Process For Conversion Of Residual Oils

Abstract

A process for hydrocracking petroleum stocks containing both residual components and metallic contaminants employing a particular catalyst. The catalyst is comprised of a metalliferous hydrogenating component on a siliceous carrier and has a minimum surface acidity and minimum Specific Volume of Pores, defined as an inter-relationship of density, pore volume and pore-size distribution.

Parent Case Text

This application is a continuation-in-part of our copending application Ser. No. 56,748, filed July 20, 1972, now abandoned.

Description

Our invention relates to a process for hydrocracking petroleum stocks containing residual components and metallic contaminants. More particularly, our invention is directed to such a process employing a siliceous catalyst having a particular surface acidity and a particular inter-relationship of density, pore volume and pore-size distribution.

Residual petroleum oil fractions are relatively less valuable than lighter distillate fractions and thus the desirability of converting the comparatively higher boiling residual materials to lower boiling, more valuable products, such as gasoline and furnace oil, is well recognized. While the hydrocracking of such residual stocks has previously been suggested, such previous hydrocracking techniques suffer from the deficiency of comparatively short catalyst life. It is believed that the shortened catalyst life obtained in the hydrocracking of residual stocks is due to a very great extent to contamination of the catalyst by deposition of metals on the catalyst surface. The seriousness and magnitude of this problem will be realized from a consideration of the following items. First, it has generally been found that the degree of metals removal from a hydrocarbon stock and thus the amount of metals available for deposition on the catalyst is related to the extent of the hydrocracking. Next, the severity of the conditions required in treatment of a residual stock is quite high in order to effect the extensive hydrocracking necessary to produce a satisfactory yield of lower boiling materials. Finally, the residual components of a petroleum stock are comprised of the larger, metals-containing molecules and thus the preponderant portion of, if not the total, metals content of a full crude is concentrated in the residual fraction. It will be seen, therefore, that catalyst deactivation due to deposition of contaminant metals on the catalyst presents a severe obstacle to the extension of catalyst life in hydrocracking residual stocks.

We have discovered a process for hydrocracking a petroleum oil that contains residual components and metallic contaminants, while effecting a decreased level of demetallization of feed stock but a satisfactorily high level of conversion, which process comprises contacting the oil at hydrocracking conditions with hydrogen and a particular type catalyst. The particular catalyst to be employed in our invention comprises a metalliferous hydrogenating component composited with a siliceous carrier containing at least about 15 percent by weight of silica based upon the carrier. The siliceous carrier must also have a surface acidity of at least about 6 cc. of ammonia per gram of carrier (measured at a nominal temperature of 340.degree.F.) Additionally, the overall catalyst composite, i.e., hydrogenating components and siliceous carrier, must have a particular inter-relationship of compacted density (D.sub.c) expressed in grams per cc. (g/cc.), total pore volume (V.sub.t) of pores having a radius in the range from 7 to 300 A expressed as cc. per gram (cc/g) and volume per cent of the total pore volume in the form of pores having radii in the range from 50 to 300 A (Z). This inter-relationship is termed the Specific Volume of Pores (SVP) and can be represented by the equation SVP = D.sub.c .times. V.sub.t .times. Z.

The particular catalyst of our invention must have an SVP of at least 20 and preferably at least about 22. Generally, the maximum feasible value for SVP is considered to be about 40, although we prefer to employ catalysts having an SVP no greater than about 35.

