Catalytic conversion of hydrocarbon distillates

United States Patent Office _p. .. ²·⁹⁴⁴’⁰⁰⁵

Patented July 5, 1960 tionation and preferential hydrogenative cracking to produce high molal yields of lower-boiling cyclic structures and of lower-boiling paraffins of higher iso-normal ratio than required' by thermodynamic equilibrium. The production of lower-boiling paraffins whose iso-normal ratio is considerably in excess of the thermodynamic equilibrium is a material factor contributing to the unusual octane characteristics of the synthetic gasoline product and illustrates the material departure of the subject isocracking process from the conventional hydrocracking or reforming processes.

Accordingly, under these conditions of isocracking operation, a C5+ gasoline product fraction boiling below about 200° F. is produced which possesses a leaded F-l octane number of at least 97. The process is also characterized by the formation of a product of somewhat reduced aromatic content (if any), as compared with that of the feed and by a hydrogen consumption above about 1000 s.c.f. per barrel of feed converted to products boiling below the initial boiling point of the feed, the latter being designated herein as “synthetic” products.

References herein to boiling points or boiling point ranges, if not specifically described as TBP values, are those measured in accordance with ASTM distillation procedure D-86. It is furthermore to be understood that, in respect to the designations of fuel boiling ranges, a 10% by volume tolerance is to be permitted in order to more closely approximate the practical limitations of refinery distillation equipment and practices. Thus, the designation of a feed boiling range between 200° F. and 450° F. resolves itself to those feeds wherein at least the 10% and 90% distillation points fall within the stated range. Further, a product designated as boiling below 200° F. includes a product fraction wherein at least the 95% point is at 200° F. or below.

In carrying out the foregoing conversion operation, the effluent from the isbcracker may be worked up in any convenient fashion. Thus, a gas recycle stream rich in hydrogen is customarily separated in a high pressure gas-liquid separation zone, which stream is recycled to the isocracker along with make-up hydrogen. Thereafter, remaining normally gaseous products are separated in a gas-liquid separation zone operated at lower pressures, leaving a normally liquid effluent portion from which the synthetic product of the present invention (preferably a C₅+ gasoline fraction boiling below about 200° F.) can be recovered either as a separate stream or along with higher boiling product fractions. In either case, the process may properly be said to be productive in substantial yields of a C₆⁺ gasoline fraction boiling below about 200° F. which possesses a leaded F—1 octane fating of at least 97.

Again, in addition to recovering a preferred gasoline product fraction boiling below about 200° F., as aforementioned, it is also possible to recover from the isocracker effluent a cut boiling above said gasoline fraction which makes an excellent jet fuel component since it is extremely low in aromatics and exhibits outstanding thermal stability, it being substantially free of deposit-forming even when maintained for extended periods at temperatures of 600° F. and above. Alternatively, such higher boiling fractions may be recycled to the isocracking unit for conversion to lower boiling (e.g., below 200° F.) gasoline fractions of high octane value, or they may be diverted to other appropriate refinery uses.

As a general proposition, the charging stocks to the subject invention process may be any of the conventional hydrocarbon distillate fractions boiling in the range of from about 200°—450° F. whose nitrogen content has been reduced, where not already sufficiently low, to a basic nitrogen content of less than 25 p.p.m. through hydrofining or otherwise, and which further satisfy the specification

2,944,005

CATALYTIC CONVERSION OF HYDROCARBON . DISTILLATES ⁶

John W. Scott, Jr., Ross, Calif., assignor to California Research Corporation, San Francisco, Calif., a corporation of Delaware

No Drawing. Filed Aug. 13,1958, Ser. No. 754,715 ¹⁰

Claims. (Cl. 20S—109)

This invention relates to a process for the catalytic 15 conversion of hydrocarbon distillates to premium gasoline fractions, and more particularly those boiling essentially below the feed stocks employed, said product gasoline fractions being clean-burning and having a leaded F-l octane rating of at least 97. 20

More specifically, this invention is directed to the application of a unique conversion process on a hydrocarbon fraction of selected boiling range which is relatively free of aromatic components as well as of those compounds containing nitrogen. This conversion process 25 involves a critical correlation of process variables to attain a multiple combination of reactions in which those of isomerization, disproportionation and selective cracking predominate, and is conducted in the presence of hydrogen over a selective catalyst composition incor- 30 porating both hydrogenation as well as highly active cracking components at elevated pressures and at distinctively controlled temperatures resulting in a substantial consumption of hydrogen. In order to adequately differentiate and define the distinctive features 35 of the subject conversion process in comparison with conventional refining operations, this process will hereinafter be referred to as an isocracking process.

