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Publication numberUS3159564 A
Publication typeGrant
Publication dateDec 1, 1964
Filing dateOct 20, 1961
Priority dateOct 20, 1961
Publication numberUS 3159564 A, US 3159564A, US-A-3159564, US3159564 A, US3159564A
InventorsInwood Texas V, Kelley Arnold E, Vaell Raoul P
Original AssigneeUnion Oil Co
Export CitationBiBTeX, EndNote, RefMan
External Links: USPTO, USPTO Assignment, Espacenet
Integral hydrofining-hydro-cracking process
US 3159564 A
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Description  (OCR text may contain errors)

Dec l, 1964 A. E. KELLEY r-:TAL

NTEGRAL HYDROFINING-HYDROCRACKING PROCESS Filed 001;. 20, 1961 LWN United States Patent O 3,l'9,564 WTEGRAL HYDRFMING-HYDRU- CRACFING PRDCESS` Arnold E. Kelley, Grange, Raoul P. `Vlaeli, Los Angeles,-

This invention relates to the catalytic hydrogenation and/or hydrocracking of hydrocarbonsto produce eg., hydrogenated aromatics, or lower boiling hydrocarbons, boiling for example in the gasoline or jet fuel range. In one aspect, the invention embraces an integral hydrolining-hydrocracking process, the hydroning treatment being integrated in such a manner that the bulk of the hydrofiner effluent boiling above the gasoline range can be sent directly, and essentially nitrogen-free, to the hydrocracking zone with a minimum of intervening cooling, depressuring, Washing, reheating and repressuring. The invention is particularly concerned with the treatment of high end-point feedstocks which are contaminated with organic nitrogen and/ or sulfur compounds.

Briey stated, the invention comprises the following essential steps: (l) the initial feedstock is subjected to f catalytic hydrofining with added hydrogen under conditions such that both ia liquid phase and a vapor phase are produced, each comprising a substantial portion of the initial feed; (2) the liquid phase, after substantial hydroning has takenl place, is stripped with hydrogen at essentially hydroning temperatures and pressures to remove ammonia, hydrogen sulfiide and light hydrocarbons; (3) the vapor phase4 effluent from the hydrofiner, together with the stripping vapors from step (2), are then cooled Sulliciently to condense out a substantial portion of the rcmaining hydrocarbons; (4) the resulting condensate is stripped with hydrogen at essentially hydrotining pressures to remove ammonia, hydrogen sulfide, and dissolved light hydrocarbons; (5) the stripped liquid phase from step (2) and the stripped condensate from step (4) yare then subjected to catalytic hydrogenation or hydrocracking substantially in the 'absence of catalyst poisons such as ammonia or hydrogen sulfide; and (6) the gasoline, or other desired hydrocarbon fraction synethesized in the hydrogenation zone, is recovered by condensation and fractionation.

According to a preferred embodiment of the invention, at least a portion of the stripping of liquid phase (step 2) is carried out integrally during the hydroiining by providing a countercurrent flow of hydrogen upwardly through the lower portion of the hydrof'lner to strip out ammonia and light hydrocarbons from the downowing liquid phase, the resulting stripping vapors being withdrawn from the hydroner at an upper point, along with the vapor phase hydrofiner eflluent. This preferred mode of operation may be carried out either (l) by admitting feed-plus-hydrogen at the top of the hydrofiner, admitting an additional hydrogen stream at the bottom, and removing vapor phase etl'luent at a mid-point, or (2) by admitting feed-plushydrogen at a mid-point to the hydrofiner, admitting an additional hydrogen stream at the bottom, and removing vapor phase effluent from the top. In either .case the liquid phase gravitates downwardly in the lower portion of the reactor, countercurrently to the 'hydrogen admitted at the bottom. The invention also embraces several other modifications which will become apparent as the description proceeds.

A major difficulty in hydrocracking processes is encountered in connection with feedstocks containing organic nitrogen compounds. These nitrogen compounds are mostly basic in character, and tend to poison the 3,159,564 Patented Dec. l, l9lfl` acidic cracking centers of hydrocracking catalysts. To overcome this difficulty, Vthe feedstock is often subjected to a preliminary catalytic hydrofining treatment in order to convert the nitrogen compounds to ammonia, and the product is then .condensed and washed to remove the ammonia. The remaining nitrogen-free hydrocarbons are then subjected to hydrocracking. The principal objection of this pretreatment process is that the facilities required for condensing, washing, reheating and repressuring the hydrofiner eflluent, `are Very expensive.

It is also known in catalytic hydrocracking that the poisoning effect of nitrogen compounds can be minimized by operating at higher hydrocracking temperatures. However, this procedure is objectionable in the case of feedstocks containing substantial amounts of high-boiling components, boiling for example above about 650 F. High-temperature hydrocracking of feedstocks containing such high-boiling constituents leads to rapid deactivation of the catalyst due to coke formation.

In the catalytic hydrogenation of aromatic hydrocarbons, many catalysts, and especially the more active ones such as platinum and other noble metals, are sensitive to poisoning by sulfur compounds. Noble metal hydrocracking .catalysts may also be sulfur-sensitive in some cases. It is therefore desir-able in many cases to remove sulfur compounds, as Well as nitrogen, from hydrogenation or hydrocracking feedstocksi This may be achieved quite efficiently by catalytic hydrofining, but subject to the same disadvantages previously noted, involving expensive interstage treatments to remove hydrogen sulfide.

