LIQUID PHASE HYDROPROCESSING WITH LOW PRESSURE DROP

Abstract

A process for hydroprocessing a hydrocarbonaceous feedstock in a continuous liquid phase utilizes a hydroprocessing catalyst comprising pills that have a largest dimension that averages no more than 1.27 mm ( 1/20 inch) and more than 100 nm to produce a hydrocarbonaceous product stream.

Description

FIELD OF THE INVENTION

The field of art to which this invention pertains is the catalytic hydroprocessing of hydrocarbons to useful hydrocarbon products. More particularly, the invention relates to catalytic hydroprocessing in continuous liquid phase.

BACKGROUND OF THE INVENTION

Petroleum refiners often produce desirable products such as turbine fuel, diesel fuel, middle distillates, and gasoline boiling hydrocarbons among others by hydroprocessing a hydrocarbon feedstock derived from crude oil or heavy fractions thereof. Hydroprocessing can include, for example, hydrocracking, hydrotreating, hydroisomerization, hydrodesulfurization and the like. Feedstocks subjected to hydroprocessing can be vacuum gas oils, heavy gas oils, and other hydrocarbon streams recovered from crude oil by distillation. For example, a typical heavy gas oil comprises a substantial portion of hydrocarbon components boiling above about 371° C. (700° F.) and usually at least about 50 percent by weight boiling above 371° C. (700° F.), and a typical vacuum gas oil normally has a boiling point range between about 315° C. (600° F.) and about 565° C. (1050° F.).

Hydroprocessing is a process that uses a hydrogen-containing gas with suitable catalyst(s) for a particular application. In many instances, hydroprocessing is generally accomplished by contacting the selected feedstock in a reaction vessel or zone with the suitable catalyst under conditions of elevated temperature and pressure in the presence of hydrogen as a separate phase in a three-phase system (gas/liquid/solid catalyst). Such hydroprocessing is commonly undertaken in a trickle-bed reactor where the continuous phase is gaseous and not liquid.

In the trickle bed reactor, an excess of the hydrogen gas is present in the continuous gaseous phase. In many instances, a typical trickle-bed hydroprocessing reactor requires up to about 1,685 Nm³ of hydrogen per m³ of oil (10,000 SCF/bbl) at pressures up to 17.3 MPa gauge (2500 psig) to effect the desired reactions. However, even though the trickle bed reactor has a continuous gaseous phase due to the excess hydrogen gas, it is believed that the primary reactions are taking place in the liquid-phase in contact with the catalyst, such as in the liquid filled catalyst pores. As a result, for the hydrogen gas to get to the active sites on the catalyst, the hydrogen must first diffuse from the gas phase into the liquid-phase and then through the liquid to the reaction site adjacent the catalyst.

Under some hydroprocessing conditions the hydrogen supply available at the catalytic reaction site may be a rate limiting factor in the hydroprocessing conversions. For example, hydrocarbon feedstocks can include mixtures of components having greatly differing reactivities. While it may be desired, for example, to reduced the nitrogen content of a vacuum gas oil to very low levels prior to introducing it as a feed to a hydrocracking reactor, the sulfur containing compounds of the vacuum gas oil will also undergo conversion to hydrogen sulfide. Many of the sulfur containing compounds tend to react very rapidly at the operating conditions required to reduce the nitrogen content to the desired levels for hydrocracking. The rapid reaction rate of the sulfur compounds to hydrogen sulfide will tend to consume hydrogen that is available within the catalyst pore structure thus limiting the amount of hydrogen available for other desired reactions, such as denitrogenation. In these circumstances, if the diffusion of hydrogen through the liquid to the catalyst surface is slower than the kinetic rates of reaction, the overall reaction rate of the desired reactions (i.e., denitrogenation, for example) may be limited by the hydrogen supply and diffusion. Ideas to overcome the limitations posed by this phenomenon of hydrogen depletion, by manufacturing hydroprocessing catalysts in small sizes and more conventional shapes such as spheres or cylinders are dismissed in conventional continuous gas phase hydroprocessing due to the concern that such small catalysts can create large pressure drops in the reactor.

The catalyst pill size in hydroprocessing is limited to larger sizes and special shapes such as pills with largest dimensions larger than 1.27 mm ( 1/20 inch) and trilobes or quadralobes with lengths as large as 3.2 mm (⅛ inch) to reduce pressure drop. The pressure drop under typical trickle-bed hydroprocessing reaction conditions is exacerbated by recycle rates of hydrogen-rich gas that are five to ten times the chemical hydrogen consumption. Even with the larger catalyst dimensions, mass flux rates have been kept below 29,300 kg/h-m² (6,000 lb/h-ft²) to avoid excessive pressure drop in hydroprocessing.

Continuous liquid phase hydroprocessing with a liquid hydrocarbon stream and solid catalyst has been proposed to convert certain hydrocarbon streams into more valuable hydrocarbon streams in some cases with less hydrogen requirements.