According to this invention, it has been found possible to utilize the distinctive features of the isocracking proc- ⁴θ ess for the production of a premium gasoline product fraction which is characterized by its clean-burning properties and by a leaded F-l octane rating of at least 97. The term “leaded,” as employed herein, designates those fuels Which contain 3 ml. tetraethyl lead (TEL) per 45 gallon. The instant process is also productive of a product fraction having particular utility as a jet fuel, it evidencing a high degree of thermal stability. However, this fraction may be readily converted in whole or in part to the aforesaid gasoline fraction, if desired, by recycling 50 it to the isocracking zone. The formation of these product fractions in optimum yields is predicated on a critical correlation of charging stock composition and process variables in the conduct of the isocracking process.

The process of the invention involves the selection of 55 a hydrocarbon fraction boiling in a range of from about 200° to 450° F. which contains not more than about 10% by volume of aromatic hydrocarbons and which has a nitrogen content, measured as basic nitrogen, of less than about 25 p.p.m. This selected charging stock is 60 introduced into an isocracking zone in admixture with at least 2000 s.c.f. hydrogen per barrel of feed for contact in said zone with a multifunctional catalyst composition comprising a hydrogenating component disposed on an active acidic cracking support. The isocracking zone 65 is operated under a pressure of at least about 200 p.s.i.g. and at catalyst temperatures falling in a range of about 450°-800° F.

By virture of the critical correlation of process variables associated with the aforementioned balance of cata- 70 lyst components, it has been found possible to effect simultaneous reactions involving isomerizatino, dispropor2,644,005 effect of nitrogen in the feed even when present in substantial amounts becomes progressively smaller as reaction temperatures increase above 800° F., and becomes substantially lost at temperatures above 850° F. Under the conditions encountered in the present isbcracking operations, on the other hand, and particularly at temperatures below about 750° F., the (basic) nitrogen compounds present in the feed or formed in the isocracking zone become chemisorbed upon the catalyst surfaces and thus drastically reduce the activity of the catalyst.

While reference is uniformly made to feeds having a maximum basic nitrogen content of 25 p.p.m., this assumes a total nitrogen content of less than about 100 p.p.m. since conventional refinery streams normally contain from about 3 to 4 times as much total nitrogen as basic nitrogen. However, in view of this known distribution of the nitrogen compounds, it is now common in the art to make only the basic nitrogen determination since this test is much easier and less expensive to make than is that for total nitrogen. It is also known that the non-basic nitrogen components present in the feed are quickly converted to basic nitrogen compounds once contact is made with the isocracking catalyst. Hence, when dealing with isocracker feed streams which contain only 25 basic nitrogen compounds or with those wherein the distribution between basic and nonbasic nitrogen compounds is other than that conventionally encountered, it is appropriate to define the isocracker feed as one which contains a total of less than about 100 p.p.m. of nitrogen.

The isocracking process of this invention is critically applicable to the processing of stocks of relatively low aromatic content, with specific reference to those boiling in the range of from about 200°-450° F., since such feeds are productive of C₅-200° F. gasoline fractions hav35 ing a leaded F-l octane rating of from about 97 to 100 or more. On the other hand, when the process is applied to a stock which, though otherwise meeting feed specifications, contains a substantial proportion of components boiling below 200° F., it is found that the resulting C5-200° F. gasoline fraction recovered from the isocracker effluent has a leaded F—1 octane number which is well below the minimal value of 97 otherwise obtained. At the other end of the scale, i.e., when applying the process of the invention to feeds boiling above about .-.- _ .. . . ..... ,.. _ gasoline product (boiling below the initial boiling point of the feed) also has a leaded octane number which is far below the desired rating of 97 or above.

The conversion or isocracking process to which the 50 foregoing charge stocks are subjected further involves a critical correlation of process variables to achieve the unique combination of conversion reactions which are manifested in the physical and chemical characteristics ______- _________________________ of the synthetic product. The principal process variables specification maximum higher reaction temperatures are 55 effecting the desired combination of reactions, and therenecessary to maintain an economic per-pass conversion level or, in other words, a per-pass production of synthetic product boiling below the initial boiling point of the feed. These requisite higher reaction temperatures entail a dis10 of having an aromatic content of not more than about 10% by volume. In this connection, the word “aromatic” is employed in the conventional sense to include all those hydrocarbons which incorporate at least one aromatic nucleus whether substituted or unsubstituted, and which may also contain other atoms such as nitrogen, oxygen and sulfur. The basic nitrogen content of the charging stock can be determined in the conventional manner by dissolving the stock in glacial acetic acid and titrating the solution with perchloric acid, also in glacial acetic acid, crystal violet commonly being employed as the indicator.