It is therefore, the overall objective of this invention to provide a convenient and and economical method for removing nitrogen and/ or sulfur compounds from hydrogen-ation or hydrocracking feedstocks. A more specific objective is to obtain all of the substantial benefits of prehydroning without the normally accruing disadvantage of the expensive condensation, washing, reheating and repressuring facilities required for separate, or nonintegral hydrofining. A still further objective is to provide hydrofining facilities which will effect a maximum removal of nitrogen and sulfur compounds Without overtreatment of the lower boiling portions, thus minimizing the overall volume of hydroning catalyst and reactor space required. Other objectives will be apparent from the more detailed description which follows.

The invention may perhaps bel more readily understood with reference to the accompanying drawing, which is a flowsheet illustrati-ng several modifications. In the succeeding description, it will be understood that the drawing has been simplified by the omission of certain conventional elements such as valves, pumps, compressors, and the like. Where heaters or coolers are indicated, it will be understood that these are merely symbolic, and in actual practice many of these will be combined into banks of heat exchangers and fired heaters, according to standard engineering practice.

Referring more particularly to FIGURE 1, the initial feedstock is brought in via line 2, mixed with recycle hydrogen from line 4, preheated to incipient hydrofining temperature in heater 6, and then passed directly into the top of hydrofiner 3, vwhich contains an upper bed of hydroiining catalyst l0 where mixed-phase hydrofining proceeds under substantially conventional, concurrent downflow conditions. The mixed-phase effluent from upper hydrofining catalyst bed l@ emerges into an inter- -space l2, where the initial phase separation takes place.

Product vapors, comprising vapor phase effluent from bed 8 and vapor-phase stripping effluent from lower catalyst bed 14, are withdrawn via line 16 and treated as hereinafter described. The liquid phase portion of efl'luent from bed 8 gravitates downwardly through lower catn) alyst bed 14, countercurrently to a stream of hydrogen admitted at the bottom of the reactor via line i8.

1n the modification illustrated, a small optional stripping column 20, containing a series of stripping trays 21, is provided below catalyst bed 14, through which the liquid portion of feed percolates countercurrently to the hydrogen stream and collects in the bottom, forming a liquid seal. The purpose of the stripping column is to assure that ammonia generated in the lower portion of catalyst bed 14 will also be removed. For efficient stripping, the hydrogen admitted via line 1S should preferably be preheated to approximately the temperature prevailing in the reactor, as by means of heater 22. The ratio of hydrogen to oil in catalyst bed 14 and stripping column preferably ranges between about 500 and 15,000 scf. per barred of oil, at reactor pressures between about 400 and 2,000 p.s.i.g. `and temperatures between about 400 and 800 F.

The stripped heavy oil accumulating in the bottom of stripping column 20 is taken off via line 24 and valve 26, in response to liquid level controller 28, and is transferred directly to hydrocracker 30, via lines 32. and 34 and heat exchanger 36, as will be subsequently described.

Suitable hydroining catalysts for use in hydrofiner 8 include for example mixtures of the oxides and/ or sulfides of cobalt and molybdenum, or of nichel and tungsten, preferably supported on a carrier such as alumina, or alumina containing a small amount of coprecipitated silica gel. Other suitable catalysts include in general the oxides and/or sulfides of the Group VIB and/or Group VIH metals, preferably supported on substantially noncracking adsorbent oxide carriers such as alumina, silica, titania, and the like. The hydrofining operation may be conducted either adiabatically or isothermally, and under the following general conditions:

HYDROFINING CONDITIONS The above conditions are suitably adjusted so as to reduce the nitrogen content of the feed to below about parts per million, and preferably below about 10 parts per million.

The vapor phase which is taken olf from hydrofiner 8 via line 16 comprises hydrogen saturated with hydrocarbons of all boiling range, and also containing ammonia and hydrogen sulfide. The light hydrocarbons, boiling in and below the gasoline range, comprise mostly hydrocarbon fragments derived from the decomposition of sulfur and nitrogen compounds in hydrofiner S, since very little true hydrocracking occurs therein. It is desired to recover at minimum expense, and in an ammoniafree condition, the bulk ofthe hydrocarbons boiling above the gasoline range for treatment in hydrocracker 30. For this purpose, the 'total vapor phase in line 16 is transferred via cooler 38 to a small separator-stripping column 40. At reactor pressures, it is normally found that the bulk `of these heavier hydrocarbons can be condensed out of the vapor phase by cooling the same to about 400- 600 F. It is undesirable to cool the Vapor any more than is required to produce a substantial desired amount of liquid condensate; generally it is unnecessary to go below about 300 F. It will be understood that stripping column 40 is operated at substantially the same pressure prevailing in the hydrofiner, or within about 100 p.s.i.g. thereof.

The cooled mixed-phase material entering the top of separator-stripper 40 separates in the upper portion thereof, the vapor phase going overhead via line 42, and the liquid portion gravitating downwardly over stripping trays 44, countercurrently to a rising stream of preheated fihydrogen admitted near the bottom thereof via line 46, thereby stripping light hydrocarbons and ammonia therefrom. Hydrogen rates similar to those prescribed for stripper 20 may be employed, and the stripping hydrogen is preheated in heater 48 to about the same temperature as :the entering mixed-phase efiiuent from cooler 38.