While the SVP itself is a determining factor in the selection of a catalyst in accordance with our invention, the values of the individual physical characteristics of the catalyst contributing to the SVP are also important. Generally, the value of D.sub.c will be greater than about 0.3 grams per cc. and preferably will be greater than 0.4 grams per cc. Similarly, the value for D.sub.c generally will not exceed about 0.8 grams per cc. and usually will be less than about 0.7 grams per cc. We prefer, however, to employ catalysts having a compacted density in the range from about 0.5 to about 0.6 g/cc. The total pore volume (V.sub.t) will usually be at least about 0.4 cc./g and preferably will be greater than 0.5 cc. per gram. Generally, however, the value of V.sub.t will not be greater than about 0.8 cc./g. The value for Z will generally be at least about 50 per cent, preferably at least about 60 per cent, and more preferably at least about 70 per cent or greater. Additionally, we prefer that at least about 50 percent of the total pore volume (V.sub.t) be in pores having radii in the range from 50 to 200 A and that the ratio of pore volume of pores having a radius from 50 to 100 A to the pore volume of pores having a radius from 100 to 200 A be in the range from about 0.75:1 to about 1.5:1.

While the carrier employed in the catalyst of our invention must have a silica content of at least 15 percent by weight based upon the carrier, we prefer to employ a carrier containing at least 20 percent by weight silica and more preferably containing at least 30 percent by weight silica. Usually the maximum silica content of the catalyst carrier of our invention will not exceed about 90 percent by weight and generally the silica content will be less than 80 percent by weight. In this connection it should be noted that we have found carriers having a silica content of less than about 70 percent to be quite satisfactory. Further, we prefer to employ carriers having a surface acidity greater than 8 cc./g and preferably greater than about 10 cc./g.

The metalliferous hydrogenating component employed in the catalyst in our invention can be any of such materials well known in the art or combinations of two or more of such materials. Generally, the hydrogenating component can be any one or more of the Group VI or Group VIII metals, their oxides or their sulfides, such as, for example, nickel, cobalt, molybdenum, tungsten or Group VIII noble metals. Usually combinations of these metalliferous components are quite satisfactory, such as, for example, nickel-cobalt-molybdenum, nickel-molybdenum, cobalt-molybdenum, and nickel-tungsten. Additionally, the catalyst composites employed in our invention can be promoted with a small quantity of a halogen such as, for example, fluorine. Usually the halogen content can be in the range from about 0.1 to 10 percent by weight based upon the total catalyst and preferably from about 0.5 to 5 percent by weight. We particularly prefer to employ fluorine as the halogen promoter in the quantity from about 1 percent to about 3 percent by weight based upon the total catalyst.

Catalysts of the type described above and suitable for employment in our invention will have a surface area of at least about 100 M.sup.2 /g. and preferably will have a surface area in the range from about 150 to about 300 M.sup.2 /g.

In the practice of our invention we prefer to employ catalysts of comparatively small particle size, i.e., particles having a smallest dimension of less than about 1/8 inch, preferably less than about 1/16 inch, and more preferably about 1/32 inch or less. Although it is not necessary that the entire catalyst bed be composed of these small size particles, the proponderant portion of the bed is preferably comprised of such small catalyst particles.

Although it is not believed it is necessary to prepare the catalysts required by our process in any particular manner, satisfactory catalysts can be prepared by first synthesizing a silica-alumina carrier in accordance with the general technique described in U.S. Pat. No. 2,469,314. Broadly, the technique of this patent comprises forming a silica sol which upon aging sets as a firm hydrogel. This hydrogel is then admixed with a solution of an aluminum salt and precipitation from the mixture is effected. In keeping with the general teachings of U.S. Pat. No. 2,469,314, particularly preferred procedures such as those described in U.S. Pat. No. Re-issue 23,438 can also be employed to provide a suitable silica-alumina carrier. After formation of this siliceous material, it can be formed into particles of desired size and shape through known techniques, such as, for example, extrusion. The metalliferous hydrogenating components required by our process can then be composited with the siliceous carrier by any of the techniques well known in the catalyst art, such as, for example, impregnation. It is believed that various other techniques for producing catalysts of required characteristics will be evident to those skilled in the catalyst art upon being informed of the required physical properties to be possessed by the catalyst employed herein.