The feeds charged to the isocracker may be of petroleum origin or they may be obtained from gilsonite, shale, coal tar or other sources. For example, a suitable feed can be obtained by distillation from certain crude petroleum stocks, one such feed (from a Minas crude) being employed in the operation of Example I given below. Another suitable feed comprises a raffinate portion remaining after separating out the aromatic compounds present in the effluent stream (normally termed a “reformate”) obtained from a catalytic reforming unit. Reformates are conventionally produced by passing straight-run, thermally cracked and/or catalytically cracked naphthas, along with hydrogen, through a reforming unit provided with a platinum-on-alumina catalyst under reforming conditions. Separation of the aromatic from the nonaromatic portions of the product stream so obtained can be effected by a variety of means, as by the use of a glycol, sulfur dioxide, furfural or other appropriate selective solvent, or by the use of an adsorbent such as silica gel. Still other appropriate feeds for use in the present invention can be obtained by the practice of said separation methods in connection with other hydrocarbon.streams.

One of the important variables in the conduct of the subject precess which has a material effect, and to that extent permits the production of the desired gasoline products, is the control of the nitrogen content of the charge stock, as indicated, an acceptable nitrogen level, expressed as basic nitrogen, is about 25 p.p.m., although appreciable further improvement is obtained as this nitrogen content is reduced to levels below 10 p.p.m. These nitrogen levels may be reached by hydrofining the feed stock by treating the same with hydrogen at elevated temperatures and pressures in the presence of a hydrogenating catalyst . . _ . .

havingTittle cracking activity. The effluent from this pre- 45 450° F., it is found that the resulting synthetic treating or hydrofining step, if satisfying the specification ’ .—j--, ^zt-—’ 1 boiling ranges and aromatic content, may be fed directly to the isocracking stage of this invention or it may be first subjected to a preliminary fractionation to recover a specification feed.

In general, the effect of a basic nitrogen content in excess of 25 p.p.m. is a reduction in catalyst activity which is reflected in both operational efficiency and product distribution. As the nitrogen content increases above the the various process variables.

As previously indicated, the catalyst employed is a multifunctional catalyst composition comprising a hydro_x__________—o_____,________a____genating-dehydrogenating component disposed on an tures if conversion is to be maintained at the desired level, 65 active acidic cracking support. The cracking component or support may comprise any one or more of such siliceous, acidic materials as silica-alumina, silica-magnesia, silica-alumina-zirconia composites, as well as certain acid-treated clays and similar materials; provided, fore identifiable with the isocracking process, are catalyst composition and reaction temperature or temperature control. This is not intended to imply that other pro— ____________r________ ____________ cess variables are not significant to the isocracking proproportionate increase in the amount of product con- 60 cess, but rather to emphasize the degree of criticality in verted to light gases and carbonaceous residues which ”‘‘ form on the catalyst surface and thus decrease catalyst activity. Any such decrease in activity must be compensated for by resorting to still higher operating temperaand thus, the on-stream portion of the cycle is shortened as limiting temperatures of 800° F. and above are reached much sooner than would otherwise be the case. The criterion employed to determine the necessity of catalyst regeneration in a constant conversion operation is the 70 however, that such acidic materials possess substantial prescribed specification limit of isocracking at 800° F. '

This effect of nitrogen is in contrast to that observed in other hydrocracking operations which, though employing an intrinsically acidic catalyst system, are conducted at temperatures above 800° F. In such operations the 75 type of acidity in the finished catalyst composition.

cracking activity, it being recognized that in some cases the acidic nature of the cracking component may be enhanced, as by the addition of halides or the practice of other known means for developing Lewis or Bjronsted §,044,O&S . . .5

'.More specifically, the Cracking component of the catalyst is preferably one having an activity, in terms of gasoline production, of at least about 25 as measured by the “Cat. A” method (J. Alexander and H. G. Shimp, National Petroleum News (1944), vol. 36, at page R-537; J. Alexander, Proc. Am. Petroleum Inst. (1947), vol. 27, at page 51).

A preferred siliceous cracking support for the subject catalyst composition is comprised of synthetically prepared composites of silica and aluminum containing from about 75 to 90% of the silica component. A material of this type, in crushed aggregate form, was employed in the various exemplary runs described herein, said material containing about 87% silica, having a Cat. A activity of 46, and a surface area of about 430 m.²/g.

