MXPA98002265A - Hydroconversion process that uses a nimo catalyst charged with phosphorus with distribution specific poro size - Google Patents

Abstract

A process for hydrotreating a hydrocarbon feedstock containing a boiling point higher than 538øC (1000øF), sulfur, metals and coal residues in products containing decreased levels of components having a boiling point higher than 538øC (1000øF), of sulfur decreased, particularly sulfur content decreased in products higher than unconverted 538øC (1000øF) and reduced sediment, which comprises: contacting said hydrocarbon feed with hydrogen under isothermal conditions of hydroprocessing in the presence of, as catalyst, a porous support of alumina containing < = 2.5% by weight of silica and comprising 2.2 to 6% by weight of Group VIII metal oxide, 7 to 24% by weight of Group VIB metal oxide and 0.3 to 2.0% by weight of phosphorus oxide, said catalyst having a Total Surface Area of 175 to 205 m / g, a Total Pore Volume of 0.82 to 98 cc / g, and a Pore Size Distribution where 29.6.0 to 33.0% of the Total Pore Volume is present as macropores of diameter larger than 250Å, 67.0 to 70.4% of the Total Pore Volume is present as micropores smaller than 250Å, < = 65% of the micropore volume is present as micropores of diameter at 250 μA of a pore fashion by volume of 110 to 130Å, less than 0.05 cc / g of pore volume is present in micropores with diameters less than 8

Description

HYDRVERSION PROCESS USING A CATALYST NiMoP CHARGED WITH PHOSPHORUS WITH PORO SIZE DISTRIBUTION SPECIFIED This invention relates to a process for hydrotreating a hydrocarbon feed. More particularly, this relates to a hydro-conversion process employing a catalyst with a specified pore size distribution that achieves improved levels of hydrversion of feed stock components having a boiling point higher than 538 ° C (1000 °). F) to products having a boiling point of less than 538 ° C (1000 ° F), improved hydrodesulphurisation, particularly sulfur removal of 538 ° C (1000 ° F) products not confounded, and reduced production of sediments that make possible operation at higher temperatures. As is well known to those skilled in the area, it is desirable to convert heavy hydrocarbons, such as those having a boiling point above about 538 ° C (1000 ° F), into light hydrocarbons which are characterized by higher value economic. It is desirable to treat the hydrocarbon feedstock, particularly petroleum residues, to achieve other objectives including hydrodesulfurization (HDS), carbon residue reduction (CRR), and hydrodemetalization (HDM) -the most current particularly includes removal of nickel components (HDNi) and vanadium compounds (HDV). Generally these processes employ hydrotreating catalysts with specific pore ranges with relatively small diameters (for example micropores, here defined as pores having diameters less than 250 A) and pores having relatively large diameters (for example macropores, here defined as pores having diameters larger than 250 A). One way to develop improved catalysts for the improvement of petroleum residues has included increasing the micropore diameters of essentially monomodal catalysts (not having significant macro-pores) to overcome the diffusion limitations. The first hydrotreating catalysts for petroleum distillates were generally monomodal catalysts with micropores of very small diameters (less than said 100 A) and distribution of coarse pore sizes. First-generation petroleum residue hydrotreating catalysts were developed by introducing a large amount of macroporosities into a pore structure of distillate hydrotreating catalysts to overcome the diffusion resistance of large molecules. Such catalysts, which are fully considered HDS / HDM bimodal catalysts, are typified by US Patent Nos. 4,746,419, 4,495,328, 4,395,329, and 4,089,774, discussed below. U.S. Patent 4,746,419 (PECK et al) discloses an improved hydrversion process for the hydro-conversion of heavy hydrocarbon feedstock containing asphaltenes, metals, and sulfur compounds, minimizing said process the production of soluble carbonaceous solids and levels of abrasion of the catalyst. The disclosed process employs a catalyst having 0.1 to 0.3 cc / g of its pore volume in pores with diameters greater than twelve hundred A and not more than 0.1 cc / g of its pore volume in pores having diameters greater than four. thousand Á. The present invention is distinguished from this reference because the prior art discloses only useful pore size distribution characteristics to minimize the production of insoluble carbonaceous solids and does not disclose a pore size distribution that could provide additional hydrodesulfurization activities. On the contrary, the catalysts of the present invention require a single pore size distribution in order to provide additional hydrodesulphurisation to the feed stock components having a boiling point higher than 538 ° C (1000 ° F) to products having a Boiling point less than 538 ° C (1000 ° F), and improved hydrodesulfurization. The present invention provides improved levels of hydroconversion of feed stock components having a boiling point higher than 538 ° C (1000 ° F) to products having a boiling point of less than 538 ° C (1000 ° F) improved hydrodesulfurization, particularly improved sulfur removal of unconverted 538 ° C (1000 ° F) products, and reduced production of sediments that make it possible to operate at higher temperatures compared to operations with a commercial vacuum hydroconversion catalyst, having a distribution of pore size that meets the requirements of this reference. U.S. Patent 4,395,328 (Hensley, Jr. Et al). It discloses a process for the hydroconversion of a hydrocarbon stream containing asphaltenes and a substantial amount of metals, comprising contacting the flow (in the presence of hydrogen) with a catalyst present in one or more fixed or bubbled beds the catalyst comprising at least one metal which can be a metal of the group VTB or group VIII, a phosphorus oxide and where the alumina support material has at least 0.8 cc / g per gram of TPV is in pores having diameters of 0-1200A, at least 0.1 cc / g of TPV is in pores having a diameter of 1200-50000Á, a surface area in the range of 140-190 mg, and the support material was formed as a compound com-pending alumina and one or more phosphorus oxides in a structured material and was thus heated with steam to increase the average pore diameter of the catalyst support material prior to impregnation with the active metals. The present invention differs from this reference in that the support of the present invention does not contain one or more phosphorus oxide, is not heated with steam to increase the average pore diameter, and requires a surface area larger than about 205%. 275 M2 / g and there is a much more precise definition of pore volume distribution. U.S. Patent 4,395,329 (Le Page et al) discloses a process of hydro-refining a feed stock containing enough metal using a catalyst containing alumina, a group VI metal and a metal of the iron group, the catalyst having an area of total surface of 120-200 m / g, a total pore volume of 0.8 to 1.2 cc / g, and a distribution of pore diameters where 0 to 10 percent of the Total Pore Volume is present as micropores with smaller diameters that 100 Á, 35 to 60 percent of the total pore volume is in pores with diameters of 100 to 600 Á, and 35 to 55 percent of the total pore volume is present as macro-pores of diameters larger than 600 Á. The present invention is distinguished from that difference because the catalyst of the state of the art requires 35 to 55 percent of TPV in pores with a diameter greater than 600 Á and the catalysts of the present invention have only 21 to 27 percent of TPVs in pores larger than 600 Á. U.S. Patent 4,089,774 (Oleck et al.) Discloses a process for the demetallisation and desulphurization of a hydrocarbon oil comprising contacting the oil with hydrogen and a catalyst, the catalyst comprising a Group VIB metal and a metal of the iron group (e.g. iron, cobalt, or nickel) in a porous support, and having a surface area of 125-210 pr / g and a TPV of 0.4-0.65 cc / g with at least 10% of TPV in pores having diameters less than 30A, less 50% of pore volume accessible to mercury being of pores having diameters of 30 to 150A, and at least 16.6% of pores accessible to mercury being of pores having diameters larger than 300Á. The present invention is distinguished from this reference in that the catalyst of the state of the art requires a relatively low Total Pore Volume of only 0.4-0.65 cc / g. whereas the catalyst of the present invention requires a much higher Total Pore Volume of 0.82-0.98 cc / g. The American patent 5, 221,656, issued to Clark et al. discloses a hydroprocessing catalyst comprising at least one hydrogenation metal selected from the group consisting of group VIB metals and consisting of Group VIII metals deposited on an inorganic oxide support, said catalyst being characterized by a larger surface area than about 220 m2 / g, a pore volume of 0.23 to 0.31 cc / g in pores with a radius larger than about 600A (for example in pores with a diameter greater than 1200A), an average of radio pores around 30 to 70A in pores with a radius of less than about 600A (for example, average pore diameters of about 60 to 140A in pores with a diameter less than about 1200A), and an increase curve of maximum pore volumes of radio from around 20 to 50A (for example around 40 to 100A diameter). In the present invention, pores having a diameter larger than 1200 Á are only about 0.15 to 0.20 cc / g and the pore volume increase curve has a maximum (e.g., Pore Fashion) at 110 to 130Á. Also, reflecting the longer range of dimensions of Pore Modes, the present catalyst has much lower surface area of 175 to 205 nr7g. A recent way of