Although a wide variety of process flow schemes, operating conditions and catalysts have been used in commercial activities, there is always a demand for new hydroprocessing methods which provide lower costs, ease of construction, higher liquid product yields and higher quality products.

BRIEF SUMMARY OF THE INVENTION

In continuous liquid phase hydroprocessing the recycle gas stream is eliminated. Therefore, the pressure drop under these conditions becomes very low even at higher than typical conventional technology trickle-bed mass fluxes. We have found that flow distribution can be problematic with very low pressure drop per length of reactant flow path.

The present invention is a process for hydroprocessing a hydrocarbonaceous feedstock which comprises introducing a liquid phase stream comprising a hydrocarbonaceous feedstock and a sufficiently low hydrogen concentration to maintain a continuous liquid phase into a hydroprocessing reactor. The hydroprocessing reactor contains hydroprocessing catalyst comprising pills that have a largest dimension that averages no more than 1.27 mm ( 1/20 inch) and more than 100 nm to produce a first hydrocarbonaceous product stream. In an aspect the pills have a largest dimension that averages no more than 0.85 mm ( 1/30 inch) and no less than 0.51 mm ( 1/50 inch). In an additional aspect, the pills have a largest dimension that averages no more than 0.51 mm ( 1/50 inch).

Other embodiments of the present invention encompass further details such as types and descriptions of feedstocks, hydrocracking catalysts, hydrotreating catalysts, and preferred operating conditions including temperatures and pressures, all of which are hereinafter disclosed in the following discussion of each of these facets of the invention.

BRIEF DESCRIPTION OF THE DRAWING

The FIGURE is a simplified process flow diagram of preferred embodiments of the present invention. The FIGURE is intended to be schematically illustrative of the present invention and not be a limitation thereof. While the FIGURE depicts a reactor as operating in a downflow mode it is presented for illustrative purposes and is not intended to exclude an upflow mode of operation.

DETAILED DESCRIPTION OF THE INVENTION

This invention uses catalyst pills with smaller effective diameter than is commercially practiced in a hydroprocessing flow scheme. The minimum particle size commercially practiced is typically at least 1.27 mm ( 1/20 inch). This invention uses a pill smaller than 1.27 mm. Under continuous liquid phase hydroprocessing conditions, a pressure drop per length of catalyst path equivalent to a conventional continuous gas phase hydroprocessing unit can be achieved with much smaller pill sizes. This smaller pill size has two significant effects. The small pill size in continuous liquid phase reaction conditions creates equivalent pressure drop, which will facilitate flow distribution, as is typical for hydroprocessing in a conventional continuous gas phase. The small pill size also reduces diffusional resistance and potentially can enhance hydroprocessing reaction rates. The small pill size can be used under continuous liquid phase hydroprocessing conditions with or without hydrogen recycle.

The methods described herein are particularly useful for hydroprocessing a hydrocarbonaceous feedstock containing hydrocarbons, and typically other organic materials, to produce a product containing hydrocarbons or other organic materials of lower average boiling point, lower average molecular weight, as well as reduced concentrations of contaminants, such as sulfur and nitrogen and the like. In an aspect, the process utilizes an initial hydrogen addition that provides all the hydrogen requirements for the reactor without the use of hydrogen sourced from a hydrogen recycle gas compressor. In other words, the hydrogen is not recycled within the hydroprocessing unit, but is supplied from outside the hydroprocessing unit.

As used herein, the term “communication” means that material flow is operatively permitted between enumerated components. The term “downstream communication” means that at least a portion of material flowing to the subject in downstream communication may operatively flow from the object with which it communicates. The term “upstream communication” means that at least a portion of the material flowing from the subject in upstream communication may operatively flow to the object with which it communicates.

The hydrocarbonaceous feedstocks that may be processed using the methods and apparatuses comprise mineral oils and synthetic oils (e.g., shale oil, tar sand products, etc.) and fractions thereof that may be subjected to hydroprocessing and hydrocracking. Illustrative hydrocarbon feedstocks include those containing components boiling above about 150° C. (300° F.), such as atmospheric gas oils, vacuum gas oils, vacuum and atmospheric residua, hydrotreated or mildly hydrocracked residual oils, coker distillates, straight run distillates, solvent-deasphalted oils, pyrolysis-derived oils, high boiling synthetic oils, cycle oils, catalytic cracker distillates, and Fischer-Tropsch derived liquids. One preferred feedstock is a gas oil or other hydrocarbon fraction having at least about 50 wt-%, and preferably at least about 75 wt-%, of its components boiling at a temperature above about 371° C. (700° F.). For example, another preferred feedstock contains hydrocarbon components which boil above about 288° C. (550° F.) with at least about 25 percent by volume of the components boiling between about 315° C. (600° F.) and about 565° C. (1050° F.). Other suitable feedstocks may have a greater or lesser proportion of components boiling in this range.

With reference to the FIGURE, an integrated hydroprocessing unit 10 is illustrated where a hydrocarbonaceous feedstock is introduced perhaps by pump into the process via a fresh hydrocarbonaceous feed lines 12 and 14. The hydrocarbonaceous feedstock is provided at a first temperature which may be a temperature well below reactor temperature such as between about 200° and about 300° F. (90° and 150° C.) because the feedstock has yet been subjected to heating.