The hydrogenating-dehydrogenating component of the catalyst may be selected from any one or more of the various group VI and group VUI metals, as well as the oxides and sulfides thereof, representative materials being the oxides and sulfides of molybdenum, tungsten, chromium and the like, together with such metals as nickel or cobalt and the various oxides and sulfides thereof. Also suitable are certain group 1(B) or group 11(B) metals, such as copper or Cadminum and their oxides and sulfides. If desired, more than one hydrogenatingdehydrogenating component may be present, e.g., composites of two or more of the oxides and/or sulfides of molybdenum, cobalt, nickel, copper, chromium and zinc. - Depending on the activity thereof, the amount of the hydrogenating-dehydrogenating component may be varied Within relatively wide limits of from about 0.1 to 15%, based on the weight of the entire catalyst composition. Within these limits, the amount of said component present should be sufficient to provide a reasonable catalyst on-stream period at required conversion levels. As a general rule, exemplary catalysts having satisfactory activity are those containing from about 1 to 10% of one or more of the oxides and/or sulfides of molybdenum, nickel or cobalt, together with the hydrogen-reduced counterparts of said oxides, it being recognized that many of these oxides are reduced to the metal state and remain as such once the oxide has been exposed to hydrogen at elevated temperatures and pressures, either before the conversion reaction takes place or under the conditions prevailing in the isocracking unit as the same is placed on-stream.

Particularly good results from the standpoint of high per-pass conversion, even at relatively low operating temperatures, coupled with good selectivity and the ability to withstand repeated regeneration with relatively minor decrease in activity, are obtained with catalysts composed of from 1 to 10% nickel sulfide deposited on the aforementioned synthetically prepared silica-alumina composites, and these catalysts constitute a preferred class for use in this isocracking process.

The following catalysts are representative of those which are well adapted for use in the present invention. NICKEL SULFIDE (2.5% Ni) ON SILICA-ALUMINA

This catalyst (No. 425-2) was prepared by impregnating 11 liters of a crushed silica-alumina aggregate (87% SiO₂) with 2896.9 grams of Ni(NO₃)₂-6H₂O, dissolved in enough water to make 8800 milliliters total solution, following which the beads were held for 24 hours at 70° F. The catalyst was then dried for 10 hours at 250° F. and thereafter calcined at 1000° F. for 10 hours. The calcined material was reduced in an atmosphere of hydrogen at 580° F. and 1200 p.s.i.g., following which the resulting nickel-bearing catalyst was sulfided in an atmosphere containing 8% H₂S in hydrogen at 1200 p.s.i.g. and 580° F., thereby converting the nickel essentially to nickel sulfide.

NICKEL SULFIDE (2.5% Ni) ON SILICA-ALUMINA

This catalyst (No. 316) was prepared by impregnating 11 liters of a crushed silica-alumina aggregate (87% '6

SiO₂) with a solution prepared by mixing 1500 milliliters water and 500 milliliters of ammonium hydroxide solution with 1082 grams of ethylenediamine tetraacetic acid (EDTA) and 469 grams of nickel carbonate, the solution being made up to a total of 4000 milliliters with water. The impregnated material was held for a period of 24 hours at 70° F., following which it was centrifuged and calcined for 10 hours at 1000° F. in air to convert the nickel chelate to nickel oxide. The catalyst was then reduced in an atmosphere of hydrogen at 650° F. and 1200 p.s.i.g. and sulfided in situ in the reactor by the use of a feed stream made up of a catalytic cycle oil (49 volume percent aromatics) to which 0.1% by volume of dimethyl disulfide had been added, at a pressure of 1200 p.s.i.g., and in the presence of approximately 6500 s.c.f. H₂ per barrel of feed

NICKEL SULFIDE (25% Ni) ON SILICA-ALUMINA

This catalyst (No. 353) was prepared by impregnating approximately 7.5 liters of a crushed SiO₂ (87% )A1₂O₃ aggregate which had been dried in air for 24 hours at 400° F., with 2183.7 grams of Ni(NO₃)₂-6H₂O dissolved in water and made up to a total of 7760 milliliters. The impregnated base material was then held for 24 hours at 70° F. and calcined for 10 hours at 1000° F. The catalyst was then sufided by treatment in an atmosphere of hydrogen containing 8% hydrogen sulfide at 1200 p.s.i.g. and 580° F.