The feed stock suitable for employment in our invention includes any petroleum hydrocarbon stock containing a significant quantity of residual components and metallic contaminants. As used herein the terms "residual," "residue" and "residual fraction or component" are meant to describe the most difficultly vaporizable components of a petroleum oil which normally cannot be vaporized at temperatures below that at which thermal decomposition or cracking would occur at atmospheric pressure. These materials encompass the highest boiling components of a crude oil. Generally, these residual components or fractions can be separated from the lighter, lower boiling components of a crude oil by employment of a vacuum distillation. Normally a residue can be described as boiling above about 1,050.degree.F. or 1,100.degree.F. Illustrative of petroleum oils containing a significant quantity of residual components are vacuum tower bottoms, atmospheric tower bottoms and reduced or topped crudes.

Since the catalysts of the class disclosed herein appear to have both the ability to effect extensive hydrocracking with diminished demetallization of the treated stock and an especially high tolerance for metallic contaminants normally tending to act as catalyst poisons, the present process is particularly advantageous in connection with the treatment of petroleum oils containing a significant quantity of metallic components as are present in residual fractions. This is not to say that the process of our invention is not operative in the treatment of non-residual containing distillate stocks, but the particular advantages provided in accordance with our invention are not obtained when hydrocracking completely distillate charge stocks. Accordingly, therefore, the charge stock to our process is comprised of at least 10 percent by volume residual material and preferably at least about 25 percent by volume residuals. It will be understood, of course, that the process of our invention provides increasingly advantageous results with increasing residual content of the charge and that the maximum advantage is obtained when treating a 100 percent residual charge stock. Similarly, the charge stock to our process will contain at least about 10 p.p.m. weight of contaminant metals, such as for example, nickel and vanadium, and usually will contain at least about 25 p.p.m. total metals content. Our invention is particularly suitable for the treatment of residual containing stocks which contain more than about 50 p.p.m. metals and especially stocks which contain more than about 75 p.p.m. metals.

The operating conditions employed in the hydrocracking operation of our invention include a temperature in the range from about 650.degree.F. to about 900.degree.F., preferably in the range from about 700.degree. to 850.degree.F. The pressure employed in this process can be in the range from about 1,000 to 10,000 psig. Preferably, we employ a hydrogen partial pressure in the range from about 1,500 to about 5,000 psia with hydrogen partial pressures in the range of 2,000 to about 3,000 psia being particularly preferred. In this connection it should be pointed out that while it is preferred to employ hydrogen of comparatively high purity such as for example 85 percent by volume or greater, it is quite satisfactory to employ hydrogen containing gas streams of the type normally found in a refinery operation, such as for example, reformer off gas, containing a minimum hydrogen concentration of about 65 percent by volume. Generally, the hydrogen feed rate will be in the range from about 2,000 to about 30,000 standard cubic feet per barrel (SCF/B), preferably in the range from about 5,000 to about 15,000 SCF/B and more preferably in the range from 6,000 to about 12,000 SCF/B.

Normally, the amount of hydrogen consumed during our hydrocracking operation will be at least about 300 SCF/B and can range up to about 3,000 SCF/B, usually, however, hydrogen consumption will be in the range from about 500 SCF/B up to about 1,500 to 2,000 SCF/B. It will be understood, of course, that hydrogen consumption will vary somewhat based upon the composition of the particular feed stock being treated as well as, to a certain extent, upon the selection of operating conditions.

In the hydrocracking operation of our invention, we also employ a liquid hourly space velocity (LHSV) in the range from about 0.1 to 10 volumes of charge stock per volume of catalyst per hour. Preferably, we employ an LHSV in the range from about 0.1 up to about 1.0 or 2.0, with space velocities in the range from about 0.2 to about 0.5 being particularly preferred.

In order to illustrate our invention in greater detail, reference is made to the following examples.

EXAMPLE I

In this example, a vacuum residue having the inspections shown in Table 1 below was subjected to hydrocracking in two separate runs employing an alumina-supported catalyst in one run and a silica-alumina catalyst, as required by this invention, in the other run.