A hydrogen-rich gaseous stream is provided via a hydrogen lines 20 and 22 via a make-up gas compressor 24. In an aspect, hydrogen in line 22 is only provided via a make-up gas compressor 24. The hydrogen supply line 20 which may perhaps be a line from a general refinery hydrogen supply provides hydrogen to hydrogen line 22. The hydrogen-rich gaseous stream from line 22 is admixed with the fresh feed in the hydrocarbonaceous feed line 14 to provide an admixture of the hydrocarbonaceous feedstock and hydrogen in line 16. The feed is heated to the appropriate reaction temperature with a heater either upstream of the joinder with the hydrogen line 22 in line 14 (not shown) or downstream thereof in line 16. The heater 18 may be one or more fired heaters and/or heat exchangers represented by fired heater 18. Alternatively or additionally, the hydrogen in line 22 may be heated by a heat exchanger 26 or other means and mixed with the fresh feed to thereby heat the hydrocarbonaceous feedstock from line 14 to the appropriate reaction temperature.

The heated, mixed stream in line 28 is introduced into a hydroprocessing reactor 40 via an inlet 42. The hydroprocessing reactor contains at least one bed 44 of hydroprocessing catalyst which in an aspect may be a fixed bed of catalyst. The hydroprocessing reactor 40 may have at least a single catalyst bed 44 and may have a plurality of catalyst beds. As mentioned above, the hydroprocessing reactor 40 is designed to be operated in a continuous liquid phase with the hydrogen requirement supplied from the combined stream of hydrogen from line 22.

The hydrocarbonaceous feedstock is subjected to hydroprocessing in a continuous liquid phase in the hydroprocessing reactor 40. Hydroprocessing can include, without limit, hydrotreating such as hydrodesulfurization, hydrocracking and hydroisomerization. Continuous liquid phase hydroprocessing involves introducing a liquid phase hydrocarbonaceous feedstock and hydrogen into a hydroprocessing reactor. The hydrogen should be present in a sufficiently low concentration to maintain a continuous liquid phase in the hydroprocessing reactor but high enough to provide sufficient hydrogen for hydroconversion of the hydrocarbon feed. In other words, a continuous plenum of hydrocarbon liquid should extend from the feed inlet 42 for the reactor 40 to the product outlet 46 for the reactor to establish a continuous liquid phase. Hydrogen gas may be present outside of the liquid plenum or inside of the liquid plenum in the forms of slugs or bubbles. At the very least, the volume of the liquid in the reactor will be greater than the volume of the gas in the reactor.

During the hydroconversion reactions occurring in the hydroprocessing reactor, hydrogen is necessarily consumed. Hydrogen may be provided to the reactor in excess or replaced by one or more hydrogen inlet points located in the reactor. The flow rate of hydrogen added at these locations is controlled to ensure that the reactor operates in a continuous liquid phase. The maximum flow rate of hydrogen that may be added to the reactor is less than the flow rate which would cause a transition from a continuous liquid phase to a continuous vapor phase.

In some aspects, the hydrocarbonaceous feedstock does not contain recycled product from a hydroprocessing reactor or other hydrocarbon diluent. In other aspects, a recycle stream 64 or diluent may be incorporated into the fresh hydrocarbonaceous feedstock prior to hydroprocessing the feedstock to provide additional volume to the process zone to provide added hydrogen-carrying capacity to the product stream or to provide additional mass to reduce the temperature rise in catalyst bed 44. In such aspects, any recycled product 64 or diluent typically is introduced into the feedstock in line 14 before a hydrogen stream in line 22 is mixed with the feedstock, and no further recycled product is incorporated into the process flow. Typically, such recycled product may be previously stripped of a vaporous hydrogen sulfide, nitrogen or nitrogen containing compositions, and any other vapor phase materials. In an aspect, recycled product in line 64 is optionally recycled to help carry hydrogen into the hydroprocessing reactor 40 via lines 16 and 28 or reduce the temperature rise in catalyst bed 44.

In one aspect, the fresh hydrocarbonaceous feed in line 14 is provided and mixed with a hydrogen flow in line 22 from a make-up gas compressor or other similar hydrogen sources. The hydrogen flow is mixed into the fresh hydrocarbonaceous feed for the hydroprocessing reactor 40 and is provided at a rate at least sufficient to satisfy the hydrogen requirement of the first reactor and subsequent reactors if any. If multiple catalyst beds are provided in a reactor, the hydrogen rate may be at least sufficient to satisfy the hydrogen requirement of the first catalyst bed and subsequent catalyst beds in the same or subsequent reactors if any. In some instances, the rate of added hydrogen will include an amount in excess of the predicted hydrogen requirements of the particular catalyst bed(s) as reserve in event the hydrogen consumption exceeds the expected rate at a particular bed or in the reactor(s) as a whole.