COBALT SULFIDE (4% Co) ON SILICA-ALUMINA

This catalyst (No. 248-2) was prepared by impregnating 2000 milliliters of a crushed SiO₂ (87%)-Al₂O₃ aggregate with 150 milliliters of an aqueous solution containing 172.5 milliliters ammonium hydroxide solution and 373 grams EDTA along with 168 grams cobalt carbonate, the solution being heated until bubbling ceased before being added to the silica-alumina material which, in turn, had previously been dried for 24 hours at 400° F. Following impregnation, the catalyst was centrifuged and calcined for 4 hours at 1000° F., thus yielding a material having an amount of cobalt oxide equivalent to 2.2 weight percent Co. A second impregnating solution was then made up as above, using 150.2 grams cobalt carbonate, 334 grams EDTA and 154 milliliters of ammonium hydroxide and added to the catalyst. Following a holding period of 24 hours at 70° F., the catalyst was centrifuged and calcined for 10 hours at 1000° F. The calcined product so obtained was then alternately reduced in hydrogen and oxided in air (repeating the cycle 5 times) at 1000° F. and 1200 p.s.i.g. The catalyst was then sulfided by treatment with an excess of a mixture comprising 10% by volume of dimethyl disulfide in mixed hexanes at 1200 p.s.i.g. and 675° F., hydrogen also being present in the amount of about 6500 s.c.f. per barrel of feed.

COBALT SULFIDE (2% Co) AND CHROMIUM SULFIDE (3.53% Cr) ON SILICA-ALUMINA

This catalyst (No. 174-5) was prepared by forming an aqueous slurry with 1130 grams of the chelate of chromium and EDTA, to which slurry was added 196 grams of cobalt carbonate, the solution being then stirred until bubbling action ceased and made up to 1779 milliliters. This solution was warmed to 140° F. and added to 2280 milliliters of the crushed SiO₂—A1₂O³ aggregate. The resulting material was then held for 24 hours at 140° F., following which it was centrifuged and calcined 10 hours at 1000° F. The calcined product was reduced in an atmosphere of hydrogen at 1200 p.s.i.g. and 675° F., following which the cobalt and chrominum metals present were converted to sulfides by treatment with an excess of a solution comprising 10% by volume of dimethy· disulfide in mixed hexanes at 1200 p.s.i.g. and 675° F., hydrogen also being present in the amount of 6500 s.c.f. per barrel of feed.

2,044,008

MOLYBDENUM SULFIDE (2% Mo) ON SILICA-ALUMINA

This catalyst (No. 226) was prepared by forming 530 milliliters of an ammoniacal solution containing 41.4 grams of ammonium molybdate. This solution was then added to the crushed SiO₂—A1₂O₃ aggregate, previously dried for 24 hours at 400° F., in an amount sufficient to yield a dried product containing the equivalent of 2 weight percent Mo. After being held for 24 hours at 70° F., tlie impregnated material was centrifuged and calcined for 5 hours at 1000° F. It was then reduced in an atmosphere of hydrogen at 1200 p.s.i.g. and 650° F., following which it was sulfided in situ by treatment under these same conditions of temperature and hydrogen pressure with a hydrofined cycle oil (49% aromatics) containing 1% by volume dimethyl disulfide.

NICKEL SULFIDE (1% Ni) AND MOLYBDENUM

SULFIDE (1% Mo) ON SILICA-ALUMINA

This catalyst (No. 296) was prepared in the following manner. 28.6 milliliters of ammonia were mixed with 80 milliliters water and added to 49.3 grams EDTA, and to this solution was added 22.3 grams of nickel carbonate. After being heated to evolve carbon dioxide, this solution was mixed with another solution prepared by dissolving 78.7 grams of ammonium holybdate in a mixture of 80 milliliters of ammonia hydroxide and 80 milliliters of water. The resulting solution, on being made up to 480 milliliters by the addition of water, was then used to impregnate 600 milliliters of the crushed SiO2—A1₂O₃aggregate. The impregnated material, after being held for 24 hours at 70° F., was centrifuged and calcined for a period of 10 hours at 1000° F. It was then reduced in an atmosphere of hydrogen at 1200 p.s.i.g. and 650° F,, following which it was sulfided under these same conditions of temperature and hydrogen pressure with a solution containing 10 volume percent dimethyl disulfide in mixed hexanes.