The liquid phase accumulating in the bottom of stripper 44 will of course contain some high-boiling hydrocarbons, but will consist primarily of light gas oil boiling in t. e 400 to 600 F. range. This liquid phase is withdrawn via line 50 and valve 52 in response to liquid level controller 54, and is then transferred via line 56 to hydrocraclter 30, where it is subjected to hydrocracking under one of two alternate modifications hereinafter described.

The vapor phase overhead from stripper i4 comprises mostly gasoline-boiling-range hydrocarbons, but will still contain a considerable proportion of light gas oil hydrocarbons. lt is also heavily contaminated with ammonia and/ or hydrogen sulfide. To effect recovery and purification of the various components, the total vapor phase in line 42 is partially cooled and condensed in cooler 5S, then mixed with wash water from line 60, and transferred via final condenser 62 to high pressure separator 64. The purpose of the wash water is to remove water-soluble impurities, including :unmom'a and hydrogen sulfide and salts thereof. Normzlly about 15 pounds of water per barrel of liquid hydrocarbon is sufficient for this purpose. It is preferable to inject the wash water prior to the final cooling in condenser 62, in order to avoid condensation of solid salts in the transfer lines, as might occur if final cooling to, eg., 200 F. were carried out before the injection of water.

A three-phase separation takes place in high pressure separator 6ft. Spent 'ash water containing dissolved impurities is withdrawn via line 66, and purified recycle hydro-gen is withdrawn via line 6d for reuse in the hydroner and stripping columns associated therewith, as previously described. Liquid hydrocarbons in separator 64 are withdrawn via line 70 and flashed into low pressure separator 72, -rom which light hydrocarbon gases are exhausted via line The liquid hydrocarbon fraction in separator 72 is transferred via line 76 to fractionating column 73, from which the desired gasoline product is taken overhead via line S0.

y As will become apparent from the succeeding description, the hydrocarbon fraction admitted to separator 72 via line 70 also includes the hydrocarbon product from hydrocracker 30. Hence, fractionating column 78 serves to recover both the minor proportion of gasoline synthesized in the hydroner, and the major portion synthesized in hydrocrackcr 30. If desired, a side-cut jet fuel or diesel fraction may be recovered from the column via line 82. The bottoms fraction from column 7 8, comprising unconverted gas oil hydrocarbons from hydrocracker 30, and remaining gas oil hydrocarbons from hydrofiner d, is taken` off via line 34 and blended with the liquid phase hydroner effluent in line 32, and the mixture is then transferred via line 34 and exchanger' 36 to hydrocracker 30. Alternatively, where stripper 116 is utilized as described below, it may be preferable to mix the bottoms fraction in line 84 with the stripped condensate in line 56, and divert the mixture Via preheatcr 1M) into the mid-section of hydrocracker 30, as will be described below in connection with alternate B'. ft should be noted further that the bottoms fraction from column 78 nec/:l not necessarily be recycled to the hydrocracker; if conversion of the feedstock is not desired, all or a portion of the bottoms fraction in line 84 may be diverted from the process for other uses, e.g., as feed to a catalytic cracking unit.

The operation of hydrocracker 30 will now be described in reference to the two principal alternate modifications referred to above.

1 Alternate A.-In this modification, the entire feed to the nydrocracker enters the top and iiows downwardly therethrough in a conventional manner. This is the simplest modification, and may be preferred in cases where the light gas oil feed derived from stripper 40 constitutes a Very minor portion of the fresh feed to the hydrocracker. To operate in this manner, valves 103 and 112 are closed, and the stripped condensate in line 56 is transferred via open valve S6 and line 8S to line 34, which contains the main feed stream derived from stripper Ztl and recycle oil line 84. The mixture is blended with recycle and fresh hydrogen derived from lines 9G and 92, and the temperature is suitably adjusted in exchanger 36, which may either be a heater or a cooler, depending upon the temperature of the feed mixture in line 34, as well as upon the desired hydrocracking temperature, which in turn will depend upon the relative activity of the hydrocracking catalyst. Normally however, since the hydrogen mixed with the feed isV relatively cool, some preheating will be required. in any case, the mixture of hydrogen and oil, at suitable inlet hydrocracking temperatures is admitted via line 94 into the top of hydrocracking reactor 34B and passes downwardly therethrough under` suitable hydrocracking conditions hereinafter prescribed.

Hydrocracker 3i) may ,be operated under either adiabatic or isothermal conditions, and in the latter case, the reaction being exotherrnic, it may be desirable to admit cool quench hydrogen at one or more points in the reactor a-s illustrated via lines 96 and 98. It will be noted that the catalyst in hydrocracker is divided into an upper bed 11N? and a lower bed 1152, separated by an interspace 1114. This has no particular significance in reff erence to alternate A, land hence in this case a single unitary catalyst bed may be utilized. Effluent from the hydrocraclier is taken off via line 106 and treated for product recovery as hereinafter described.