In other aspects, hydrogen is added to the fresh feed stream to provide sufficient hydrogen to exceed the saturation point of the hydrocarbonaceous liquid so that a small vapor phase is present throughout the substantially liquid phase. Thus, there is, in some aspects, sufficient additional hydrogen in the small vapor phase to provide additional dissolved hydrogen to the continuous liquid hydrocarbon phase as the reaction consumes hydrogen. For example, the amount of added hydrogen to each catalyst bed may be about 10 to 20 wt-% greater than the expected collective hydrogen requirements of each bed of hydroprocessing catalyst. In yet other aspects, it is expected that the amount of hydrogen may be up to about 100 percent of saturation to about 1000 percent of the saturated liquid phase hydrocarbon. Excess hydrogen is carried in the effluent from the catalyst bed and/or reactor in solution, in gaseous phase, or both in gaseous phase and in solution in the liquid effluent streams. In this aspect, no other hydrogen is added to the catalyst bed or reactor. In other aspects, hydrogen may be added to a downstream catalyst bed or a downstream reactor that supplements hydrogen exiting the upstream catalyst bed or upstream reactor, respectively. It will be appreciated, however, that the rate of hydrogen added to a reactor or a catalyst bed can vary depending on the feed composition, operating conditions, desired output, and other factors. In an aspect, the liquid hydrocarbonaceous feed may include about 67 to about 135 Nm³ hydrogen per m³ oil (400 to 800 scf/bbl). In this aspect, a continuous gas phase may exist along with the continuous liquid phase extending from hydrogen inlet to product outlet. As such, about 4 to about 25 Nm³ hydrogen per m³ oil (25 to 150 scf/bbl) may exit a respective catalyst bed or reactor at its outlet.

The “hydroprocessing” that may be performed in the hydroprocessing reactor may be “hydrotreating”. Hydrotreating is a process wherein hydrogen gas is contacted with hydrocarbon in the presence of suitable catalysts which are primarily active for the removal of heteroatoms, such as sulfur and nitrogen from the hydrocarbon feedstock. In hydrotreating, hydrocarbons with double and triple bonds may be saturated. Suitable hydrotreating catalysts for use in the present invention are any known conventional hydrotreating catalysts and include those which are comprised of at least one Group VIII metal, preferably iron, cobalt and nickel, more preferably cobalt and/or nickel and at least one Group VI metal, preferably molybdenum and tungsten, on a high surface area support material, preferably alumina. Other suitable hydrotreating catalysts include zeolitic catalysts, as well as noble metal catalysts where the noble metal is selected from palladium and platinum. It is within the scope of the present invention that more than one type of hydrotreating catalyst be used in the same reaction vessel. The Group VIII metal is typically present in an amount ranging from about 2 to about 20 wt-%, preferably from about 4 to about 12 wt-%. The Group VI metal will typically be present in an amount ranging from about 1 to about 25 wt-%, preferably from about 2 to about 25 wt-%.

Preferred hydrotreating reaction conditions include a temperature from about 204° C. (400° F.) to about 482° C. (900° F.), a pressure from about 3.5 MPa (500 psig) to about 17.3 MPa (2500 psig), a liquid hourly space velocity of the fresh hydrocarbonaceous feedstock from about 0.1 hr⁻¹ to about 10 hr⁻¹ with a hydrotreating catalyst or a combination of hydrotreating catalysts.

The “hydroprocessing” that may be performed in the hydroprocessing reactor may be “hydrocracking”. Hydrocracking refers to a process in which hydrocarbons crack in the presence of hydrogen to lower molecular weight hydrocarbons. Depending on the desired output, the hydrocracking zone may contain one or more beds of the same or different catalyst. In one aspect, for example, when the preferred products are middle distillates, the preferred hydrocracking catalysts utilize amorphous bases or low-level zeolite bases combined with one or more Group VIII or Group VIB metal hydrogenating components. In another aspect, when the preferred products are in the gasoline boiling range, the hydrocracking zone contains a catalyst which comprises, in general, any crystalline zeolite cracking base upon which is deposited a minor proportion of a Group VIII metal hydrogenating component. Additional hydrogenating components may be selected from Group VIB for incorporation with the zeolite base.

The zeolite cracking bases are sometimes referred to in the art as molecular sieves and are usually composed of silica, alumina and one or more exchangeable cations such as sodium, magnesium, calcium, rare earth metals, etc. They are further characterized by crystal pores of relatively uniform diameter between about 4 and about 14 Angstroms (10⁻¹⁰ meters). It is preferred to employ zeolites having a relatively high silica/alumina mole ratio between about 3 and about 12. Suitable zeolites found in nature include, for example, mordenite, stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite and faujasite. Suitable synthetic zeolites include, for example, the B, X, Y and L crystal types, e.g., synthetic faujasite and mordenite. The preferred zeolites are those having crystal pore diameters between about 8-12 Angstroms (10⁻¹⁰ meters), wherein the silica/alumina mole ratio is about 4 to 6. One example of a zeolite falling in the preferred group is synthetic Y molecular sieve.