approaching the development of improved catalysts to process petroleum residues has meant the use of catalysts having intermediate micropore diameters between the above described HDS and HDM monomodal catalysts, as well as with sufficient macroporosities to overcome the limitations of diffusion for low HDS oils (eg sulfur removal of hydrocarbon products from a hydrotreated petroleum residue having a boiling point above 538 ° C (1000 ° F)) but limited micro-porosities to limit poisoning of the interior of the particles catalytic Catalysts with micropores of intermediate diameters between the above-described HDS and HDM monomodal catalysts with limited micro-porosities include those of the American patents 4,941,964, 5,047,142 and 5,399,259 and comprising those of the US patent application serial number 08 / 425,971 which is divisional of the now US Patent 5,435,908 discussed above. US Patent 4,941,964 discloses a process for hydrotreating a sulfur and metal containing feed comprising said process contacting said feed with hydrogen and a catalyst such that the catalyst is maintained under isothermal conditions and is exposed to a uniform feed quality. , the catalyst comprising an oxide of a Group VIII metal, an oxide of a Group VIB metal and 0 to 2.0% by weight of a phosphorus oxide in a porous alumina support, having a surface area of 150 to 210 pr. / g and a Total Pore Volume (TPV) of 0.50 to 0.75 cc / g in such a way that 70 to 85% of the TPV is of pores having a diameter of 100 to 160A and 5.5 to 22.0% of the TPV is of pores having a diameter greater than 250 Á. US Patent 5,047,142 discloses a catalyst composition useful in the hydroprocessing of a metal-sulfur-containing feed batch, comprising a nickel or cobalt oxide and a molybdenum oxide in a porous support of alumina such that the gradient of molybdenum of the catalyst has a value of less than 6.0 and 15-30% of the nickel or cobalt is in the form of an acidic extract, having a surface area of 150 to 210 m2 / g, a total of Total Volume of Pores (TPV) ) from 0.50 to 0.75 cc / g, and a pore size distribution such that less than 25% of TPV is of pores having diameters less than 100A, 70.0% to 85.0% of the TPV is of pores having diameters of 100 to 160A and 1.0 to 15.0% of the POS is of pores having diameters greater than 250A. US Patent 5,399,259 discloses a process for hydrotreating a feed containing asphaltenes, metals and sulfides comprising contacting said feed with hydrogen and a catalyst such that the catalyst is maintained under isothermal conditions and is exposed to a uniform quality of feeding, the catalyst comprises 3 to 6% by weight of a metal oxide of group VIII, 14.5 to 24% by weight of an oxide of a metal of Group VIB and 0 to 6% of a phosphorus oxide in a porous support of alumina , having a surface area of 165 to 230 m2 / g and a Total Pore Volume (TPV) of 0.5 to 0.8 cc / g in such a way that 5% of TPV is of pores with diameters less than about 80A, at least 65 % of the pore volume in pores with diameters less than 250A is pores with diameters of ± 20A of a Pore Fashion of about 100 to 135A and 22-29% of the TPV is of pores having a diameter greater than 250A. The present invention is distinguished from this reference in that the state of the art requires a relatively low Total Pore Volume of only 0.5 to 0.8 cc / g, and a relatively low microporosity of 22-29% TPV in pores having larger diameters than 250Á. On the contrary, the catalyst of the present invention requires a larger Total Pore Volume of 0.82 to 0.98 cc / g and a much higher porosity level of 29.6 to 33.0% of the TPV in pores having diameters larger than 250A. In the patent application of the United States with serial number 08 / 425,971 there is disclosed a hydrotreatment process using, as a catalyst, a porous support of alumina with pellet diameters of 0.032 to 0.038 inches (0.81 to 0.96 mm) comprising 2.5 a 6% by weight of a non-noble metal oxide of group VIII, 13 to 24% by weight of a metal oxide of group VIB, less than or equal to 2.5% by weight of silicon oxide, typically around 1.9 at 2% of silicon oxide added intentionally, and 0 to 2% of a phosphorus oxide, preferably less than 0.2% by weight of a phosphorus oxide, with no phosphorus-containing component intentionally added during catalyst preparation, said catalyst having a total Surface Area of 165 to 210 m2 / g, a Total Volume of Pores of 0.75 to 0.95 cc / g, and a Pore Diameter Distribution by means of which 14 to 22% of the Total Pore Volume present as diameter macropores >; 1000Á, 22 to 32% of the Total Pore Volume is present in diameters > 250Á, 68 to 78% of the Total Pore Volume is present as diameter pores < 250Á, 26 to 35% of the Total Pore Volume is present as mesopores of diameters > 200Á, 34 to 69% of the Total Pore Volume is present as secondary micropores of diameters between 100 and 200A, from 5 to 18% of the Total Pore Volume is present as primary micropores of diameter < 100 Á, and > 57% of the micro-pore volume is present as micropore of diameter between ± 20A of a pore fashion by volume of 100 to 145A. On the contrary, the present invention employs, as a catalyst, a porous support of nominally pure alumina with diameters of 0.81 to 1.12 mm (0.032 to 0.044 inches), preferably 0.99 to 1.2 mm (0.039 to 0.044 inches) peelings. 2.2 to 6% by weight of a non-noble metal oxide of Group VIII, 7 to 24% by weight of a metal oxide of group VIB, less than or equal to 2.5% by weight of silicon oxide (for example silica: Si02), preferably 1.3 to 2.5% by weight of intentionally added silica during the pre-setting of the catalyst, and 0.3 to 2% by weight of a phosphorus oxide, preferably 0.5 to 1.5% by weight of a phosphorus oxide, with component containing phosphorus intentionally added the catalyst preparation, said catalyst having a Total Surface Area of 175 to 205 mVg, a Total Pore Volume of 0.82 to 0.98 cc / g, and a Pore Diameter Distribution by means of which from 29.6 to 33.0% of the Total Pore Volume is pr As macropores with a diameter greater than 250A, 67.0 to 70.4% of the Total Pore Volume is present as micropores smaller than 250A, < 65% of the volume of micropores is present as micropores of diameters of ± 25A of a pore fashion by volume of 110 to 130A, and less than or equal to 0.05 cc / g of volume of micropores is present in micropores with diameters less than 80A . Another recent line of development of catalysts, has been to develop a catalytic hydroconversion of feed stock having a boiling point higher than 538 ° C (1000 ° F) has contemplated the use of catalysts having intermediate micropores between the above-described monomodal catalyst HDS and HDM with pore volumes in the range HDS type and with sufficient macroporosities to overcome the diffusion limitations for feed stock conversion having a boiling point higher than 538 ° C (1000 ° F) in products having a point of boiling less than 538 ° C (1000 ° F), but with limited macroporosities to limit the poisoning of the interiors of catalyst particles. Such catalysts are described in U.S. Patent 5,397,456 and the copending U.S. patent application serial number 08 / 130,472 discussed below. US Pat. No. 5,397,456 discloses a catalyst composition useful in the hydroconversion of a metal-sulfur-containing feedstock comprising said metal oxide of group VIII and a metal oxide of group VIB and optionally phosphorus in a porous support of alumina, the catalyst having a Total Surface Area of 240 to 310 nr / g, a Total Pore Volume of 0.5 to 0.75 cc / g, and a Pore Diameter Distribution by means of which 63 to 78% of the Po-Ro Volume Total is present as micropores with diameters between 55 and 115A and 11 to 18% of the total volume of pores is present as macropores with diameters larger than 250A. The heavy feed stock is brought into contact with hydrogen and the catalyst. The catalyst is maintained in isothermal conditions and is exposed to a uniform feed stock quality. The process is particularly effective to achieve desired levels of hydroconversion of feed stock having a boiling point higher than 538 ° C (1000 ° F) in products having a boiling point of less than 538 ° C (1000 ° F). The present invention is distinguished from this reference because the state of the art requires a catalyst with a Pore Diameter Distribution of 63 to 78% of the Total Pore Volume is present as micropores of diameters between 55 and 115A and 11 to 18% of the Total Pore Volume is present as macropores with diameters larger than 250A, considering that the catalyst employed in the present invention has only 20 to 25% of the Total Volume of Pores present as of diameters between 55 and 115A and 29.6 to 33.0% of the Total Volume of Pores is present as macropores of larger diameters than 250Á. In the American application with serial number 08 / 130,472 is disclosed a hydrotreating process and catalyst where 50 to 62.8% of the TPV is present as of diameters between 55 and 115A and 20 to 30.5% of the Total Volume of Pores is present as macropores of diameters larger than 250Á. In this case, the catalyst preferably has only about 20 to 25% of the TPV present in pores having diameters of between 55 to 115A. None of the types of catalysts identified above in the state of the art have been found to be effective in achieving all the desired process improvement needs. The first catalysts in the state of the art address the need to improve hydrodesulfurization and / or demetallization when measured in the final liquid product. A recent line of development of catalysts has been to develop improved catalysts for low HDS oils (for example, removal of sulfur compounds selectively from unconverted hydrocarbon products having a boiling point higher than 538 ° C (1000 ° C). F) of a hydroprocess operating with significant hydroconversion of the components of the feed stock (eg petroleum residues) having a boiling point higher than 538 ° C (1000 ° F) for products having a boiling point lower than 538 ° C (1000 ° F)). More recent developments of low oil HDS catalysts have been to develop HDS catalysts of low oils with some degree of sedimentation control allowing the use of some higher temperatures and reduced sediment creation. However, none of the low oil HDS catalysts described above provides improved levels of hydroconversion of feed stock components having a boiling point higher than 538 ° C (1000 ° F) for products having a lower boiling point than 538 ° C (1000 ° F), while, at the same time, reduce sediment production. An additional line of catalyst development has been to develop improved catalysts to convert components of feed stocks having a boiling point higher than 538 ° C (1000 ° F) to products having a boiling point lower than 538 ° C (1000 ° F). The most recent development has been to provide hydroconversion catalysts with some degree of sedimentation control allowing the use of some higher temperatures and reduction of sediment production. Although the hydroconversion catalysts described above provide hydroconversion levels of feed stock components having an ebu-Ilition point higher than 538 ° C (1000 ° F) to products having a boiling point lower than 538 ° C (1000 ° C) F), these do not provide the desired level of sulfur removal obtained from the low oil HDS catalysts described above and these hydroconversion catalysts still produce some amount of sediment. It would be desirable to have available a catalyst that would provide improved hydrodesulfurization, particularly improved base HDS, and would not produce any sediment and which could also resist operations at very high temperatures, thus it may be possible to reach a much higher level. high hydroconversion without the undesirable formation of sediments. Undesirable low levels of hydroconversion represent a problem that is particularly acute for those refineries that operate units in the vacuum of hydroprocessing waste at or near its sedimentation limit. Such refineries, in the absence of sediments, could increase the temperature and thus the hydroconversion levels. It is an object of the present invention to provide a process for hydroconverting a hydrocarbon feedstock, particularly, for hydroconverting feed stock components that have a boiling point higher than 538 ° C (1000 ° F) in products having a boiling point of less than 538 ° C (1000 ° F) while simultaneously stirring high amounts of Sulfur from the unconverted product flow 538 ° C + (1000 ° F +). It is also an object of the present invention to provide low Existing and Accelerated IP Sediment values in the boiling point products at 343 ° C + (650 ° F +) (discussed below under "Sediment Measurement"). It is also an object of the present invention to allow the use of much higher operating temperatures and minimum sediment. According to certain aspects, this invention is directed to a process for hydroprocessing a hydrocarbon feedstock containing components with a boiling temperature around 538 ° C (1000 ° F), and sulfur, metals, and residues. of carbon comprising: contacting said hydrocarbon feedstock with hydrogen under conditions of isothermal hydroprocessing in the presence of, as a catalyst, a porous support of alumina containing <2.5% by weight of silica or comprising 2.2 to 6% by weight of a non-noble metal oxide of group VIII, 7 to 24 by weight of a metal oxide of Group VIB, and 0.3 to 2% of a phosphorus oxide , said catalyst having a Total Surface Area of 175 to 205 m2 / g, a Total Pore Volume of 0.82 to 0.98 cc / g, and a Pore Diameter Distribution by means of which from 29.6 to 33.0% of the Pore Volume Total is present as macropores with a diameter greater than 250A, 67.0 to 70.4% of the Total Pore Volume is present as micropores smaller than 250A, < 65% of the volume of micropores is present as micropores of diameters of ± 25A of a pore fashion by volume of 110 to 130A, and less than or equal to 0.05 cc / g of volume of micropores is present in micropores with smaller diameters that 80A, to thereby form a hydroprocessed product containing lower content of components boiling about 538 ° C (1000 ° F) and residues of sulfur, metals, and carbon; and recovering said hydroprocessed products containing reduced contents of components boiling above 538 ° C (1000 ° F), carbon residues, metals and sulfur, recovering said hydroprocessed product containing reduced sulfur content in the portion of the hydroprocessed boiling products about 538 ° C (1000 ° F), and recovering said hydroprocessed product containing reduced sediment content in the portion of the hydroprocessed product that boils around 343 ° C (650 ° F). The catalyst of the present invention makes the operation possible at around + 5.6 ° C (+ 10 ° F) and around a conversion of + 8% by weight of 538 ° C (1000 ° F) compared to the operations with the first generations of H-OIL catalysts. This constitutes a substantial economic advantage. DESCRIPTION OF THE INVENTION Stock of feed The hydrocarbon feed that can be charged to the process of this invention can include high boiling, heavy oil cuts typified by crude oil gas, crude oil gas vacuum, petroleum coke, waste of crude oil, vacuum residues, etc. The process of the present invention is particularly useful for treating high boiling oils containing components that boil above 538 ° C (1000 ° F) to convert them to products boiling below 538 ° C (1000 ° F). The filler may be an oil fraction having an initial boiling point of about 343 ° C (650 ° F) characterized by the presence of a high undesirable content of components boiling above 538 ° C (1000 ° F), and sulfur, carbon and metal waste; and such charges may be subject to hydrodesulfurization (HDS). In particular, the load may be non-diluted vacuum residue. A typical load that can be used is a Medium / Heavy Arabian Vacuum Residue having the properties shown in Table I below: It is a particular feature of the process of this invention that it can allow the treatment of a hydrocarbon charge, particularly those containing components that boil above about 538 ° C (1000 ° F), to form products that are characterized by an increase in content of components boiling above 538 ° C (1000 ° F) and by decreasing the content of undesirable components typified by sulfur, metals, and carbon residues. It is another feature of the process of the present invention that this provides improved removal of sulfur from unconverted 538 ° C (1000 ° F) products. It is another feature of the process of the present invention that it provides the above-mentioned improvements with no or very little sediment formation as measured by the Accelerated and Existing IP Sediment values of the products boiling at 343 ° C + ( 650 ° F +). It is another characteristic of the process of the present invention to allow operations at high temperatures with the consequent high levels of 538 ° C + at 538 ° C- (1000 ° F + at 1000 ° F-) which can be achieved with the use of first class catalysts. generation. Sediment Measurement It is a particular feature of the catalysts of this invention to allow the operation to be carried out under conditions that yield a substantial decrease in sediment content in the product flow after treatment. The load of a hydroconversion process is typically characterized by a very low content of sediment of 0.01 percent maximum by weight (% by weight). The sediment is typically measured by a sample test by the standard "Shell Hot Filtration Solid Test" (SHFST). See Jour. Inst. Pet. (1951) 37 pages 596-604 Van Kerknoort et al. The typical hydroprocessing process in the state of the art commonly yields the Shell Hot Filtration Solids above about 0.17% by weight and as high as about 1% by weight in the products 343 ° C + (650 ° F +) recovered in the lower flash drum (BFD). The production of significant amounts of sediment is undesirable from the point of view that it results in depositions in units downstream which for their proper functioning must be removed. This requires of course that the unit be stopped for an undesirably long period of time. Sediment is also undesirable in the products because it is deposited in and within several pieces of the hydroprocessing equipment units downstream and interferes with the proper operation of, for example, heat exchangers, fractionating towers, etc. Very high levels of sediment formation (for example 1% by weight), are not frequent by those refiners that operate units of hydroprocessing of waste operating at vacuum in moderate levels of conversion of components of stock of feeding having boiling points larger than 538 ° C (1000 ° F) in products having boiling points less than 538 ° C (1000 ° F) (say, conversion of 40-65 percent by volume -% by volume-) and relatively low but still undesirable values of sediments (for example a sediment limit of 0.17% by weight). In the present invention, the IP 375/86 test method for the determination of total sediments has been very useful. The test method is described in the designation D 4870-92 ASTM. The IP method 375/86 was designed for the determination of total sediments in residual fuels and is very appropriate for the determination of total sediments in the product boiling point of 343 ° C + (650 ° C +). The boiling product of 343 ° C + (650 ° C +) can be tested directly for total sediment which is designated as the "Existing IP Sediment Value". It has been found that the Existing IP Sediment Test provides essentially equivalent test results to the Shell Hot Filtration Solids Test described above. Nevertheless, it has been noted that even the products of boiling point at 343 ° C + (650 ° C +) that give low values of Existing IP Sediment, can produce additional sediments after storage. Thus, a more severe test