Alternate B.-Since the liquid phase hydrofiner effluent from stripper 2b has a considerably higher average molecular weight than the strippedcondensate from stripper 40, it may be advantageous in some cases to take advantage oi the different optimum cracking temperatures for these two fractions by introducing them separately into the reactor in zones where optimum temperatures prevail for the respective fraction. This` is particularly feasible in cases where the initial feed to the hydrofiner contains substantial proportions, e.g., more than labout 20% by volume, of hydrocarbons boiling in the 400- 600 F. range, in which case the stripped condensate in line 56 will comprise a substantial proportion of the total feed to the hydrccracker. Selective treatment of the two feed fractions is effected in this alternate by maintaining upper hydrocracking bed 1111i at arelatively lower average temperature than lower hydrocracking bed 192, and feeding the heavy fraction of feed into the top of the reactor while feeding the lighter fraction into the top of the lower hydrocracking catalyst bed. The differential temperatures required in this mode of operation can in most cases be maintained by allowing hydrocracking to proceed in reactor 3@ under more nearly adiabatic conditions, with a minimum of quenching via lines 96 and 98. It is contemplated that upper hydrocracking bed 161i may be maintained at an average bed temperature about 2O 100 F. lower than the average bed temperature in lower hydrocraclring bed 192. Preferably, upper bed 108 is maintained at an average temperature between about 500 and 750 F., and lower bed at 162 at Stiff-800 F.

To effect selective treatment of the feed fractions in this manner, valve 86 is closed and valve 1% is opened,

thereby diverting stripped condensate in line S6 through preheater 110 into interspace 104, where it mingles with the downflowing effluent from upper bed 11W. It is also desirable in this case to mix the light feed with hydrogen via line 111 and open valve 112. By operating in the manner prescribed, upper hydrocracking bed 113i) serves primarily to crack the heavy gas oil feed to light gas oil, while lower bed 162 serves primarily to convert light gas oil components to gasoline. Maximum catalyst efficiency, and efficiency of conversion to gasoline is obtained.

In either of the above alternate modes of operation, the hydrocracker effluent in line 106 is next passed through a partial condenser 114 and thence into an optional separator-stripper 116, which functions similarly to stripping column 40. The purpose of separator-stripper 116 is to provide for economical recovery of a substantial proportion of unconverted oil from the hydrocracker before it is cooled to final condensation temperatures, thus minimizing the utilities and facilities required for re-heating the recycle oil. Normally, the temperature of the hydrocracker eiiiuent is reduced in cooler 114 to about 400700 F., and about 500-l0,000 s.c.f. of hydrogen per barrel of liquid condensate is admitted near the bottom of the stripper via preheater 118 and line 12th, to effect countercurrent shipping of light hydrocarbons from the condensate. The condensate accumulating in the bottom of stripper 116 comprises the heaviest components of the unconverted oil, and is more or less continuously removed via line 122 and valve 124, in response to liquid level controller 126, and recycled to the hydrocracker via lines 128 and 34.

The vapor phase from separator-stripper 116 is taken overhead via line 130, and comprises the uncondensed vapor phase effluent from the hydrocracker as well as the stripping vapors from stripper 116. This combined vapor phase is then transferred via nal condenser 130 into high pressure separator 132, from which hydrogenrich recycle gas is taken olf via line 134 and reused in the process as previously described. The condensed hydrocarbons in separator 132 are withdrawn via line 136 and blended with the liquid hydrocarbons in line 70, and the resulting mixture is flashed into low pressure separator 72 and fractionated in column 78 as previously described.

As indicated above, the use of stripper-separator 116 is optional, and in many cases, especially in small installations, may not be economically justifiable. In these cases, it is preferable simply to omit the stripper, and send the hydrocracking effluent in line 106 directly through final condenser 130 and high pressure separator 132. Where the stripper is employed however, it will be apparent that the recycle bottoms fraction recovered from column 78 will have a relatively lower average molecular weight than the recycle condensate from stripper 116. Its boiling range may approximate that of the condensate from stripper 40; hence the feasibility of combining these fractions for selective treatment in lower hydrocracking bed 102 as noted above.

Many variations are contemplated from the processing scheme described above and in the drawing. For example, as shown in FIGURE 2, it is entirely feasible to operate hydrofiner` 3 with the total feed plus hydrogen entering interspace 12 via line 140, the vapor phase passing upwardly through bed 1li and the liquid phase gravitating downwardly through lower bed 14, countercurrently to the rising hydrogen stream admitted via line 18. The total Vapor phase effluent is then taken off at the top of the hydrofiner via line 142 and sent to separatorstripper d@ via cooler 38. Liquid phase product is taken off as before, via line 24. This modification is advantageous in providing for a more efficient hydrofining of the vapor phase in the absence of liquid phase.

The process conditions in hydrocracker 3@ are suitably adjusted so as to provide about 30-8{)% conversion to gasoline, or other desired product, per pass. It is preferred at the same time to adjust the operating conditions, preferably temperature, so Vas to permit relatively long runs between regenerations, i.e., from about 2-8 months. For these purposes, it will be understood that pressures in the high ranges will'be used in connection with temperatures in the high range, while the lower operative pressures will normally be used in conjunction with the 4 lower temperatures. The range of operative conditions contemplated for, hydrocracker are as follows:

HYDROCRACKING CONDTONS Those skilled in the art will readily understand that when ranges of operating conditions are specified as above, a large number of determinative factors are involved. Thus, highly active catalysts, or fresh catalysts at the beginning of a run, will be used in conjunction with lower temperatures than will less active or partially deactivated catalysts. Also, the lower temperature ranges will normally be used in conjunction with feedstocks having high end-points, i.e., above about 750 F. The lower limit of pressure to be utilized in a given operation will normally depend upon the desired run length. Lower pressures generally result in a more rapid deactivation of the catalyst, and hence where extremely long run lengths are desired, pressures of above about 1,000 p.s.i.g. are mandatory. However, economically feasible run lengths are normally obtainable with most catalysts and feedstocks within the 6002,000 p.s.i.g. pressure range.