The natural occurring zeolites are normally found in a sodium form, an alkaline earth metal form, or mixed forms. The synthetic zeolites are nearly always prepared first in the sodium form. In any case, for use as a cracking base it is preferred that most or all of the original zeolitic monovalent metals be ion-exchanged with a polyvalent metal and/or with an ammonium salt followed by heating to decompose the ammonium ions associated with the zeolite, leaving in their place hydrogen ions and/or exchange sites which have actually been decationized by further removal of water. Hydrogen or “decationized” Y zeolites of this nature are more particularly described in U.S. Pat. No. 3,130,006.

Mixed polyvalent metal-hydrogen zeolites may be prepared by ion-exchanging first with an ammonium salt, then partially back exchanging with a polyvalent metal salt and then calcining. In some cases, as in the case of synthetic mordenite, the hydrogen forms can be prepared by direct acid treatment of the alkali metal zeolites. In one aspect, the preferred cracking bases are those which are at least about 10 percent, and preferably at least about 20 percent, metal-cation-deficient, based on the initial ion-exchange capacity. In another aspect, a desirable and stable class of zeolites is one wherein at least about 20 percent of the ion exchange capacity is satisfied by hydrogen ions.

The active metals employed in the preferred hydrocracking catalysts of the present invention as hydrogenation components are those of Group VIII, i.e., iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum. In addition to these metals, other promoters may also be employed in conjunction therewith, including the metals of Group VIB, e.g., molybdenum and tungsten. The amount of hydrogenating metal in the catalyst can vary within wide ranges. Broadly speaking, any amount between about 0.05 percent and about 30 percent by weight may be used. In the case of the noble metals, it is normally preferred to use about 0.05 to about 2 wt-%.

The method for incorporating the hydrogenating metal is to contact the zeolite base material with an aqueous solution of a suitable compound of the desired metal wherein the metal is present in a cationic form. Following addition of the selected hydrogenating metal or metals, the resulting catalyst powder is then filtered, dried, pelleted with added lubricants, binders or the like if desired, and calcined in air at temperatures of, e.g., about 371° to about 648° C. (about 700°to about 1200° F.) in order to activate the catalyst and decompose ammonium ions. Alternatively, the zeolite component may first be pelleted, followed by the addition of the hydrogenating component and activation by calcining.

The foregoing catalysts may be employed in undiluted form, or the powdered zeolite catalyst may be mixed and copelleted with other relatively less active catalysts, diluents or binders such as alumina, silica gel, silica-alumina cogels, activated clays and the like in proportions ranging between about 5 and about 90 wt-%. These diluents may be employed as such or they may contain a minor proportion of an added hydrogenating metal such as a Group VIB and/or Group VIII metal. Additional metal promoted hydrocracking catalysts may also be utilized in the process of the present invention which comprises, for example, aluminophosphate molecular sieves, crystalline chromosilicates and other crystalline silicates. Crystalline chromosilicates are more fully described in U.S. Pat. No. 4,363,718.

By one approach, the hydrocracking conditions may include a temperature from about 232° C. (450° F.) to about 468° C. (875° F.), a pressure from about 3.5 MPa (500 psig) to about 16.5 MPa (2400 psig) and a liquid hourly space velocity (LHSV) from about 0.1 to about 30 hr⁻¹. In some aspects, the hydrocracking reaction provides conversion of the hydrocarbons in the process stream to lower boiling products, which may be the conversion of at least about 5 vol-% of the process flow. In other aspects, the per pass conversion in the hydrocracking zone may be in the range from about 15 percent to about 70 percent and, preferably, the per-pass conversion is in the range from about 20 percent to about 60 percent. In such aspects, the processes herein are suitable for the production of naphtha, diesel or any other desired lower boiling hydrocarbons.

The “hydroprocessing” that may be performed in the hydroprocessing reactor may be “hydroisomerization”. Hydroisomerization may also include catalytic dewaxing. Hydroisomerization is a process in which in one aspect at least about 10 percent, in another aspect, at least about 50 percent and, in yet another aspect, about 10 to about 90 percent of the n-paraffins of the hydrocarbon feedstock are converted into iso-paraffins effective to provide an effluent with at least one of a cloud point value of about 0° C. (32° F.) or less, a pour point value of about 0° C. (32° F.) or less, and/or a cold filter plugging point (CFPP) value of about 0° C. (32° F.) or less. In general, such hydroisomerization conditions include a temperature from about 260° C. (500° F.) to about 371° C. (700° F.), a pressure from about 1.38 MPa (200 psig) to about 8.27 MPa (1200 psig), a liquid hourly space velocity of the fresh hydrocarbon feedstock from about 0.1 hr⁻¹ to about 10 hr⁻¹. However, other hydroisomerization conditions are also possible depending on the particular feedstocks being treated, the compositions of the feedstocks, desired effluent compositions, and other factors.