Table 1 ______________________________________ Inspections Gravity .degree.API 4.7 Sulfur, wt.% 5.61 Nitrogen, wt.% 0.44 Carbon residue, ASTM D-524 wt.% 23.12 Nickel, PPM 38 Vanadium, PPM 142 Carbon, wt.% 83.68 Hydrogen, wt.% 9.88 Oxygen, wt.% 0.39 ______________________________________

Both of the catalysts were 1/32 inch extrudates and were comprised of nickel, cobalt and molybdenum as the hydrogenating components.

The silica-alumina catalyst required in this invention was prepared by first forming a silica sol at a pH of less than 7 and aging of the sol to form a silica hydrogel, after which the hydrogel was suspended in a solution of an aluminum salt and precipitation then effected to produce a material containing about 25 percent by weight alumina. This preparative technique is in accordance with the disclosure of U.S. Pat. No. 2,469,314. This silica-alumina material was then extruded to provide the 1/32 inch extrudates mentioned above and the metalliferous hydrogenating components were added by impregnation of the extrudates. More specifically, the extrudates were first impregnated with a solution of ammonium paramolybdate and then dried. After drying, the molybdenum-containing extrudate was subjected to a second impregnation with a solution of nickel and cobalt nitrates. Subsequent to the second impregnation, the material was dried and then calcined at 1,000.degree.F. Inspection data for both the alumina supported catalyst and the silica-alumina catalyst required in this invention are shown in Table 2 below.

Table 2 __________________________________________________________________________ Inspections Alumina Silica-alumina of this invention __________________________________________________________________________ Compacted density, g/cc. 0.765 0.569 Surface acidity cc NH.sub.3 /g (measured on support) at 342.degree.F. 9.48 11.26 Surface area M.sup.2 /g 146.9 208.9 Pore vol. cc./g 0.47 0.58 Pore size distribution, % of pore volume 200-300 A radius 4.0 ) 5.5 ) ) ) 100-200 A radius 36.2 ) 78.1 27.9 ) 72.3 ) ) 50-100 A radius 37.9 ) 38.9 ) 40-50 A radius 7.0 9.1 30-40 A radius 7.6 8.2 20-30 A radius 7.3 10.4 7-20 A radius 0.0 0.0 Specific vol. of pores cc./100 cc. 28.0 23.8 __________________________________________________________________________

After start up, both runs were continued for a period sufficient to take the edge off the catalyst and to achieve lined out operation.

The present discussion is directed to the main conversion portion or lined out portion of the runs. At a throughput of about 130 volumes of feed per volume of catalyst, the temperature required to maintain a 50 percent conversion was about 770.degree.F. for the alumina based catalyst and about 775.degree.F. for the silica-alumina based catalyst. At this same point in the runs the alumina based catalyst removed about 98 weight per cent of the metals, i.e. nickel and vanadium, present in the feedstock while the silica-alumina based catalyst removed only about 88 percent of the metals present in the feedstock. Both runs were continued and the operating temperature of each run was adjusted so as to maintain the 50 percent conversion. At a throughput of approximately 200 volumes of feed per volume of catalyst the temperature required to maintain the 50 percent conversion in the run employing the alumina based catalyst had increased to about 775.degree.F. while the temperature required in the run employing the silica-alumina based catalyst had increased to only about 778.degree.F. At this point in the runs, the alumina based catalyst was still effective to remove approximately 98 percent of the metals in the fresh feed while metals removal with the silica-alumina supported catalyst had declined to about 84 percent. Throughout the balance of the run the metals removal effected with the alumina based catalyst remained substantially constant in the range from about 97.5 to 98 percent metals removal. On the other hand, however, the metals removal effected with the silica-alumina based catalyst continued to decline to a level of about 55 percent, or just slightly greater, at a throughput of about 740 volumes per volume and then appeared to stabilize at about this level until the run was terminated at a throughput of about 1,000 volumes per volume.