for sediments has been developed. In this modified test, heat 50 grams of product from a boiling point to 343 ° C + (650 ° C +) are heated to about 90 ° C and mixed with about 5 cm3 of hexa-decane reagent grade. The mixture is set at about 100 ° C for about an hour. The resulting sediment is then measured by the test method IP 375/86. The values obtained from this modified test are designated the "Accelerated IP Sediment values". As it is recommended that the IP / 86 test method be restricted to samples containing less than or equal to about 0.4 to 0.5 weight percent sediment, the sample size is reduced when high sediment values are observed. This leads to an important reduction in the values for those samples with very high sediment content. It will be noted that the catalysts of this invention, characterized by i) around 0.15 to 0.20 cc / g of pores in the range >; 1200 A, ii) around 21-27% of TPV in pores in the range of > 600 Á, iii) 29.6 to 33.0% of the TPV in pores having a diameter > 250 Á, iv) 67.0 to 70.4% of the TPV in micropores with a diameter smaller than 250 Á, v) > 65% of the volume of micropores is present as micropores of diameter within ± 25 Á of a pore mode by volume of 110 to 130 Á, vi) of 20 to 25% of the volume of microporous pores with diameter 55 to 115Á and vii ) less than 0.05 cc / g micropore volume in micropores with diameters less than 80 A, -they are particularly advantageous in that they allow to reach the flow of hydrocarbon product containing the lowest sediment content in the highest conversion, while producing products characterized by sulfur, coal waste and low metal content. It is a characteristic of the catalysts of this invention that with the hydrotreatment products with sediment < 0.15% by weight, when measured by the IP Sediment Existing test in the portion of boiled hydroprocessed products above 343 ° C (650 ° F), typically as low as 0.0-0.1% by weight, preferably 0.0 to 0.04% by weight, that is, 0.05% in weight. Reaction Conditions In the practice of the process of this invention (as typically conducted in a one-stage Robinson reactor in pilot plant operations), the hydrocarbon feedstock is brought into contact with hydrogen under isothermal hydrotreating conditions in presence of the catalyst. The operating pressure can be from 10.4 to 69 MPa (1,500 to 10,000 PSIG), preferably from 12.4 to 17.3 MPa (1,800 to 2,500 PSIG), say 15.5 MPa (2,250-PSIG). Hydrogen is loaded into the Robinson reactor at a rate of 360 to 1,800 MJ / M3 (2,000 to 10,000 SCFB), preferably 540 to 3,240 M3 / M3 (3,000 to 8,000 SCFB), say 1,260 M3 / M3 (7,000 SCFB). The Speed Space Time of the liquid (LHSV) is generally from 0.1 to 1.5, say 0.56 volumes of oil per hour per volume of liquid contained in the reactor. The operating temperature is typically 371 to 482 degrees C (700 to 900 degrees F), preferably to 399-468 degrees C (750 to 875 degrees F), say (404 degrees C). The operation is essentially isothermal. The temperature may vary generally through the bed less than about 11.1 degrees C (20-degrees F). In another more preferred embodiment of the process of the present invention, the liquid and gaseous tributaries of the Robinson reactor of a previously described step is directed to a second stage Robinson reactor containing the same weight of catalyst that has been loaded into the Robinson reactor. of the first step and which is operated at essentially the same temperature and pressure of the Robinson reactor of the first step. The average temperature difference between the reactors of the first and second stage is 0 to 16.7 degrees C (0 to 30 degrees F), preferably between 0 and 8.3 degrees C (0 to 15 degrees F), say 0 degrees C (0 degrees F). No additional amount of hydrogen is normally injected into the Robinson reactor of the second stage. The liquid effluent passes through the Robinson reactor of the second stage to a similar LHSV to that of the Robinson reactor of the first stage. The liquid effluent from the Robinson reactor of the first stage is uniformly brought into contact with the gaseous effluent containing hydrogen and the second catalyst charge under isothermal conditions in the Robinson reactor of the second stage. No attempt is made to keep the catalytic activity constant by continuously or periodically removing the portions of used catalysts and replacing the removed material with fresh catalyst in the Robinson two-stage reactor system. The catalyst starts as a fresh catalyst and accumulates catalyst age in barrels per pound. The average temperature is defined as the average of the temperatures of the first and second stage reactors. The average operating temperature is typically 371 to 482 degrees C (700 to 900 degrees F), preferably 399 to 468 degrees C (750 to 875 degrees F), say 404 degrees C (760 degrees F). Above all, the hydrocarbon charge passes through the entire processing system (for example the Robinson reactors of first and second stage) to an effective LHSV of 0.05 to 0.75, say 0.28 volumes of oil per hour per volume of liquid contained in the reactor. In general, the reaction can be carried out in one or more crops of the reactor tanks (CSTRs), such as a Robinson reactor, in which the catalyst is exposed to a uniform feed quality. In a particularly preferred embodiment of the present invention, a feed stock of hydrocarbon containing metal and sulfur is catalytically processed using the process configuration (TM) H-OIL. H-OIL owns the bubbling bed process (co-owned with Hidrocarbon Research, Inc. and Texaco Development Corporation) for the catalytic hydrogenation of residual and heavy oils to produce high-grade distilled petroleum products and an unconverted base product suitable for mixing with a low sulfur fuel oil. The bubble bed system operates under essentially isothermal conditions and makes it possible to expose catalyst particles to a uniform feed quality. In the H-OIL process, a catalyst is contacted with hydrogen and a hydrocarbon feed stock containing metal and sulfur by means of which it is ensured that the catalyst is maintained under essentially isothermal conditions and exposed to a uniform feed quality. Preferred means for achieving such contact includes contacting the hydrogen feed with catalyst in a simple bubbled bed reactor, or in a series of two to five bubbled bed reactors, being particularly preferred with a series of two ebullated fluidized bed reactors. This hydroprocessing process is particularly effective in achieving high hydrodesulfurization levels with residual vacuum feed stocks. In the H-OIL process the hydrocarbon charge is admitted to the first stage reactor of a two-stage bubbled H-OIL unit in the liquid phase at 343-454 degrees C (650-850 degrees F), preferably 371 to 441 degrees C (700 to 825 degrees F) and 6.9 to 24.2 MPa (1000 to 3500 PSIA), preferably 10.4 to 20.7 MPa (1500 to 3000 PSIA). The hydrogen gas is admitted to the first stage reactor of a two-stage bubbled H-OIL unit in amounts of 360 to 1800 M3 / M3 (2000 to 10000 SCFB), preferably 540 to 1440 M3 / M3 (3000 to 8000 SCFB). The hydrocarbon charge passes through the bubbling bed reactor of the first stage at an LHSV of 0.16-3.0 Hr "", preferably 0.2 to 2 Hr "1. During the operation, the catalyst bed is expanded to form a bed bur - Buckling with a defined upper level The operation is essentially isothermal with a typical maximum temperature difference between inlet and outlet of 0 to 27.8 degrees C (0 to 50 degrees F), preferably 0 to 16.7 degrees C (0 at 30 degrees F.) The gaseous and liquid effluents from the effluent from the first stage are 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 reactor from First stage The difference in the average temperature between the first and second stage reactors is 0 to 16.7 degrees C (0 to 30 degrees F), preferably Oa 8.3 degrees C (0 to 15 degrees F). of additional hydrogen can optionally be injected The reactor of the second stage is used to restore the hydrogen consumed in the reactions in the first stage reactor. In the H-OIL process, the constant catalytic activity by periodic or continuous withdrawals of portions of used catalyst and replacement of the removed material with fresh catalyst. The fresh catalyst is added generally in a proportion of 0.05 to 1.0 pounds per barrel of fresh feed, preferably 0.20 to 0.40 pounds per barrel of fresh feed. An equal volume of used catalyst is removed and discarded to maintain a constant inventory of catalysts in the bulk volume. The catalyst replacement is carried out in such a way that equal amount of fresh catalyst is 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. However, the alumina can be alpha, beta, teta, gamma, or alumina range, with the alumina range being preferred. The alumina charge that may be in carrying out this invention may be commercially available from the suppliers or may be prepared by various standardized processes in which 85-90 parts of pseudoboemite alumina is mixed with 10 to 15 parts of recycled fines. The silica (SiO2) can be incorporated in small amounts typically below about 2.5% on the base of the finished catalyst, and preferably 1.3 to 2.5% by weight on the base of the finished catalyst. Acid is added and the mixture is fluffed and then struded in an auger type struder through a die having cylindrical holes dimensioned to produce a calcined substrate having a diameter of 0.81 to 1.12 mm. (0.032 to 0.044 inches), preferably 0.99 to 1.12 mm. (0.039 to 0.44 inches). The extrudate is air dried at a final temperature of generally 121 to 135 degrees C (250 to 275 degrees F) resulting in strings with 20 to 25% calcined solids. The air-dried extrudate is then calcined by air in an indirect fire oven for 0.5 to 4 hours in an atmosphere of air and steam at generally about 538 to 621 degrees C (1000 to 1150 degrees F).