Suitable hydroisomerization catalysts are any known conventional hydroisomerization catalysts. For example, suitable catalysts can include zeolite components, hydrogenation/dehydrogenation components, and/or acidic components. In some forms, the catalysts can include at least one Group VIII metal such as a noble metal (i.e., platinum or palladium). In other forms, the catalyst may also include silico alumino phosphate and/or zeolite alumino silicate. Examples of suitable catalysts are disclosed in U.S. Pat. No. 5,976,351; U.S. Pat. No. 4,960,504; U.S. Pat. No. 4,788,378; U.S. Pat. No. 4,683,214; U.S. Pat. No. 4,501,926 and U.S. Pat. No. 4,419,220; however, other isomerization catalysts may also be used depending on the feedstock composition, operating conditions, desired output, and other factors.

In an aspect of the invention, the hydroprocessing reactor 40 houses a hydroprocessing catalyst bed 44 containing hydroprocessing catalyst. In an aspect, the catalyst bed is fixed, such that the catalyst particles do not leave the bed with the exiting hydrocarbon. The hydroprocessing catalyst comprises pills that have a largest dimension that averages no more than 1.27 mm ( 1/20 inch) and more than 100 nm to produce a first hydrocarbonaceous product stream. In a further aspect, the hydroprocessing catalyst pills have a largest dimension that averages no more than 0.85 mm ( 1/30 inch) and no less than 0.51 mm ( 1/50 inch). In an even further aspect, the hydroprocessing catalyst pills have a largest dimension that averages no more than 0.51 mm ( 1/50 inch) and more than 100 nm. The small catalyst pills are useable in continuous liquid phase hydroprocessing because the overall mass flux may be lower when the gas-to-liquid ratio is such that a liquid continuous phase is formed as described above and because the substantially lower gas rates have a lower gas velocity over the small catalyst pills, thus generating a smaller pressure drop through the bed than a continuous gas phase operation would. However, pressure drop also serves a distributional purpose, so some pressure drop is desired. The smaller catalyst pills also serve to distribute the feed across the catalyst bed by virtue of axial dispersion over that which would be provided by conventionally large catalyst particles used for continuous gas phase hydroprocessing. Lastly, the small catalyst pills provide more surface area and a shorter diffusion path into the pores of the smaller catalyst pills, thereby enhancing hydroconversion to products.

The catalysts of the present invention may take any shape such as extruded trilobes or quadralobes, pellets, oil dropped spheres, layered spheres, spray dried particles, etc. without limitation.

We have found that with smaller catalyst particles, the hydroprocessing reactor can be operated with a higher mass flux. Conventional hydroprocessing reactors are rarely operated to approach a mass flux of 29,300 kg/h-m² (6,000 lb/h-ft²). We have found that the hydroprocessing reactor can be operated with a mass flux of more than 29,300 kg/h-m² with the catalyst pill sizes of the present invention. We have found that the hydroprocessing reactor can be operated at a mass flux that can reach 50,000 kg/h-m² (10,000 lb/h-ft²).

A hydrocarbonaceous product stream exits from the hydroprocessing reactor 40 via an outlet 46 in a hydroprocessed effluent line 48. The hydrocarbonaceous product may be recovered as product or further processed.

In an aspect, at least a portion of the hydrocarbonaceous product may be fed to a subsequent hydroprocessing reactor 50 via inlet 52. In this aspect, hydroprocessing reactor 40 is a first hydroprocessing reactor. The subsequent hydroprocessing reactor 50 contains at least one bed 54 of hydroprocessing catalyst which in an aspect is a fixed bed of catalyst. The hydroprocessing reactor 50 may have at least a single catalyst bed 54 and may have a plurality of catalyst beds. The hydroprocessing reactor 50 may be operated in a continuous liquid phase with the hydrogen requirement supplied from unreacted hydrogen exiting from the first hydroprocessing reactor 40 in line 48. In such a case, a large excess of hydrogen will be added to the first hydroprocessing reactor via lines 22, 16 and 28. In another aspect, hydrogen may be added to the subsequent hydroprocessing reactor 50 by bypassing the first hydroprocessing reactor via lines 30, 68 and 32.

The hydroprocessing that occurs in the subsequent hydroprocessing reactor 50 may be different from the hydroprocessing that occurs in the first hydroprocessing reactor 40. For example, hydrotreating may be conducted in the first hydroprocessing reactor 40 and either hydrocracking or hydroisomerization may be conducted in the subsequent hydroprocessing reactor 50. In another example, hydrocracking may be conducted in the first hydroprocessing reactor 40 and either hydrotreating or hydroisomerization may be conducted in the subsequent hydroprocessing reactor 50. In a further example, hydroisomerization may be conducted in the first hydroprocessing reactor 40 and either hydrotreating or hydrocracking may be conducted in the subsequent hydroprocessing reactor 50. Moreover, the same type of hydroprocessing may be conducted in both the first hydroprocessing reactor 40 and the subsequent hydroprocessing reactor 50. Lastly, more than one subsequent reactor 50 may be used in an aspect using the same or different type of hydroprocessing than the other hydroprocessing reactors.