During the operation of these runs the data further indicated that the temperature required for maintenance of the 50 percent conversion level with both catalysts was substantially equal at a level of about 780.degree.F. in the period encompassing a throughput from about 250 to about 260 volumes of feed per volume of catalyst. Thereafter, temperature increase required to maintain the 50 percent conversion in each of the runs continued at the same rate that had previously been established so that the temperature required to maintain the 50 percent conversion level with the silica-alumina catalyst was only 782.degree.F. at a throughput of 340 volumes per volume, while a temperature of 786.degree.F. was required at the same throughput with the alumina based catalyst. From extrapolation of these data it can be determined that the alumina based catalyst will reach a cut off temperature of 800.degree.F. in order to maintain a 50 percent conversion before achieving a throughput of 550 volumes per volume, while the silica alumina catalyst will require an operating temperature of only 788.degree.F. at a throughput of 550 volumes per volume. In fact the run with the silica alumina based catalyst was continued and required a temperature of 793.degree.F. to maintain the 50 percent conversion at 600 volumes per volume throughput.

From this example it can be seen that, although the large-pored alumina based catalyst may provide some temperature advantage over the large-pored silica alumina based catalyst of this invention during the earlier stages of operation, such temperature advantage is rapidly lost. Further it can be seen that the overall rate of deactivation of the alumina catalyst is substantially greater than the deactivation rate of the silica-alumina catalyst of this invention, thus showing that the silica-alumina based catalyst of this invention will provide a substantially greater catalyst life than that obtainable with the alumina based catalyst. It will also be seen that this deactivation rate appears to be due primarily to the greater metals removal effected by the alumina based catalyst as compared to that produced by the silica-alumina based catalyst. Moreover, it should be noted that the level of metals removal achieved by the alumina catalyst is substantially constant throughout the main conversion portion of the run while the level of metals removal achieved with the silica-alumina based catalyst of this invention advantageously decreases as the run progresses and it is believed that this lower and decreasing rate of metals removal of the silica-alumina based catalyst is effective to increase further the life of such catalyst.

EXAMPLE II

In this example the same vacuum residue as employed in Example I was subjected to hydrocracking in two separate runs employing the silica-alumina catalyst of this invention as described in Example I in one run and a second silica-alumina catalyst representative of typical commercially available silica-alumina hydrocracking catalysts in the other run. The catalysts employed in the runs of this example were 1/32 inch extrudates comprised of nickel, cobalt and molybdenum as the hydrogenating components. The inspection data for the typical commercial silica-alumina catalyst are shown in Table 3 below.

Table 3 ______________________________________ Inspections Commercial Silica-alumina ______________________________________ Compacted density, g/cc. 0.757 Surface acidity cc NH.sub.3 /g (measured on support) at 342.degree.F. 11.14 Surface area M.sup.2 /g 217.5 Pore vol. cc/g. 0.46 Pore size distribution, % of pore vol. 200-300 A radius 2.0 ) ) 100-200 A radius 7.7 ) 54.4 ) 50-100 A radius 44.7 ) 40-50 A radius 14.9 30-40 A radius 13.4 20-30 A radius 12.7 7-20 A radius 4.4 Specific vol. of pores cc/100 cc 19.0 ______________________________________

Both of the runs of this example were conducted employing the operating conditions of temperature, pressure and space velocity shown in Table 4 below. Table 4 also shows the inspection data obtained at various times during the course of the runs of this example.