Catalysts of the Present Invention - Pore Size Distribution The catalyst that can be employed is characterized by a Total Surface Area (TSA), Total Pore Volume (VPT), and Pore Diameter Distribution (Pore Dimension Distribution PSD ). The Total Surface area is 175-205 m2 / g, preferably 175-205 m2 / g, that is 178 m2 / g. The Total Pore Volume (TPV) can be 0.82-0.98, preferably 0.82-0.90, that is 0.83 cc / g. Less than 0.05 cc / g micropore volume is present in micropores with diameters less than 80A. Micropores of diameter in the range of less than 250A are present in an amount of about 67.0-70.4% of the Total Pore Volume, preferably 67.0-69.1% TPV, ie 67.0% of the TPV. Preferably > 65% of the volume of micropores is present as micropores with a diameter between ± 25A of a pore fashion by volume (for example V / MAX) of 110-130Á. The amount of Total Pore Volume in the range of 55-115Á is only about 20-25% and preferably 20.8%. The Pore Dimension Distribution is such that 29.6-33.0% of the Total Pore Volume, and preferably about 33.0% is present as macropores of diameter greater than 250A. The amount of Total Pore Volume in pores with a diameter greater than 600A is only about 21-27% and preferably 26.6% TPV. The amount of Total Pore Volume in pores having a diameter greater than 1200A is only about 0.15-0.20 cc / g and preferably 0.20 cc / g. It should be noted that the pore percentages in the finished catalyst are essentially the same as in the substrate charge of gamma alumina from which it is prepared, although the Total Surface Area of the finished catalyst may be 75-85%, it is say 80.0% of the gamma alumina substrate charge from which it is prepared (for example 75-85% of a support surface area of 205-275 m '/ g, ie 221 m2 / g). It is also convenient to note that the Medium Pore Diameter per surface area measured by the Mercury porosimeter of the finished catalyst is essentially the same as that of the gamma alumina substrate load from which it is prepared. It is also noted that the Pore Dimension Distribution (percent of the total) in the finished catalyst can be essentially the same as that of the alumina load from which it is prepared (unless most of the distribution of pore volume in a given range is close to a "breaking point" -for example 80 amstroms or 250 amstroms, in which case a small change in the pore amounts of a given dimension can modify the reported value of the pore volume that fall into a reported range). The Total Pore Volume of the finished catalyst can be 75% -98%, that is 81.3% of the alumina load from which it is prepared. The Catalysts of the Present Invention - Metal Charges The extrudates of the alumina charge can be charged with metals to yield a catalyst product containing a Group VIII non-noble metal oxide in an amount of 2.2-6% by weight, preferably 3.0 -3.5% by weight, ie 3.3% by weight and a metal oxide of Group VIB in a quantity of 7-24 by weight, preferably 12.5 to 15.5% by weight, ie 14.4% by weight. The Group VIII metal can be a non-noble metal such as iron, cobalt, or nickel. This metal can be charged to the alumina generally from 10% -30%, ie 15% aqueous solution of a water soluble salt (for example a nitrate, acetate, oxalate, etc.). The preferred metal is nickel, used as about 12.3% by weight of aqueous solution of nickel nitrate hexahydrate Ni (NO-6H_0) The metal of group VI B can be chromium, molybdenum or tungsten.This metal can be charged in alumina generally from 10% -40% ie 20% of an aqueous solution of a water-soluble salt.The preferred metal is molybdenum, employed in about 15.5% by weight aqueous solution of ammonium molybdate tetrahydrate. It is a characteristic of the catalyst of the invention to contain from 0.3 to 2% by weight of P2O5, preferably around 0.5 to 1.5% by weight. This level of phosphorus oxide loading is very small, representing only 0.13 to 0.87% by weight of elemental phosphorus. The phosphorus component can be charged to the alumina as 0 to 4% by weight, say 1.1% by weight of 85% by weight aqueous solution of phosphoric acid H3PO4 in water. As described above, silica Si02 can be incorporated in the catalyst support prior to impregnation and can therefore be present in small amounts, typically below 2.5% by weight and preferably 1.3 to 2.5% by weight, although the benefits of the invention can be achieved without the addition of silica. These metals in the catalyst and phosphorus can be charged to the alumina support by impregnation therein with a solution of the former. Although it is preferable to load the metals simultaneously, it is possible to charge each one separately. Small amounts of H 0 can be added to stabilize the impregnation solution. It is preferable that the catalyst be impregnated by 90 to 105% filler, preferably 97 to 98%, ie 97% of the pore volume of the substrate with the solution containing the required amounts of metal and phosphorus. The loaded is followed by drying and calcination at 482-677 ° C (900 to 1250 °), preferably 621-654 ° C (1150 to 1210 ° F), say 638 ° C (1180 ° F) for 0.5 to 5 hours, that is 1 hour. Another feature of the catalyst composition of the present invention is that the measurement of the constant K of the hydrodesulfurization microactivity rate (HDS) of the catalyst of the present invention tailored to the constant K of the HDS microactivity rate of a standard hydroprocessing catalyst (i.e., Cri-terium HDS-1443B, a prior art catalyst commercially available for use in the hydroprocessing of waste oils), has a value greater than or equal to 0.5 to 1.0, preferably 0.6 to 0.85. As used in this description, the phrase "HDS microactivity" means the intrinsic hydrodesulfurization activity of a catalyst in the absence of diffusion, as measured according to the HDS microactivity test (HDS-MAT), described as follows. In the HDS MAT test, a given catalyst is ground to a fraction of 30-60 MESH (0.071-0.013 MM) and pre-sulphided at 399 degrees C (750 degrees F) with a gas flow containing 10% H2S / 90% H2. The catalyst is then exposed to a sulfur containing feed, ie benzothiophene, which acts as a test model of sulfur compounds, at reaction temperature and with hydrogen flow for about 4 hours. Samples are taken periodically and analyzed by gas chromatogy for the conversion of benzothiophene to ethylbenzene, to thereby indicate the hydrodesulfurization properties of the catalyst being tested. The activity is calculated on both a catalyst weight basis and catalyst volume to take into account any difference in densities between catalysts. The conditions for an HDS-MAT test are as follows: Temperature: around 288 degrees C (around 550 degrees F) Pressure: around atmospheric pressure Stock of Feeding around 0.857 moles of benzothiophene in reactive grade of normal eptane Space Velocity 4 hr "1 Charge of Catalyst 0.5 gram The kinetics of the reactor used in the HDS-MAT test is of the first order, by flow contact. In the temperature and space velocity established above the constant K can be expressed as k = ln (1 / -1-HDS) where HDS is the fractional value of the hydrodesulphurisation obtained by a given catalyst under the conditions stated above, a state catalyst of the technique, commercially available sold to be used in the hydroprocessing of residual oils (HDS-1443B criterion of catalysts) was evaluated with the HDS-MAT test under the conditions established above and was found to have a percentage value of HDS of 73% in weight and a corresponding constant K of 1.3093. The catalyst of the present invention requires that the proportions of the K HDS-MAT constants, relative to those obtained with the HDS-1443B criterion, have values of preferably greater than or equal to 0.5 to 1.0, preferably 0.6 to 0.85 the catalysts of the patent American number 5,047,142 requires having values greater than 1.0, preferably greater than 1.5. It is another feature of the composition of the catalysts of the present invention that the molybdenum oxide, preferably MoO .., is distributed in the porous alumina support described above in such a way that the molybdenum gradient is about 1.0. As used in the present description, the phrase "molybdenum gradient" means the molybdenum / atomic aluminum ratio observed on the outer surfaces of the catalyst pellets relative to the molybdenum / atomic aluminum ratio observed on surfaces of a sample from the same catalyst that has been shredded to a fine powder, the atomic ratio being measured by x-ray photoelectron spectroscopy (XPS), sometimes referred to as electronic spectroscopy for chemical analysis (ESCA). The molybdenum gradient is suspected to be strongly affected by the impregnation of molybdenum in the catalyst support and the subsequent drying of the catalyst during its preparation. The ESCA data were obtained in an ESCALAB MKII instrument available in V.G. Cientific Ltd., which uses a 1253.6 eV magnesium X-ray source. Generally, the finished catalysts of the present invention will be characterized by the properties set forth in Table II where the columns show the following: a) The first column lists the crude 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 volumes occupied by pores that fall in designated ranges-such as one percent by volume of Total Pore Volume (% TPV) or as a percentage of Pore Volume volume in micropores with diameter less than 250A -por example% of Pore Volume in the micropores - or in cc of Pore Volume per gram of catalyst; the Poro Fashion by volume measured by the mercury porosimetry (dV / dD MAX); the Volume of Pore that falls in ± 25Á of dV / dD MAX in the region of smaller than 250Á; and, Area of Surface in m2 / g. b) The third column lists the properties of the best known catalyst mode, Example I. The fourth column lists the specific properties of a second sample, Example II, made by the same formula as in Example I. c) the remaining columns list properties for other catalysts of the state of the art. The catalyst can be evaluated in a Robinson Reactor two stages, a Continuous Withdrawal Tank Reactor (CSTR) which evaluates the catalytic activity under conditions simulating those of the two-stage H-OIL bubbled bed unit. The feed stock is a Medium / Heavy Arabian Oil Vacuum Residue set above. The evaluation is carried out for four or more weeks at a catalyst age of 1.86 or more barrels per pound.