In an aspect, the hydrocarbonaceous feedstock in line 12 may be separated into a first portion of fresh feed in a first hydrocarbonaceous portion line 14 and at least one additional portion of fresh feed in a second hydrocarbonaceous portion line 60. The first portion of fresh feed is fed in lines 16 and 28 into the first hydroprocessing reactor 40 and at least one portion of the resulting first hydrocarbonaceous product stream in line 48 is mixed with the additional portion of hydrocarbonaceous feed in the second hydrocarbonaceous portion line 60 via lines 62 and 68 and the mixture is fed via line 32 into the subsequent hydroprocessing reactor 50 through inlet 52. The additional portion of fresh feed in second hydrocarbonaceous portion line 60 serves to quench the first hydrocarbonaceous product stream in line 48 and regulate temperature before entering the subsequent hydroprocessing reactor 50.

The subsequent hydroprocessing reactor 50 may have a single catalyst bed 54 or may have a plurality of catalyst beds. The subsequent hydroprocessing reactor 50 may be operated in a continuous liquid phase or a continuous gas phase. The continuous liquid phase is preferred with the hydrogen requirement supplied from the mixed stream of unreacted hydrogen from the first hydroprocessing reactor 40 in line 48 and in an aspect hydrogen also optionally provided from lines 30, 68 and 32. The hydrogen concentration should be kept sufficiently low to maintain a continuous liquid phase if desired but high enough to provide sufficient hydrogen for hydroconversion of the hydrocarbon feed. A subsequent hydrocarbonaceous product stream exits the subsequent hydroprocessing reactor 50 through outlet 56 into line 34.

In an aspect of the invention, the subsequent hydroprocessing reactor may house a fixed catalyst bed and contain hydroprocessing catalyst comprising pills that have a largest dimension that averages no more than 1.27 mm ( 1/20 inch) and more than 100 nm to produce the subsequent hydrocarbonaceous product stream. In a further aspect, the hydroprocessing catalyst pills in the subsequent hydroprocessing reactor 50 have a largest dimension that averages no more than 0.85 mm ( 1/30 inch) and no less than 0.51 mm ( 1/50 inch). In an even further aspect, the hydroprocessing catalyst pills in the subsequent hydroprocessing reactor 50 have a largest dimension that averages no more than 0.51 mm ( 1/50 inch) and more than 100 nm. The hydroprocessing catalyst pills in the subsequent hydroprocessing reactor 50 need not have the same general size as in the first hydroprocessing reactor 40. The subsequent hydroprocessing reactor may also be operated with a higher mass flux of more than 29,300 kg/h-m² (6,000 lb/h-ft²) when loaded with catalyst having sizes according to the present invention. We have also found the mass flux can reach 50,000 kg/h-m² (10,000 lb/h-ft²) if operated with the catalyst having sizes of the present invention.

The first hydrocarbonaceous product stream in line 48 or the subsequent hydrocarbonaceous product stream in line 34 may be transported into a separation zone 70. A vaporous stream is removed from the separation zone 70 via line 72 and may be further separated into a hydrogen rich stream, contaminants, such as hydrogen sulfide and ammonia, and low boiling point hydrocarbons. In this case, stream 72 may be cooled, low boiling point hydrocarbons separated in a flash drum with the ensuing gas being scrubbed with an aqueous solution of absorbent such as an amine. The resulting hydrogen-rich stream may be sent to a general refinery hydrogen supply, but is not directly recycled back to the hydroprocessing reactor 40 or 50 unless optionally recycled through the make-up gas compressor 24 after mingling with the general refinery hydrogen supply. Consequently, the hydrogen line 22 is out of downstream communication with the hydroprocessing reactors 40 or 50 but optionally through a make-up gas compressor 24. The remaining liquid product is removed from the separation zone via line 74 and is directed in line 78 to further processing or to a fractionation zone for further separation into its constituents.

An alternative embodiment is shown in phantom in the FIGURE. The remaining liquid phase is removed from the separation zone via line 74 and, optionally, a portion of the liquid phase is externally recycled in line 76, such that the external recycle is added as a diluent as desired to one or all of the streams of fresh hydrocarbonaceous feed 14 or 60 through lines 64 or 66, respectively. In another aspect, the external recycle is added as a diluent entirely to the first portion of fresh feed 14 in line 64. The remaining liquid phase from the separation zone 70 is directed by line 78 to further processing treatments and/or to a fractionation zone for further separation into its constituents.

Product recycle in line 64 may be mixed with the first portion of hydrocarbonaceous feed in line 14 and hydrogen from line 22 and is carried in lines 16 and 28 into the first hydroprocessing reactor 40. Product recycle in line 66 may be mixed with the second portion of hydrocarbonaceous feed in line 60 and transported in line 62. Optional hydrogen from line 30 may be mixed into line 62 and the mixture may be carried in line 68 to be mixed with first hydrocarbonaceous effluent from line 48 and the mixture fed to the subsequent hydroprocessing reactor 50 in line 32. In an aspect, the hydrogen in line 30 is not preheated before mixing with the first hydrocarbonaceous effluent in line 48.