Table 4 __________________________________________________________________________ Silica-alumina of this Invention Catalyst Age:Days 9 33.5 66.5 109.5 121 LHSV vol./hr./vol. 0.35 Reactor temperature .degree.F. 775 780 790 815 820 (95% H2) Pressure:psig 2500 2500 2500 2800 2800 Inspections Gravity: .degree.API 18.7 18.1 14.5 12.7 12.7 Sulfur, wt% 0.50 0.88 2.60 3.78 3.75 Nickel, ppm 5.1 8.4 22 21 -- Vanadium, ppm 7.9 16 49 38 -- Distillation, vacuum 5% 507 492 489 454 448 10% 610 590 586 539 538 30% 896 868 877 862 843 40% 986 967 -- 999 -- Commercial Silica-Alumina Catalyst Age:Days 10.8 34.0 LHSV vol./hr./vol. 0.35 0.35 Reactor temperature .degree.F. 775 790 (95% H2) Pressure:psig 2500 3000 Inspections Gravity:.degree.API 23.3 13.7 Sulfur, wt% 0.31 3.46 Nickel:ppm 8.1 -- Vanadium:ppm 21 -- Distillation, vacuum 5% 508 536 10% 585 562 20% 712 747 30% 826 -- 50% 1009 -- __________________________________________________________________________

From the data shown in Table 4 above, it will be seen that although at a very early stage of operation (i.e., 9 days vs. 10.8 days) the amount of metals removed employing the silica-alumina catalyst of this invention was noticeably greater than the metals removal provided by the commercial silica-alumina, the commercial silica-alumina had become so deactivated by 34 days of operation that a 15.degree. increase in operating temperature together with a 500 pound increase in pressure was ineffective to provide adequate cracking activity to warrant continuation of the run. This can be seen from the change in API gravity of product from 23.3 to 13.7 and an increase of the sulfur content of the product from 0.31 up to 3.46 percent. During a comparable period of operation up to 33.5 days the silica-alumina catalyst of the present invention required only a 5.degree.F. increase in operating temperature with no alteration in pressure in order to maintain substantially a constant level of conversion. This can be seen from the fact that the API gravity and sulfur content of the product changed only slightly. Quite significantly, however, it will be seen that the amount of metals removal effected with the catalyst of this invention had declined substantially from a total metals content in the product of 13 ppm up to 24.4 ppm.

Further examination of the above data will show that at 66.5 days of operation the silica-alumina catalyst of this invention required a temperature of only 790.degree.F. and a pressure of 2,500 psig in order to provide a conversion level noticeably greater than that obtained after only 34 days of operation with the commercial silica-alumina. Additionally, it will be noted that a further decrease in metals removal was effected during the period from 33.5 to 66.5 days.

The remaining data regarding the run employing the silica-alumina catalyst of this invention demonstrate the continued sufficiency of activity of such catalyst so as to permit operation for a period extending through 120 days. Thus, it can be seen that the silica-alumina based catalyst of this invention, although showing a somewhat higher metals removal during earlier stages of operation as compared to the commercial silica-alumina, demonstrates a reduction in the rate of metals removal as the time on stream progresses and tends to stabilize at this reduced rate of metals removal. Further, it would appear that the silica-alumina catalyst of this invention is capable of retaining its cracking activity while tolerating a greater quantity of metal deposition. On the other hand, however, the commercial silica-alumina, although showing an initial lower rate of metals deposition, evidently deactivates extremely rapidly apparently due to the fact that such catalyst cannot tolerate any significant quantity of deposited metals before showing a decline in activity to a completely unsatisfactory level.

EXAMPLE III

In this example the same vacuum residue as employed in Examples I and II was subjected to hydrocracking in two separate runs employing in one run the silica-alumina catalyst of this invention as described in Example I and in the other run a second silica-alumina catalyst of comparable specific volume of pores but with a low surface acidity, outside the range required by the present invention. Again as in Examples I and II, the catalysts employed in the runs of this example were 1/32 inch extrudates comprised of nickel, cobalt and molybdenum as the hydrogenating components. The inspection data for the low surface acidity silica-alumina catalyst are shown in Table 5 below.