The hydrocracking feedstocks which may be treated herein include in general any mineral oil fraction having an initial boiling point above the conventional gasoline range, i.e., above about 400 F., and having an endboiling-point of up to about 1,000 F. This includes straight-run gas oils, coker distillate gas oils, deasphalted crude oils, cycle oils derived from catalytic or thermal cracking operations and the like. These fractions may be derived from petroleum crude oils, shale oils, tar sand oils, coal hydrogenation products and the like. Specifically, it is preferred to employ feedstocks boiling between about 400 and 900 F., having an API gravity of 15 to and containing at least about 20% by volumn of of acid-soluble components (aromatics-l-olefins). Such oils may also contain from about 0.1% to 5% of sulfur and from about 0.01% to 2% by weight of nitrogen.

Hydrogenation feedstocks may comprise any of the foregoing hydrocracking feeds, and in addition many other aromatic concentrates as eg., naphthalene or methyl naphthalene fractions.

It should be noted further that the process of this invention is especially amenable to the treatment of feedstocks boiling predominantly in the 550900 P. range. Where the feedstock is composed mainly of hydrocar bons boiling in the 400600 F. range, the alternate two-stage hydrocracking process disclosed in co-pending application Serial No. 137,366, tiled September 11, 1961, may be more desirable, assuming that the size of the operation justifies the capital inventment in a two-stage process. The process of this invention is primarily directed to the treatment of feedstocks which (l) contain a substantial proportion of heavy materials boiling above about 650 F., and (2) contain less than about 30% by volume of light components in the 400-600 F. range.

The hydrocracking catalyst to be employed in the hy- Y drocracking unit described above may consist of any desired combination of a refractory cracking base with a suitable hydrogenating component. Suitable cracking bases include for example mixtures of two or more refractory oxides such as silica-alumina, silica-magnesia, silica-zirconia, alumina-boria, silica-titania, silica-zirconiatitania, acid treated clays and the like. Acidic metal phosphates such as aluminum phosphate may also be used. The preferred cracking bases comprise composites of silica and alumina containing about 50%-90% silica; coprecipitated composites of silica, titania and zirconia containing between 5% and 75% of each component; partially dehydrated, zeolitic, crystalline molecular sieves, e.g., of the X of Y crystal types, having relatively uniform pore diameters of about 8 to 14 Angstroms, and comprising silica, alumina and one or more exchangeable zeolitic cations.

Any of the foregoing cracking bases may be further promoted by the addition of small amounts, eg., l to 10% by weight, of halides such as tluorine, boron triiiuoride or silicon tetratluoride.

The foregoing cracking bases are compounded, as by impregnation, with from about 0.5% to 25% (based on free metal) of a Group Vil?. or Group VIII metal promoter, eg., an oxide or sulfide of chromium, tungsten, cobalt, nickel, or the corresponding free metals, or any combination thereof. Alternatively, even smaller proportions, between about 0.05% and 2% of the metals platinum, palladium, rhodium or irridium may be employed. The oxides and suldes of other transitional metals may also be used, but to less advantage than the foregoing.

A particularly suitable class of hydrocracking catalysts is composed of about 75 %95 by weight of a coprecipitated base containing 5%-75% Si02 5%-75% ZrO2, and 5 %75 TiOZ, and incorporated therein from about 5 %*25 based on free metal, of a Group Vlll metal or metal sulfide, e.g., nickel or nickel sulfide.

The molecular sieve type cracking bases, when compounded with a hydrogenating metal, are particularly useful for hydrocracking at relatively low temperatures of 40G-700 F., and relatively low pressures of '400-1,500 p.s.i.g. It is prefered to employ molecular sieves having a relatively high SiO2/Al203 ratio, eg., between about 2.5 and 6.0. The most active forms are those wherein the exchangeable zeolitic cations are hydrogen and/or a polyvalent (preferably divalent) metal such as magnesium, calcium or zinc.A In particular, the Y molecular sieves, wherein the SiO2/Al203 ratio is about 5, are preferred, either in their hydrogen form, or a divalent metal form, preferably magnesium.

The Y molecular sieves can be prepared by heating an aqueous sodium alumino-silicate mixture at temperatures between about 25 and 125 C. (preferably 80-125 C.) until crystals are formed, and separating the crystals from the mother liquor. When a colloidal silica sol is employed as the source of silica, the aqueous sodium alumina-silicate mixture may have a composition as follows, expressed in terms of mole-ratios:

NazO/SOZ 0.2-0.8 SiO2/Al203 10-30 H2O/N212O 25-60 When sodium silicate is used as the silica source, the optimum molar proportions are as follows:

Na2O/Si02 0.6-2.0 SiO2/Al2O3 10-30 HZO/NaZO 30-90 The decationized, or hydrogen form of the Y zeolite may be prepared by ion-exchanging the alkali metal cations with ammonium ions, or other easily decomposable cations such as methyl subsittuted quaternary ammonium ions, and then heating to eg., 300-400 C., to drive off the ammonia. It is preferred that the degree of decationization, or hydrogen exchange, be at least about 40% Aof the maximum theoretically possible. The final composition should contain less than about 5% by weight 0f Nago.

The polyvalent metal (e.g., Mg, Ca, Sr, Ba, Zn, Mn, Ni, Cr, etc.) forms of the Y zeolite are prepared by conventional cation exchange methods, as e.g., by treating the sodium zeolite with an aqueous solution of magnesium chloride, calcium chloride, etc. It is preferred to replace at least about 40% of the monovalent metal cations with their equivalent of polyvalent cations; the activity of the catalyst generally incre-ases as the degree of displacement of monovalent metals increases.