* 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.

Preferred Modality In the practice of the process of this invention, the catalyst, preferably in the form of extruded cylinders of 0.99 to 1.1 mm (0.039 to 0.044 inches) in diameter and about 3.8 mm (0.15 inches) in length, may be located in the first and second stage reactors of a two-stage H-OIL unit. The hydrocarbon charge is admitted to the lower portion of the reactor bed of the first stage in the liquid phase at 343-454 degrees C (650 to 850 degrees F), preferably 371 to 441 degrees C (700 to 825 degrees F) and 6.9 to 24.2 MPa (1000 to 3500 PSIa), preferably 10.4 to 20.7 MPa (1500 to 3000 PSIa). The hydrogen gas is admitted to the reactor of the first stage of the H-OIL unit of the two-stage bubble bed in quantities of 360 to 1800 M3 / M3 (2000 to 10000 SCFV), preferably 540 to 1440 M3 / M3 ( 3000 to 8000 SCFV). The hydrocarbon charge passes through the bubbling bed reactor of the first stage at an LHSV of 0.16 to 3.0 hr ~ ", preferably 0.2 to 2 hr" 1. During operation the first stage reactor catalyst bed is expanded to form a bubble bed with a defined upper level. The operation is essentially isothermal with a typical maximum temperature difference between inlet and outlet of 0 to 27.8 degrees C (0 to 50 degrees F), preferably 0 to 16.7 degrees C (0 to 0 degrees F). The gaseous and liquefied 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 average temperature difference between the reactors of the first and second stage is 0 to 16.7 degrees C (0 to 30 degrees F), preferably 0 to 8.3 degrees C (0 to 15 degrees F). A little additional hydrogen may additionally be injected into the second stage reactor to replenish the hydrogen consumed by the reactions in the first stage reactor. During operation, the catalyst bed of the second stage reactor is also expanded to form a bubble bed with a defined upper level. The constant catalytic activity is maintained by periodic and continuous withdrawals of used catalyst portions and replacements of the removed material with fresh catalyst. The fresh catalyst is typically added in a ratio of 0.0.5 to 1.0 pounds per barrel of fresh feed, preferably 0.20 to 0.40 pounds per barrel of fresh feed. An equal volume of used catalyst is removed and discarded to maintain a constant inventory of catalyst in the base volume. The replacement of the catalyst is carried out in such a way that equal quantities of fresh catalyst are added to the reactor of the first stage and the reactor of the second stage of a two-stage H-OIL unit. In a less preferred embodiment, the reaction can be carried out in one or more continuous withdrawal tank reactors (CSTR) which also provides essentially isothermal conditions. During the passage through the reactor, preferably containing a bubbled bed, the hydrocarbon feed stock is converted to low boiling products by the hydrotreater / hydrocracking reaction. Practice of the Present Invention In a typical embodiment, using a two-stage Robinson reactor pilot unit, a load containing 60 to 95% by weight ie 88.5% by weight of compounds with a boiling point above 538 degrees. C (1000 degrees F) can be converted to a hydrotreated product containing only 28 to 45% by weight, ie 42.0% by weight of products with a boiling point above 538 degrees C (1000 degrees F). The sulfur of the original charge is from 3 to 7% by weight, typically 5.1% by weight; the sulfur content of the unconverted components 538 degrees C (1000 degrees F +) in the product is 0.5 to 3.5% by weight, typically 1.6% by weight. In another embodiment, using a two-stage Robinson Reactor pilot unit operating at +5.6 degrees C (+10 degrees F) above normal operating temperatures and at a higher value of catalyst aging, a load containing 60 to 95% in Weight, ie 88.5% by weight of compounds with a boiling temperature above 538 degrees C (1000 degrees F) can be converted to a hydrotreated product containing only 24 to 38% by weight, ie 35.4% by weight of compounds with boiling point above 538 degrees C (1000 degrees F) the sulfur content of the unconverted components 538 degrees C + (1000 degrees F +) in the product is 0.5 to 3.5% by weight, typically 2.2% by weight. In both modalities, the values of Existing IP sediments of the products 343 degrees C + (650 degrees F +) leaving the reactor are extremely small; < 0.17% by weight. ADVANTAGES OF THE INVENTION It will be apparent those skilled in the art that this invention is characterized by advantages such as the following. (a) Allows the performance increase of hydrocarbon products boiling at a temperature below 530 ° C (1000 ° F); (b) Allows reaching the above-mentioned yields with little or no sediment when measured by the values of Existing IP Sediment of the boiling point products at 343 ° + (650 ° F) (c) allows an improved level of removal of sulfur as seen in the observed hydrodesulfurization (HDS) of the total liquid product and the substantial improvement, low level of sulfur in the unconverted flow at a temperature of 538 degrees C (1000 degrees F); and (d) It allows to improve the reduction of levels of coal waste, and nickel and vanadium removal. The practice of the process of this invention will be apparent to those skilled in the field from the following where all parts are parts by weight if nothing else is said. DESCRIPTION OF THE SPECIFIC MODALITIES More Well-known Reactor Mode Data The same amounts of catalyst of Example I are placed in the reaction chambers (of the Robinson reactors in the first and second stages). The hydrocarbon feedstock (eg, the undiluted Arab / Heavy Dilution vacuum residue, described in Table I) is admitted in the liquid phase to the Robinson reactor of the first stage at a temperature of 404 degrees C (760 degrees F) and a pressure of 15.5 MPa (2250 PSIG). The hydrogen gas is admitted to the Robinson reactor of the first stage in an amount of 1260 MVMJ (7000 SCFV). The hydrocarbon charge passes through the Robinson reactor of the first stage at a liquid hourly space velocity (LHSV) of 0.56 volumes of oil per hour per volume of liquid admitted. This is equivalent to a catalyst space velocity (SCV) of 0.130 barrels of hydrocarbon charge per pound of catalyst per day. The hydrocarbon feed is uniformly brought into contact with the hydrogen and the catalyst under isothermal conditions in the Robinson reactor of the first stage. The liquid and gaseous effluents of the first stage Robinson reactor is then directed to the Robinson reactor of the second stage which is operated at essentially the same temperature and pressure of the first stage Robinson reactor. The difference in average temperature between the first and second stage reactors is nominally 0 degrees C (0 degrees F). No additional hydrogen is injected into the Robinson reactor of the second stage. The liquid effluent passes through the Robinson reactor of the second stage at a liquid hourly space velocity (LHSV) of 0.56 volumes of liquid effluent per hour per volume of liquid admitted. This is equivalent to a catalytic speed (SCV) of 0.130 barrels of liquid effluent per pound of catalyst per day. The effluent The Robinson reactor liquid from the first stage is uniformly brought into contact with the gaseous effluent containing hydrogen and the second catalyst charge under isothermal conditions in the Robinson reactor of the second stage. No attempt is made to maintain a constant catalytic activity by continuous or periodic withdrawals of portions of used catalyst and replacements of the material removed with fresh catalyst in the Robin-son reactor system of the second stage. The catalyst