The foregoing catalysts are those wherein the hydrogenating component has been sulfided prior to placing the unit on stream. If desired, however, the sulfiding step can be omitted, with the catalyst being placed in service in the pre-existing oxide form or in that resulting from a prereducing step wherein the oxide is converted in large part to the corresponding metal as it is subjected to an atmosphere of hydrogen, e.g., at 1200 p.s.i.g. and 575-600° F. In any event, the oxide, if not so pre-reduced, is inherently converted to said reduced state once the unit is placed in operation. When a nonsulfided catalyst is used, particular care should be taken in starting up the unit so as to avoid temperature runaways tending to foul the catalyst and unduly reduce its activity, an effect largely attributable to the overly strong hydrogenating action of the metal catalyst. Such temperature control can be effected by bringing the unit on stream at somewhat lower temperatures than would be employed with the sulfided catalyst, by reducing the feed rate, or by increasing the proportion of recycle gas supplied to the unit. However, once it appears that the catalyst has lost its unduly high exothermic characteristics (usually after about 10-25 hours) the run may be continued in the conventional fashion.

In the operation of the isocracking process, the charge stock may be introduced to the reaction zone as either a liquid, vapor or mixed liquid-vapor phase, depending upon the temperature, pressure, proportions of hydrogen and boiling range of the charge stocks utilized. This charge stock is introduced in admixture with at least 2000 s.c.f. of hydrogen per barrel of total feed (including both fresh as well as recycle feed), and this amount of hydrogen may range upwardly to 15,000-20,000 s.c.f. or more per barrel of feed. From about 1000 to 2000 s.c.f. of hydrogen is consumed in most instances in the isocracking reaction zone per barrel of total feed converted to syns thetic product, i.e., that boiling below the initial boiling point of the fresh feed. The hydrogen stream admixed with incoming feed is conventionally made up of recycle gas recovered from the effluent from the isocracking zone, together with fresh make-up hydrogen. The hydrogen content of the recycle stream in practice generally ranges upwardly of 75 volume percent.

Basically, the pressures employed in the isocracking zone are in excess of about 200 p.s.i.g. and may range upwardly to as high as 3000 or even 5000 p.s.i.g., with a preferred range being from about 400 to 1500 p.s.i.g.

Generally, the isocracker feed may be introduced to the reaction zone at a liquid hourly space velocity (LHSV) of from about 0.2 to 5 volumes of hydrocarbon (calculated as liquid) per superficial volume of catalyst, with a preferred rate being from about 0.5 to 3 LHSV.

In one method of practicing the invention process, the isocracking reaction is conducted at a given space rate under conditions of relatively constant conversion of at least 20% per pass, and preferably at constant converions falling in the range of about 20 to 70% per pass. Under this type of operation, the catalyst temperature is periodically increased to maintain the per-pass conversion at relatively constant levels. Alternatively, the process may be conducted at a constant temperature, in which case the per-pass conversion will gradually decline and the on-stream portion of the processing cycle will be terminated at an arbitrary conversion level. In either case, the decline in conversion level can also be offset to some extent by lowering the space velocity of the feed, though this procedure is not normally recommended. Additionally, the process may be conducted as a staged constant-temperature process in which the catalyst temperature is maintained constant for a periodic interval or for a time interval as determined by a prescribed drop in conversion level, after which the temperature is incrementally increased to again approximate the initial conversion level, and this temperature staging repeated over the conversion portion of the processing cycle. Still other methods of operating the process within the spirit of the present invention will suggest themselves to those skilled in the art.

As hereinabove prescribed, the process may be conducted at average catalyst temperatures in the range of about 450° to 800° F. In the preferred practice of this invention, the temperature at vzhich the reaction is initiated in a given on-stream period should be as low as possible (commensurate with the maintenance of adequate per-pass conversion levels, as discussed above) since the lower the starting temperature the longer will be the duration of the said on-stream period. For any given coversion, the permissible starting temperature is a function of catalyst activity since the more active catalysts (i.e., those capable of effecting a relatively high per-pass conversion under given operating conditions) naturally permit the unit to be placed on-stream at lower starting temperatures than would otherwise be the case. More specifically, the conversion reaction is preferably initiated at temperatures below about 730° F., and more preferably between 450° and 675° F. In some cases it may be desirable to initiate the reaction at temperatures below 450° F., with higher temperatures then being reached in a relatively short period of time as the catalyst becomes conditioned. With all except the most refractory feed stocks, and assuming the use of a catalyst of relatively good activity, it has been found that satisfactory conversion levels can be maintained while operating with average catalyst temperatures below about 730° F. during at least the first half of the on-stream portion of any given processing cycle (or, to put the matter in an equivalent fashion, during that portion of the cycle which is productive of at least one-half of the total product formed during the entire cycle), and this method of operation is observed in a preferred practice of the invention.