Mixed, hydrogen-polyvalent metal forms of the Y zeolite are also contemplated. Generally such mixed forms are prepared by subjecting a polyvalent metal-exchanged zeolite to further ion exchange with ammonium cations which 'are later driven olf during a thermal activation treatment. Here again, it is preferred that at least about 40% of the monovalent metal cations be replaced with any desired combination of hydrogen ions and polyvalent metal cations.

The hydrogenating promoter may be incorporated into the molecular sieve type cracking bases by any method which gives a suitably intimate admixture. Among acceptable methods are (l) cation exchange using an aqueous solution of a suitable metal salt wherein the metal itself forms the cation; (2) cation exchange using an aqueous solution of a suitable metal compound in which the metal is in the form of a complex cation with coordination complexing agents such as ammonia, followed by thermal decomposition of t-he cationic complex; (3) conventional impregnation with an aqueous solution of a suitable metal salt, followed by drying and thermal decomposition of the metal compound. The ion-exchange methods (l) and (2) are much to be preferred in that a more uniform and complete subdivision of the metal on the zeolite is obtained.

Method (l) above is general-ly employed to introduce metals of the iron group, while method (2) is generally best adapted for the noble metals of Group VIII. When method (l) is employed to introduce an iron group metal, it is desirable to carry out subsequent thermal activation treatments in a nonoxidizing or reducing atmosphere in order to avoid oxidizing the metal and'displacing it from the zeolite lattice. But in the case of the Group VIH noble metals such precautions are generally unnecessary, and thermal decomposition of the cationic complex can be carried out in air if desired.

The ion-exchange of hydrogenating metal onto the zeolite may be carried out by the usual methods described above in connection with the ion-exchange of polyvalent metals generally. In fact, the two steps may be combined if desired by using an aqueous solution of a mixture of salts, or a single salt in cases where the desired polyvalent metal is also the desired hydrogenating metal. Generally, however, itis preferable to carry out first the ion-exchange step for forming the hydrogen and/or polyvalent metal zeolite, and then perform the ion-exchange step providing the hydrogenating metal.

As in the case of the X molecular sieves, the Y sieves also contain pores of relatively uniform diameter in the individual crystals. In the case of X sieves, the pore diameters may range between about 6 yand 12 A., and this is likewise the case in the Y sieves, although the latter usually are found to have crystal pores of `about 8 to l0 A. in diameter.

Non-cracking hydrogenation catalysts for use herein may comprise any of the previously noted transitional metal hydrogenating components, preferably supported on a neutral, adsorbent carrier having an extended surface area of e.g., 50-300 square meters per gram. Suitable carriers include for example, activated alumina, silica gel, activated charcoal, clays and the like. Hydrogenating catalysts comprising a noble metal such as platinum, palladium or rhodium are especially preferred since they are highly active and most susceptible to poisoning by sulfur compounds. Hydrogenating conditions are Ain general similar to the hydrocracking conditions above described, except that temperatures in the lower range of about 400-600 F. are generally preferred.

The following example is cited to illustrate more concretely, exemplary process conditions and results, but is not intended to be limiting.

Example This example illustrates the results obtainable in a typical operation of the process described in the drawing, with the liquid phase from the hydroner being treated serially in two superimposed beds of hydrocracking catalyst, and the liquid condensate from stripper 40 being treated only in the `second of the two beds. The feed is a 400-850 F. boiling range coker gas-oil containing 2% of sulfur and 0.3% of nitrogen by weight, and having an API gravity of 22. The hydrofining catalyst is composed of about 4% cobalt sulfide plus 16% of molybdenum sulfide impregnated on a silica-stabilized (5% SiO2) activated alumina support. The hydrocracking catalyst is a magnesium Y molecular sieve containing about 3% by weight of zeolitic magnesium, and loaded with 0.5% by weight of palladium (Linde hydrocracking catalyst MB 53 82). The operating conditions are as follows:

TABLE 1 Hydrofiner 8 Hydrocracker 30 Operating Conditions Stripper Upper Lower 40 Upper Lower bed 10 bed 14 bed bed 102 LHSV 1. 5 2. 0 2. 0 2. 0 Pressure, p.s.i.g 1, 550 1, 550 1 540 1, 530 1, 530 Temperature, Ik Start of run 700 700 500 450 520 750 500 670 700 Hz/oil s.c.f./b 3, 000 3, 000 8, 000 8, 000

Both beds Conversion to 400 F. endpoint gasoline, vol. per cent per pass -3 -2 65 H Average bed tempera-tures.

On the basis of an operation utilizing 10,000 barrels per day of initial feed, and with total recycle of unconverted oil to the hydrocnacker, the approximate yields are as follows.

In lthe foregoing example, about 30% by volume of the total feed to the vhydroliner is recovered as condensate from separator-stripper 40. If this condensate is blended with the stripped liquid phase from the hydroner, and the mixture then subjected `to hydrocracking at 1.0 space velocity, and an isothermal bed temperature adjusted to give 65% Conversion per pass to gasoline, similar product distribution and yields are obtained, except for somewhat higher dry gas and humane makes. The run length is also somewhat decreased as a result of a more rapid rate of catalyst deactivation.