starts as fresh catalyst and accumulates catalytic aging generally expressed in barrels per pound. The average temperature is defined as the average of the temperatures of the reactors of the first and second stage. Especially, the hydrocarbon charge passes through the entire process system (for example the Rombinson reactors of the first and second stages) to an LHSV of 0.28 volumes of oil per hour per volume of liquid admitted. This is equivalent to a CSV of 0.065 barrels of hydrocarbon charge per catalyst li- bra per day. As will be discussed below, the temperatures of the reactors of the first and second stage can be brought to higher levels with the catalysts of the present invention. The products are collected and analyzed over a range of catalytic aging from 0.195 to 1.08 barrels per pound (corresponding approximately to the third to the sixteenth day of the evaluation) to yield the average data shown in Table III below: TABLE II I PROPERTY% Sulfur Removal Value 79.6% Coal Residue Reduction 58.0% Ni Reduction 73.0% Reduction 94.9% i Hydroconversion of Materials 538 ° C + at 538 ° C- (1000 ° F + at 1000 ° F-) (Base * by weight) 52.6 of kinetically adjusted Hydroconversion (to 0. 0650 bbl / day / weight and 404 ° C (760 ° F) of materials of 538 ° C + to 538 ° C- (1000 ° F + to 1000 ° F-) (Base% by weight) 52.6 From Table III, it is apparent that the process of the present invention allows to increase the conversion of materials with a boiling point above 538 degrees C (1000 degrees F) by 46.9% by weight; and the sulfur, the carbon residue, and the metals are removed. In distillation to recover (1) a first cut from the initial boiling point at 343 ° C (650 ° F), (2) a second cut from 343 to 538 ° C (650 ° F to 1000 ° F) , and (3) a third cut above 538 ° C (1000 ° F), the following is noted: From Table IV above, it is apparent that the sulfur content is decreased in all the products of the fractions (from 5.1% in the feed). In distillation to recover (4), or cut off compounds that have a boiling point of about 343 ° C (650 ° F) and higher, the following is noted: From this Table V, it is apparent that the process of the present invention can operate at about 46.9% conversion of feed components with boiling points or above 538 ° C (1000 ° F) to products with lower boiling point at 538 ° C (1000 ° F) without producing any sediment (when measured by both Existing and Accelerated sediment tests). EXAMPLE A COMPARISON WITH FIRST-GENERATION CATALYSTS Comparative data between the catalysts of Example I of the present invention and a first generation of H-OIL nickel / molybdenum catalysts (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 provides: a) No sediment against an undesirable level with an H-OIL nickel / molybdenum catalyst of the first commercially available generation; b) An improved level of 538 ° C + at 538 ° C- (1000 ° F + at 1000 ° F-); c) An improved level of sulfur removed as seen in the observed hydrodesulfurization (HDS) of the liquid product and the substantial improvement, of the low level of sulfur in the unconverted flow 538 ° C (1000 ° F) d) Improved levels of reduction of coal waste and nickel and vanadium removal; Y * * Kinetics of the Order (assuming equal constant for the reactors of the first and second stages; Activation Energy = 65 kcal / mol.

EXAMPLE B DATA A + 5.6 ° C (10 ° F) In the evaluation of the catalysts of Example I of the present invention, the temperature of the reactors was increased to around 5.6 degrees C (10 degrees F) over a period of 2.0 days at a final temperature of approximately 410 degrees C (770 degrees F) (for example the first stage, the second stage, and average temperatures). The product was collected and analyzed over a catalyst age range from 1.28 to 1.86 barrels per pound (corresponding approximately to the 19th to the 28th day of the evaluation) The comparative data between the catalyst of the present invention operating at around +5.6 grades C (+10 degrees F) compared to the H-OIL Molybdenum / nickel catalysts of the first generation (criterion HDS-11443B) at the same age of the catalysts are given in Table VII. The process of the present invention is superior in which it provides: (a) low sediment at a conversion of 60 ° by weight of products 538 degrees C + to 538 degrees C- (1000 degrees F + to 1000 degrees F-) against an undesirable level with catalysts H-OI1 molybdenum / nickel of the first generation operating at only 52% by weight of product conversion 538 degrees C + at 538 degrees C- (1000 degrees F + at 1000 degrees F-); (b) An improved weight percent conversion level of 538 degrees C + at 538 degrees C- (1000 degrees F + to 1000 degrees F-); • by the observed data and once the data from both catalysts are kinetically adjusted to the CSV objectify) an improved level of sulfur removal as can be observed in the hydrodesulfurization (HDS) of the total liquid product and the substantially improved, lower level of sulfur in the unconverted flow 538 degrees C + (1000 degrees F +); and (d) improved levels of carbon residue reduction and nickel and vanadium removal; It was noted that the sulfur levels of the cuts 343 ° C + at 538 ° C + (650 ° F + at 1000 ° F +) bp (approximately the composition of an oil gas in vacuum) were slightly higher with the catalyst of Example I of the present invention operating at around + 5.6 ° C (+ 10 ° F) compared to the level obtained with the first generation catalysts when both had an age of 1.28 to 1.86 barrels per pound. The catalyst of the present invention in addition to providing low sediment results for the boiling point cuts of 343 degrees C + (650 degrees F +), also showed improved operability. The evaluation was smooth at both 404 degrees C (760 degrees F) and 710 degrees C (770 degrees F). On the other hand, the evaluation of catalysts of the first generation showed evidence of encas-tramiento due to accumulated sediment during the course of its evaluation. Operations with the first-generation catalysts became somewhat erratic at around 1.54 BBL / lb of catalyst age and the unit had to be stopped and partially cleaned before the evaluation of the first generation catalysts could be completed With so many problems due to the sediment, it was felt that the temperature could not be increased to any degree with the catalysts of the first generation.

TABLE VII EXAMPLE B Catalyst Test Results in Robinson Reactor Two Stage One Step, Pure Residue, Not Diluted, Once Through Hydrogen Age = 1.28 to 1.86 Barrels per Pound (Increased temperature by 5.6 ° C (10 ° F) from Example I) Catalyst Example I la. Generation 9HDS-1443B ^ 3V, Bt ./ b / Di) \ Jo51 L 643 Temperature "F (" -) 70.3 (410) 0 (05) (Average between stages) 343 ° C + (c5- ° F +) Sediment IP E .-i -tente (% »n weight) 0. 5 '15 Sediment IP Accelerated (% in μ> -so) 0.33) .59 1__ -iZilil -------------. - ie Sulfur Removed 75 .9 *. i "^ iu: >) - -le Resiiuo ie Coal 57.3" * ie Nickel Removed 7J .9 i Je anadii. Remov i io 94.3 4 P hi-r - "'.er.ijn ie materials - •' - - -. * <" - (1 C 0? ° Ft- a 1 i, r-) (? er pes-) - '' 1 - -l 6 re Ajjstado (CSV y • ") i» ulr -investment of materials ° a i8 ° C- (1000ßF + a lij? o'F-) S. C 51.7 1 Weight%) r 1 1 <343 ° C (<65 ° F) jy avedad Specifi c, j / C ~ "J 84 Sulfur,% by weight?:? C rt < = • 2: 343 i-538 ° C (650 ° F-1) 0 ° F) Specific Gravity, j / ~:. 3 Sulfur,% by weight .1J • orte 1: o38 ° C + II i ° F + i 3varity Specifi -a, 'Azuf e,% en pes n * Criterion HDS-1443 Catalyst H-OIL. ** Kinetics of the Order (assuming equal constant for the reactors of the first and second stages; Activation Energy = 65 kcal / mol.