§,844,065

The low-temperature aspect of the subject iso-cracking process is a distinguishing feature Which is evidenced in product quality and yield, as well as in process advantages. One Of the major contributing factors to the unusually high octane characteristics of the synthetic gasoline fraction produced in accordance with the subject process is the preponderant production of iso-paraffins over that dictated by thermodynamic equilibrium. Such abnormal production of isoparaffins is in contrast to the conventional hydrogenation processing, such as conventional hydrocracking, which produce synthetic paraffins substantially at or below the thermodynamic iso-normal equilibrium ratio. In the present process, it has been found that the iso-normal paraffin ratio in the synthetic product increases with lower operating temperatures and is particularly increased as average catalyst temperatures are maintained below about 730° F.

These specifications of low-temperature operation also reflect the high yields of synthetic liquid product, e.g., product boiling below the initial boiling point of the feed. At operating temperatures below about 730° F. the synthetic liquid yields are consistently high, whereas at higher temperatures the liquid yield tends to fall off in appreciable measure. As a corollary to this drop in synthetic liquid yield there is obtained a corresponding increase in the production of Ci-C₃ gases. Thus, while methane production in operations effected at average catalyst temperatures below about 730° F. is substantially negligible, methane production rapidly becomes a limiting factor in the process at higher temperatures as evidenced by the fact that the amount thereof formed increases approximately 15-fold between a temperature of 700° F. and one of 800° F. Such methane production not only represents loss of feed stock but also (where hydrogen is being recirculated, as it must be in commercial operation) necessitates hydrogen purification, withdrawal of a portion of the contaminated hydrogen, or use of higher pressures in order to maintain the hydrogen partial pressure of the recycled gas within acceptable limits. This not only increases the cost of the plant, but also reduces yield due to the increased proportion of undesirable fixed gases in the product.

The following examples (other than comparative Example II) are presented to illustrate the practical application of the process of this invention in a number of its embodiments.

Example I. (Run 8-994A )

The feed stock for this run was obtained as a straight run distillate from a crude petroleum oil of Minas origin, the feed having the following inspections:

Gravity, ° API 58.0

Aniline point, °F. 142

Composition, vol. percent:

Paraffins 64

Naphthenes 31

Aromatics 5

Basic nitrogen content, p.p.m. <2

F-l octane number:

Clear 32.0 +3 ml. TEL58.5

Distillation, D-86:

Start ------------------------------° F__ 236

10% ------------------------------°F__ 254

3.0%--------------------------------° F__264

50% ______________________________° F__276

70%________________________________» F__287

90% ______________________________⁰ F__313

End point__________________________⁰ F348

The foregoing feed stock, along with 6500 s.c.f. H₂per barrel of feed, was preheated to 580° F. and contacted at a space velocity of 0.83 and pressure of 1200 p.s.i.g. with an isocracking catalyst comprising nickel .sulfide (2.5 wt. percent Ni) on silica-alumina, and cata10 io lyst having generally the same composition as No. 425-2 described above. The stream to the unit was gradually raised in temperature until the conversion to product boiling below 200° F. reached 60% per pass, the temperature at this point (56 hours) being 647° F. During the remaining portion of the run it was found that this conversion level could be maintained with but little further increase in temperature, the latter being 653° F. after 76 hours, at which point the feed was changed to initiate the operation described below in Example II.

In making this run, the portion of the product effluent boiling above 200° F. was recycled to the iso-cracking zone, i.e., the operation was one of extinction recycle. Tests conducted in companion operations show that this recycle stream, while responding to conversion in the isocracking zone in substantially the same fashion as fresh feed, makes an excellent jet fuel component since it is extremely low in aromatic content and is substantially free of fouling or deposit-forming characteristics when maintained at high temperatures for extended periods of time.

Based oh an analysis of the product recovered during the 20 hour period after conversion reached 60%, it was found that the volume percent yield of C₄-200° F. product (based on volume of feed converted to 200° F.~ product) was 109.2%. Of the C₄ portion of this synthetic product, 74.5% was made up of isobutane, an iso/normal C₄ ratio of about 3, which contrasts with the thermodynamic equilibrium value of 0.83. The yield of methane was extremely low, the same being 0.016% in terms of the weight of feed converted to 200° F.product.