Results analogous to those indicated in the foregoing example are obtained when other hydrocracking catalysts and conditions, other feedstocks and other hydroning conditions within the broad purview of the above disclosure are employed. It is hence not intended to limit the invention to the details of the example or the drawing, but only broadly as defined in the following claims.

We claim:

1. In a hydrogenation process wherein a hydrocarbon feedstock, a substantial portion of which boils above about 550 F., and which contains an organic impurity from the class consisting of sulfur compounds and nitrogen compounds, is rst subjected to hydroiining over a sulfactive hydrofining catalyst at elevated pressures and temperatures such that a substantial portion of feed in the hydroiiner is in the liquid phase and a substantial portion is in the vapor phase, and wherein a selected, purified, heavy fraction of the resulting hydrofiner effluent is then subjected to a catalytic hydrogenating reaction at elevated temperatures and pressures over a hydrogenation catalyst which is sensitive to poisoning by said impurity, the improved method for recovering and purifying said selected heavy fraction from the hydrofiner effluent with a mimmum of interstage cooling, depressuring, washing, reheating and repressuring, which comprises:

(l) subjecting said liquid phase, after substantial hydrofining has taken place, to stripping with hydrogen at essentially hydrofining temperatures and pressures to strip out dissolved volatile decomposition products of said impurity;

(2) cooling vapor phase efiiuent'from said hydrofining in fadmixture with stripping vapors from step (l) to a temperature substantially below its initial temperature without substantial depressuring, to effect a partial condensation of liquid hydrocarbons therefrom;

(3) stripping the liquid condensate from step (2) with hydrogen without substantial depressuring to remove dissolved volatile decomposition products of said impurity; and

(4) combining stripped liquid phase efliuent from step (l) with stripped condensate from step (3) to form said selected puinified heavy fraction.

2. A process `as defined in claim 1 wherein remaining vapor phase hydrofiner effluent from step (2) is further cooled and condensed and washed with water to remove decomposition products of said impurity, and .wherein a high-boiling `fraction of the resulting washed condensate is combined with said seiected purified heavy fraction.

3. A process as defined in claim 1 wherein said volatile decomposition products comprise ammonia, and wherein said hydrogenating catalyst is a hydrocracking catalyst comprising a hydrogenating metal supported on a solid, adsorbent, acidic cracking base having an extended surface area.

4. A process as defined in claim 1 wherein said volatile decomposition products comprise hydrogen sulfide, and wherein said hydrogenating catalyst comprises a Group VIII noble metal supported on a solid, adsorbent base having an extended surface area.

5. A process as defined in claim 1 wherein said feedstock contains a substantial proportion of hydrocarbons boiling above about 650 F., and less than about 30% by volume of hydrocarbons boiling below about 600 F.

6. A process as defined in claim l wherein said stripping of liquid phase hydrofining product is carried out during hydrofining by passing said liquid phase downwardly countercurrently to a rising stream of hydrogen in at least a portion of said hydrofining catalyst.

7. A process .as defined in claim l wherein said hydrofining and at least a portion of said stripping of liquid phase hydrofining product are effected simultaneously by (1) passing initial feedstock plus hydrogen downwardly through a first bed of hydrofining catalyst; (2) effecting a liquid phase-vapor phase separation of efiiuentfrom said first catalyst bed under hydrofinring conditions of pressure and temperature; and (3) passing the liquid phase from said separation' downwardly through a second bed of hydrofining catalyst countercurrently to a rising stre-am of hydrogen, thereby effecting simultaneous hydrofining and stripping of said liquid phrase.


8. A process =as defined in claim 1 wherein said hydrofining and at least aportion of said stripping of liquid phase lhydrofining product are effected simultaneously by (l) subjecting said initial feedstock plus hydrogen to a preliminary liquid-vapor phase separation under substantially hydrofining conditions of temperature and pressure; (2) passing the resulting vapor phase through a first bed of hydroning catalyst; and (3) passing the resulting liquid phase from said separation downwardly through a second bed of hydrofining catalyst counter-currently to a rising stream of hydrogen, thereby effecting simultaneous hydrofining and stripping of said liquid phase.

9. A process for hydrccracking a hydrocarbon feedstock initially contaminated with organic nitrogen compounds, and containing a substantial proportion of hydrocarbons boiling above yabout 550 F., which comprises, (l) subjecting said feedstock to hydrofining with added hydrogen over a sulfactive hydrofining catalyst under hydrofining conditions such that a substantial portion of the feed is in the liquid phase 'and a substantial portion is in the vapor phase; (2) subjecting said liquid phrase after substantial hydrofining has taken place to stripping with hydrogen at essentially hydrofining ternperatures and pressures to strip out dissolved ammonia; (3) cooling vapor phase effluent from said hydrofining in yadmitture with stripping vapors from step (2) to a tempcrature substantially below its initial temperature without substantial depressuring, to effect a partial condensation of liquid hydrocarbons therefrom; (4) stripping thc liquid condensate from step (3) with hydrogen without substantial depressuring to remove dissolved ammonia; (5) subjecting said stripped liquid phase hydrofiner effluent plus added hydrogen to hydrocracking lin a first bed of hydrocrajcking catalyst at an average bed temperature T1; (6) blending efiiuent from said first hydrocracking catalyst bed with said stripped condensate; and (7) subjecting the resulting blend to hydrocrackng in a second bed of hydrocracking catalyst at Aan average bed temperature T2 which is substantially higher than T1.

