatures). Product was collected and analysed over a range of catalyst age from 1.28 to 1.86 barrels per pound (corresponding approximately to the 19th to 28th days of the evaluation). Comparative data between the catalyst of the present invention operating at about + 10°F ( + 5.6°C) compared to the first generation nickel/molybdenum H-OIL catalyst (Criterion HDS-1443B) at the same catalyst ages are given in Table VII. The process of the present invention is superior in that it gives:

(a) Low sediment at 60 wt % 1000°F+ to 1000°F- (538°C + to 538°C-) conversion versus an undesirable level with the first generation nickel/molybdenum H-OIL catalyst operating at only 52 wt % 1000°F+ to 1000°F- (538°C+ to 538°C-) conversion;

(b) An improved level of 1000°F+ to 1000°F- (538°C+ to 538°C-) wt % conversion by the observed data and once the data from both catalysts are kinetically adjusted to the target CSV; (c) An improved level of sulphur removal as seen in the observed hydrodesulphurisation (HDS) of the total liquid product and the substantially improved, lower level of sulphur in the unconverted 1000°F+ (538°C + ) stream; and (d) Improved levels of carbon residue reduction and nickel and vanadium removal.

It was noted that the sulphur levels of the 650°F+ to 1000°F +

(343°C+ to 538°C + ) bp boiling cut (approximating the composition of a vacuum gas oil) were slightly higher with the Example I catalyst of the present invention operating at about + 10°F (5.6°C) compared to the level obtained

with the first generation catalyst when both were at catalyst ages of 1.28 through 1.86 barrels per pound.

The catalyst of the present invention, besides giving low sediment results for the 650°F+ (343°C + ) boiling cut, also showed improved operability. The evaluation went smoothly at both 760°F (404°C) and 770°F (410°C). On the other hand, the first generation catalyst evaluation showed evidence of plugging due to accumulated sediment during the course of the run. Operations became somewhat erratic with the first generation catalyst at about 1.54 bbl/pound catalyst age and the unit had to be shut down and partially cleaned out before the evaluation of the first generation catalyst could be completed. With so much trouble due to sediment, it was felt that temperatures could not be raised any higher with the first generation catalyst.

* Criterion HDS-1443B H-OIL catalyst.

* * 1st order CSTR kinetics (assuming equal rate constants for the 1 st- and nd- stage reactors); Activation Energy = 65 kcal/mole.