Claims (19)

1. CLAIMS 1. A process for hydrotreating a hydrocarbon feedstock containing components with a boiling point above 538 ° C + (1000 ° F +) and sulfur, metals, and carbon residues, which comprises: contacting said feed of hydrocarbon with hydrogen under isothermal conditions of hydroprocessing in the presence of, as a catalyst, a porous support of alumina containing < 2.5% by weight of silica, and comprising 2.2 to 6% by weight of a Group VIII metal oxide, 7 to 24% by weight of a Group VIB metal oxide and 0.0 to 2.0% by weight of a phosphorus oxide. , said catalyst having a Total Surface Area of 175 to 205 m2 / g, a Total Pore Volume of 0.82 to 98 cc / g, and a Pore Size Distribution where 29.6 to 33.0% of the Total Pore Volume is present as macropores larger than 250A, 67.0 to 70.4% of the Total Pore Volume is present as micropores smaller than 250A, > 65% of the micropore volume is present as micropores of diameter in ± 25A of a pore fashion by volume of 110 to 130A, less than 0.05 cc / g of pore volume is present in micropores with diameters less than 80A, For by medium of this form a hydro-processed product containing a decreased content of components that have their boiling point above 538 ° C + (1000 ° F +) and sulfur, metals and carbon residues, and recover said hydroprocessed product containing decreased content of boiling point components above and sulfur, metals and carbon residues, and recovering said hydroprocessed product containing decreased contents of sediment in the portion of the hydroprocessed products with a boiling point above 343 ° C (650 °).
2. 2. A process as claimed in claim 1, wherein said metal oxide of group VIB is molybdenum oxide in an amount of 12.5 to 15.5% by weight.
3. 3. A process as claimed in claim 1 or 2 wherein Group VIB metal oxide is nickel oxide in an amount of 3.0 to 3.5% by weight.
4. 4. A process as claimed in any preceding clause where the content of Si02 is < 1.3-2.5% by weight.
5. 5. A process as claimed in any of the preceding claims wherein the phosphorus oxide is P20 =, present in the amount < 0.5-1.5% by weight.
6. 6. A process as claimed in any of the preceding claims wherein said Total Surface Area is 175 to 195 pr / g.
7. 7. A process as claimed in any of the preceding claims wherein said Total Pore Volume is 0.82-0.90 cc / g.
8. 8. A process as claimed in any of the preceding claims wherein the distribution of pore diameters of the catalyst is further characterized in that 21-27% of the Total Pore Volume is present in pores with a diameter > 600A and 0.15-0.20 cc / g of the Total Pore Volume is present in pores having a diameter > 1200 Á.
9. 9. In a process for hydrotreating a hydrocarbon feedstock containing components with a boiling point above 538 ° C + (1000 ° F +) and sulfur, metals and carbon residues to form hydroprocessed product containing components with boiling point above 538 ° C + (1000 ° F +) and sulfur, metals and carbon residues decreased and recover said hydroprocessed product containing components with a boiling point above 538 ° C + (1000 ° F +) and sulfur, metals and re-residues of Reduced carbon, an improvement that allows operations at a temperature of 5.6 ° C (10 ° F) higher than normal hydrotreating conditions, increased conversions of components with boiling point above 538 ° C + (1000 ° F +) to product with a boiling point above 538 ° C + (1000 ° F +) at 8% by weight, and reduces the values of the Existing IP Sediment Test in the portion of the hydroprocessed product with boiling point n above 343 ° C (650 ° F) to 0.02% by weight comprising contacting said hydrocarbon feed with hydrogen under isothermal conditions of hydroprocessing in the presence of, as catalyst, a porous support of alumina containing <2.5% by weight of silica, and comprising 2.2 to 6% by weight of a Group VIII metal oxide, 7 to 24% by weight of a Group VIB metal oxide and 0.3 to 2.0% by weight of an oxide of phosphorus, said catalyst having a Total Surface Area of 175 to 205 m2 / g, a Total Pore Volume of 0.82 to 98 cc / g, and a Pore Size Distribution where 29.6 to 33.0% of the Total Pore Volume is present as macropores larger than 250A, 67.0 to 70.4% of the Total Pore Volume is present as micropores smaller than 250A, < 65% of the micropore volume is present as micropores of diameter in ± 25A of a pore fashion by volume of 110 to 130A, less than 0.05 cc / g of pore volume is present in micropores with diameters less than 80A.
10. 10. A process as claimed in claim 9 wherein the Pore Diameter Distribution of the catalyst is further characterized 21 to 27 percent of the TPVs in pores larger than 600 Á and the pores having a diameter greater than 1200 Á of the Total Volume of Pores, they are only around 0.15 to 0.20 cc / g.
11. 11. A process as claimed in claims 9 to 10 wherein the catalyst comprises a support containing 1.3 to 2.5% by weight of silica, and carrying 3.0 to 3.5% by weight of nickel oxide,
12. 12.5 to 15.5% by weight of a molybdenum oxide and 0.5 to 1.5% by weight of a phosphorus oxide, said catalyst having a Surface Area Total of 175 to 195 m2 / g, a Total Pore Volume of 0.82 to 0.98 cc / g. 12. A hydrotreating catalyst characterized by stability at a temperature of 5.6 ° C (10 °) over the normal conditions of the hydrotreatment process, consisting essentially of: a porous support of alumina containing < 2.5% by weight of silica, and comprising 2.2 to 6% by weight of a Group VIII metal oxide, 7 to 24% by weight of a Group VIB metal oxide and 0.3 to 2.0% by weight of an oxide of phosphorus, said catalyst having a Total Surface Area of 175 to 205 prVg, a Total Pore Volume of 0.82 to 98 cc / g, and a Pore Size Distribution where 29.6 to 33.0% of the Total Pore Volume is present as macropores of diameter larger than 250A, 67.0 to 70.4% of the Total Pore Volume is present as micropores of diameter less than 250A, < 65% of the micropore volume is present as micropores of diameter in ± 25A of a pore fashion by volume of 110 to 130A, less than 0.05 cc / g of pore volume is present in micropores with diameters less than 80A.
13. 13. A hydrotreating catalyst as claimed in any of claims 12 wherein the pore diameter distribution of the catalyst is characterized in that 21 to 27 percent of the TPVs in pores larger than 600 Á and the pores having a diameter larger than 1200 Á and 0.15 to 0.20 cc / g of the Total Pore Volume is present in pores having a diameter of > 1200Á.
14. 14. A hydrotreating catalyst as claimed in claim 12 or claim 12 wherein said metal oxide of Group VIB is molybdenum oxide in an amount of 12.5 to
15. 15.5% by weight. 15. A hydrotreating catalyst as claimed in claim 12 or claim 13 wherein said Group VIB metal oxide is nickel oxide in an amount of 3.0 to 3.5% by weight.
16. 16. A hydrotreating catalyst as claimed in claims 12 to 15 wherein the SiO2 content is < 1.3-2.5% by weight.
17. 17. A hydrotreating catalyst as claimed in any of claims 13 to 16 wherein the phosphorus oxide is P205 present in the amount < 0.5-1.5% by weight.
18. 18. A hydrotreating catalyst as claimed in any of claims 12 to 17 wherein said Total Area Area is 275-195 m2 / g.
19. 19. A hydrotreating catalyst as claimed in any of claims 12 to 18 wherein said Total Pore Volume is 0.82 to 0.90 cc / g. SUMMARY A process for hydrotreating a hydrocarbon feedstock containing a boiling point higher than 538 ° C (1000 ° F), sulfur, metals and carbon residues in products containing decreased levels of components having a higher boiling point that 538 ° C (1000 ° F), decreased sulfur levels, particularly decreased sulfur content in products higher than unconverted 538 ° C (1000 ° F) and reduced sediment, which comprises: contacting said feed of hydrocarbon with hydrogen under isothermal conditions of hydroprocessing in the presence of, as a catalyst, a porous support of alumina containing < 2.5% by weight of silica and comprising 2.2 to 6% by weight of a Group VIII metal oxide, 7 to 24% by weight of a Group VIB metal oxide and 0.3 to 2.0% by weight of a phosphorus oxide, said catalyst having a Total Area Area of 175 to 205 m2 / g, a Total Pore Volume of 0.82 to 98 cc / g, and a Pore Size Distribution where 29.6.0 to 33.0% of the Total Pore Volume is present as macropores larger than 250A, 67.0 to 70.4% of the Total Pore Volume is present as micropores smaller than 250A, < 65% of the micropore volume is pre-sat as micropores of diameter in ± 250A of a pore fashion by volume of 110 to 130A, less than 0.05 cc / g of pore volume is present in micropores with diameters less than 80A.