10. A process as defined in claim 9 wherein T1 is between about 500 and 750 F. and T2 is between about 550 and 800 F l1. In a hydrocnackng process wherein a wide-boilingrange hydrocarbon feedstock is subjected to catalytic hydrocracking in a hydrocracking zone in which relatively lhigh temperatures prevail in a downstream section and relatively low temperatures prevail in an upstream section, the improvement which comprises dividing said feedstock into a relatively low-boiling fraction and a relatively highboiling fraction, passing said high-boiling fraction plus hydrogen serially through said low-temperature section then through `said high-temperature section, and passing said low-boiling fraction only through said high-temperature section in admixture with efiiuent from said low-temperature section.

12. A process as defined in claim 11 wherein said hydrocracking catalyst comprises a Group VIII metal hydrogenating component deposited upon a zeolitic, aluminosilicate molecular sieve cracking base.

Goretta et al Nov. 3, 1959 Kelley etal Sept. 13, 1960

Patent Citations
Cited PatentFiling datePublication dateApplicantTitle
US2911352 *Oct 31, 1957Nov 3, 1959Standard Oil CoProcess for manufacture of high octane naphthas
US2952626 *Aug 5, 1957Sep 13, 1960Union Oil CoMixed-phase hydrofining of hydrocarbon oils
Referenced by
Citing PatentFiling datePublication dateApplicantTitle
US3242229 *Jan 13, 1965Mar 22, 1966Texaco IncHydrocarbon conversion process
US3248316 *May 1, 1963Apr 26, 1966Standard Oil CoCombination process of hydrocracking and isomerization of hydrocarbons with the addition of olefins in the isomerization zone
US3260663 *Jul 15, 1963Jul 12, 1966Union Oil CoMulti-stage hydrocracking process
US3268438 *Apr 29, 1965Aug 23, 1966Chevron ResHydrodenitrification of oil with countercurrent hydrogen
US3318802 *Jan 21, 1965May 9, 1967Exxon Research Engineering CoHydrocracking process employing a crystalline alumino-silicate activated with a chlorine compound
US3354078 *Feb 4, 1965Nov 21, 1967Mobil Oil CorpCatalytic conversion with a crystalline aluminosilicate activated with a metallic halide
US3364132 *Sep 19, 1966Jan 16, 1968Universal Oil Prod CoHydrocarbon conversion process to produce gasoline from high boiling hydrocarbon oils by hydrocracking and reforming
US3461061 *Jun 13, 1966Aug 12, 1969Universal Oil Prod CoHydrogenation process
US3912620 *Mar 20, 1974Oct 14, 1975Atlantic Richfield CoLubricating oil production utilizing hydrogen in two catalytic stages
US4040944 *Feb 22, 1973Aug 9, 1977Union Oil Company Of CaliforniaGroup 6b andor group 8 metal(s) hydrogenation catalyst on aluminum silicate zeolite; high-octane gasoline
US4058449 *May 21, 1976Nov 15, 1977Institut Francais Du PetroleGroup 6b or 8 metal containing catalyst
US4338186 *Nov 17, 1980Jul 6, 1982Suntech, Inc.Hydrotreatment, hydrocracking
US4565621 *Jan 9, 1984Jan 21, 1986Union Oil Company Of CaliforniaRegeneration
US4584287 *Apr 15, 1985Apr 22, 1986Union Oil Company Of CaliforniaRare earth-containing Y zeolite compositions
US4585751 *Jun 24, 1985Apr 29, 1986Phillips Petroleum CompanyHydrotreating catalysts
US4604187 *Jun 3, 1985Aug 5, 1986Union Oil Company Of CaliforniaIon exchanging sodium y zeolite with rare earth oxides and group eight oxide
US4680105 *May 5, 1986Jul 14, 1987Phillips Petroleum CompanyHydrodemetallization of oils with catalysts comprising nickel phosphate and titanium phosphate
US4705768 *Jan 20, 1987Nov 10, 1987Phillips Petroleum CompanyCoprecipitate of metal phosphates
US5279726 *Mar 21, 1991Jan 18, 1994Union Oil Company Of CaliforniaCatalyst containing zeolite beta and processes for its use
US5350501 *Feb 28, 1991Sep 27, 1994Union Oil Company Of CaliforniaA beta zeolite and a zeolite type Y for refining hydrocarbons
US5447623 *May 2, 1994Sep 5, 1995UopZeolite type beta and Y with hydrogenation component
US5536687 *May 31, 1994Jul 16, 1996UopCatalyst containing zeolite Beta
US6402935 *Nov 23, 1999Jun 11, 2002Uop LlcHydrocracking process
US6517705 *Mar 21, 2001Feb 11, 2003Uop LlcHydrocracking process for lube base oil production
US7001503 *Jan 14, 2000Feb 21, 2006Japan Energy CorporationMethod and apparatus for stripping sulfur-containing compounds from hydrocarbon feed stock in hydrorefining of petroleum distillates
US7749373Dec 13, 2005Jul 6, 2010Haldor Topsoe A/SHydrocracking process
U.S. Classification208/59, 208/216.00R, 208/254.00H, 208/2, 208/210, 208/89
International ClassificationC10G67/02, C10G65/00, C10G67/00, C10G65/12, C10G47/00
Cooperative ClassificationC10G47/00, C10G65/12, C10G67/02
European ClassificationC10G67/02, C10G65/12, C10G47/00