Table 5 __________________________________________________________________________ Inspections Low Acidity Silica-Alumina __________________________________________________________________________ Compacted density, g/cc 0.768 Surface acidity cc NH.sub.3 /g (measured on support) at 342.degree.F. 4.7 Surface area M.sup.2 /g 128.6 Pore vol. cc/g 0.41 Pore size distribution % of pore vol. 200-300 A radius 1.5 ) ) 100-200 A radius 18.7 ) 81.9 ) 50-100 A radius 61.7 ) 40-50 A radius 9.1 30-40 A radius 6.3 20-30 A radius 2.5 7-20 A radius 0.0 Specific vol. of pores cc/100 cc 25.8 __________________________________________________________________________

It will be recalled that the silica-alumina of this invention described in Example I had a surface acidity of 11.26 and a specific volume of pores of 23.9.

The particular operating conditions employed for the two runs of this example together with charge stock and products inspections are set forth in Table 6 below.

Table 6 __________________________________________________________________________ Silica-alumina Silica-alumina low surface Catalyst of this Invention acidity __________________________________________________________________________ Age: Days 9 7.7 LHSV vol./hr./vol. 0.35 0.35 Reactor temperature: .degree.F. 775 775 (95% H.sup.2) Pressure: psig 2500 2500 Inspections Charge Gravity: .degree.API 4.7 18.7 16.9 Sulfur: wt.% 5.61 0.50 0.48 Nitrogen: wt.% 0.44 0.19 0.19 Carbon Residue: wt.% 23.12 6.89 6.70 Nickel: ppm 28 5.1 3.9 Vanadium: ppm 142 7.9 6.0 Distillation vacuum, .degree.F. 5% 507 543 10% 610 637 20% 776 776 30% 896 894 40% 986 1001 __________________________________________________________________________

From the data of Table 6 above it will be seen that under substantially comparable operating conditions the silica-alumina catalyst of the present invention was effective at 9 days of operation to provide a satisfactory degree of hydrocracking as indicated by the product API gravity of 18.7 and the production of substantial quantities of lower boiling components. As distinguished from this, the silica-alumina catalyst of low surface acidity, after a comparatively shorter period of operation, had a noticeably lower cracking activity than the catalyst of this invention as can be seen from the API gravity of the product (16.9) as well as the smaller quantity of lower boiling components in the product. Furthermore, it will be noted that the product obtained with the low surface acidity silica-alumina catalyst contained somewhat less than 10 parts per million of metals while the product obtained with the catalyst of our invention contained 13 parts per million of metals.

Claims

1. A process for hydrocracking a petroleum oil containing residual components and metallic contaminants, which process comprises contacting the oil at hydrocracking conditions with hydrogen and a catalyst comprising a metalliferous hydrogenating component composited with a siliceous carrier containing at least about 15 percent by weight of silica based upon the carrier, the carrier having a surface acidity of at least about 6 cc. of ammonia per gram of carrier measured at a nominal temperature of 340.degree.F., and the catalyst having an inter-relationship of compacted density, total pore volume and volume percent pore volume of pores having radii in the range from 50 to 300 A

2. The process of claim 1 wherein the Specific Volume of Pores is at least

3. The process of claim 1 wherein the compact density is in the range from about 0.3 to about 0.8 grams per cc., the total pore volume is in the range from about 0.4 to about 0.8 cc per gram and the volume percent of the total pore volume in the form of pores having radii in the range of

4. The process of claim 1 wherein at least 50 percent of the total pore

5. The process of claim 1 wherein the ratio of pore volume of pores having a radius from 50 to 100 A to the pore volume of pores having a radius from

6. The process of claim 1 wherein the catalyst carrier contains from about

7. The process of claim 1 wherein the carrier has a surface acidity greater

8. The process of claim 1 wherein the metalliferous hydrogenating component is selected from the group consisting of Group VI and Group VIII metals,

9. The process of claim 1 wherein the petroleum oil contains at least about

10. The process of claim 1 wherein the petroleum oil contains at least about 10 parts per million by weight of contaminant metals.