In an aspect in which no subsequent hydroprocessing reactor 50 is present, the hydroprocessed product in line 48 may be delivered to the separator 70. Product recycle may be fed in lines 76 and 64 just to the hydroprocessing reactor 40 via lines 16 and 28.

Without further elaboration, it is believed that one skilled in the art can, using the preceding description, utilize the present invention to its fullest extent. The preceding preferred specific embodiments are, therefore, to be construed as merely illustrative, and not limitative of the remainder of the disclosure in any way whatsoever.

In the foregoing, all temperatures are set forth in degrees Celsius and, all parts and percentages are by weight, unless otherwise indicated.

From the foregoing description, one skilled in the art can easily ascertain the essential characteristics of this invention and, without departing from the spirit and scope thereof, can make various changes and modifications of the invention to adapt it to various usages and conditions.

Claims

1. A process for hydroprocessing a hydrocarbonaceous feedstock which comprises: introducing a liquid phase stream comprising a hydrocarbonaceous feedstock and a sufficiently low hydrogen concentration to maintain a continuous liquid phase into a hydroprocessing reactor containing hydroprocessing catalyst comprising pills that have a largest dimension that averages no more than 1.27 mm ( 1/20 inch) and more than 100 nm to produce a first hydrocarbonaceous product stream.

2. The process of claim 1 wherein the pills have a largest dimension that averages no more than 0.85 mm ( 1/30 inch) and no less than 0.51 mm ( 1/50 inch).

3. The process of claim 1 wherein the pills have a largest dimension that averages no more than 0.51 mm ( 1/50 inch).

4. The process of claim 1 further comprising: introducing at least a portion of the first hydrocarbonaceous product stream into a subsequent hydroprocessing reactor containing hydroprocessing catalyst with a sufficiently low hydrogen concentration to maintain a continuous liquid phase to produce a subsequent hydrocarbonaceous product stream.

5. The process of claim 4 wherein said hydroprocessing catalyst in said subsequent hydroprocessing reactor comprises pills that have a largest dimension that averages no more than 1.27 mm ( 1/20 inch) and more than 100 nm.

6. The process of claim 5 further comprising mixing said first hydrocarbonaceous product stream with additional hydrocarbonaceous feedstock before entering into said subsequent hydroprocessing reactor.

7. The process of claim 1 wherein the hydroprocessing reactor is operated at a mass flux of more than 29,300 kg/h-m2 (6,000 lb/h-ft2).

8. The process of claim 1 wherein the hydroprocessing catalyst comprises a fixed bed.

9. A process for hydroprocessing a hydrocarbonaceous feedstock which comprises: introducing a liquid phase stream comprising a hydrocarbonaceous feedstock and a sufficiently low hydrogen concentration to maintain a continuous liquid phase into a hydroprocessing reactor containing hydroprocessing catalyst comprising pills that have a largest dimension that averages no more than 0.85 mm ( 1/30 inch) and no less than 0.51 mm ( 1/50 inch).

10. The process of claim 9 further comprising: introducing at least a portion of the first hydrocarbonaceous product stream into a subsequent hydroprocessing reactor containing hydroprocessing catalyst with a sufficiently low hydrogen concentration to maintain a continuous liquid phase to produce a subsequent hydrocarbonaceous product stream.

11. The process of claim 10 wherein said hydroprocessing catalyst in said subsequent hydroprocessing reactor comprises pills that have a largest dimension that averages no more than 0.85 mm ( 1/30 inch) and no less than 0.51 mm ( 1/50 inch).

12. The process of claim 10 further comprising mixing said first hydrocarbonaceous product stream with additional hydrocarbonaceous feedstock before entering into said subsequent hydroprocessing reactor.

13. The process of claim 9 wherein the hydroprocessing reactor is operated at a mass flux of more than 29,300 kg/h-m2 (6,000 lb/h-ft2).

14. The process of claim 9 wherein the hydroprocessing catalyst comprises a fixed bed.

15. A process for hydroprocessing a hydrocarbonaceous feedstock which comprises:

introducing a stream comprising a hydrocarbonaceous feedstock and hydrogen into a hydroprocessing reactor containing a fixed bed of hydroprocessing catalyst comprising pills having a largest dimension that averages no more than 0.51 mm ( 1/50 inch).

16. The process of claim 15 further comprising: introducing at least a portion of the first hydrocarbonaceous product stream into a subsequent hydroprocessing reactor containing hydroprocessing catalyst with a sufficiently low hydrogen concentration to maintain a continuous liquid phase to produce a subsequent hydrocarbonaceous product stream.

17. The process of claim 16 wherein said hydroprocessing catalyst in said subsequent hydroprocessing reactor comprises pills that have a largest dimension that averages more than 100 nm.

18. The process of claim 16 further comprising mixing said first hydrocarbonaceous product stream with additional hydrocarbonaceous feedstock before entering into said subsequent hydroprocessing reactor.

19. The process of claim 15 wherein the hydroprocessing reactor is operated at a mass flux of more than 29,300 kg/h-m2 (6,000 lb/h-ft2).