The C₅-200° F. (TBP) gasoline fraction recovered from the isocracking zone in this operation had the following inspections:

Gravity, ° API________

Composition:

Paraffins_________

Naphthenes_______

Aromatics___

Octane number:

F-l-j-3 ml. TEL

Example II (Run 8-994 B)

The run of this example represented a continuation of that described above in the Example I, the only essential difference being the substitution of a different

83.8 percent-_ ____do____ —__do.____

99.8 feed stock at the expiration of the 76 hour Operating period there described. This new feed stock, like that 50 of Example I, was also a straight run distillate obtained from a Minas crude, the primary difference between the two feeds being that the one here employed contained approximately 27% of components boiling below 200° F. as determined by Hypercal distillation methods. This 55 feed had the following inspections:

Gravity, ° API 62.6

Aniline point, °F.141.0

Composition, vol. percent:

Paraffins 67

Naphthenes 29

Aromatics 4

Basic nitrogen content, p.p.m. <2

Distillation, ASTM D-86:

⁶⁵ 10%

30%

50% 70% 90%

-----------------------------°F„158

-----------------------------⁰ F__ 196

-----------------------------·» F— 217

-----------------------------» F__ 245

-----------------------------’ F__ 265

-----------------------------°F- 307

End point_________________________° F__ 340

In this run, which was conducted without recycle of 200° F.+ product, it was found that the conversion (in terms of feed boiling above 200° F. converted to product 75 boiling below said temperature) could be maintained at

2,944,005

60% per pass while holding catalyst temperatures substantially constant at 650° F. The yield of C₄-200° F. product was 105 volume percent. The C₅-200° F. portion of said product had the following inspections:

Gravity, ° API--------------------Composition, vol. percent:

Paraffins ------ 82

Naphthenes -----------------------—16

Aromatics ------------------- 2

Octane number:

F-l clear-|-3 ml. TEL-------—--------94.2

Comparing the C₅-200° F. product obtained in this operation with that of Example I, it will be seen that the latter had a leaded F-l octane rating (99.8) which is far higher than that exhibited by the corresponding fraction of the instant operation (94.2). From this and other factors it is judged to be critical to a practice of the present invention to employ a feed stream which is essentially free of components boiling below 200° F. if a C₅-200° F. gasoline fraction of high octane rating is to be recovered from the isocracker effluent.

Example 111 {Run 8-934)

This run was conducted with a raffinate feed stock obtained by extracting a catalytic reformate with silica gel. The feed had the following inspection characteristics:

Gravity, ° API_______________________:60.2

Composition, vol. percent:

Paraffins---------------—---— ?1

Naphthenes —: 26

Aromatics 3

Basic nitrogen content p.p.m. —- 0

Octane number:

F-l clear 32 +3 ml. TEL 64

TBP distillation:

Start______________________________°F— 205

End point _—_—----------------------° F— 367

The foregoing feed was isocracked using catalyst No. 316, as described above, said catalyst containing nickel sulfide (2.5% Ni) on a support made up of silicaalumina. The catalyst had previously been employed in seven isocracking runs and had been regenerated at the conclusion of each of said runs and reduced in an atmosphere of hydrogen before being again sulfided for use in the succeeding run. In this particular operation, the catalyst had already been in service for approximately 42 hours using a similar raffinate feed before being switched to the feed of the present operation for an additional 48-hour operating period. To complete the run description, it may be noted that still a third raffinate stock of generally similar characteristics was thereafter supplied to the catalyst for an ensuing 50 hour period, the total run length thus being 140 hours.

The present operation was conducted by preheating the feed admixed with 0.1 volume percent dimethyl disulfide and 6500 s.c.f. H₂ per barrel of feed, to 760° F. and passing the heated mixture through the catalyst at a LHSV of 1.05 and pressure of 1200 p.s.i.g. The temperature was maintained constant throughout the run, and conversion thus fell off as the run progressed due to catalyst fouling. At the point of the run where the present feed stock was introduced (42 hours) the conversion to product boiling below 205° F. was 70—75% per pass, and at the end of 90 hours conversion was approximately 52% per pass. The portion of the product effluent boiling above 205° F. was recycled to the catalyst along with fresh feed and makeup hydrogen, the operation thus being one of extinction recycle.

Analysis of the product disclosed that approximately 0.17 to 0.2 weight percent methane was formed, based on feed converted to 205° F.~ product, a rate of methane production which is approximately 10-15 times higher than that observed in operations conducted at 10 650° F. (Example I). The yield of C₄-205° F. (TBP) product was 107.2%, while that of the C₅-205° F. (TBP) product was 60.6%. The latter product had the following inspections:

₁₅ Gravity, ° API 84.6

Composition, vol. percent:

Paraffins 94

Naphthenes 5

Aromatics ---------------------— 1 ₂₀ Octane number:

F-l+3 ml. TEL 97

By reducing the cut point of the product obtained in the foregoing operation from 205° F. to 200° F., thereby freeing the product of low octane n-heptane, it is 25 estimated that the leaded octane number would be increased to a value of 99.

This application is a continuation-in-part of application Serial No. 594,968, filed June 29, 1956.