GB2136826A - Refining and cracking carbonaceous materials - Google Patents

Abstract

A process for converting into lighter viscosity products the heavy fractions of petroleum or other hydrocarbon containing components comprises the use of steam and catalysts, at least one of which is prepared by treating an alcoholic solution of alkali metal hydroxide with hydrogen sulphide. The bottom and top fractions from the catalytic treatment may be each subjected to catalytic treatment.

Description

SPECIFICATION Cleavage and hydrogenation of refractory petroleum residue products, such as asphaltenes, resins, and the like This invention relates to the treatment of petroleum and petroleum residues; more specifically, this invention pertains to the treatment of petroleum, petroleum residues in general, but especially high boiling residues according to ASTM D-1180 and D-1160 method, or residual fractions having no boiling points according to these methods or which are destructively distilled with formation of carbon residues. Further, this invention concerns an interdependent treatment of component parts of the above petroleum and petroleum residues for improving the overall yields and space-time velocities including the use of somewhat higher pressures, such as pressures up to 1 50 psi to aid further in the recovery of the desirable products.

PRIOR ART In general, the prior art has attempted the upgrading of heavy petroleum fractions with limited success. Many of the petroleum components, as residues, have been especially refractory to further treatment. By stripping all of the oil components from the bottoms, such as by steam stripping and the like, or by further treating of these such as by hydrogen transfer reactions, partial coking, delayed coking and the like, these bottom fractions have been improved and/or further components obtained therefrom. For example, by using a hydrogen donor solvent, some of the hydrogen from the solvent is transferred to the higher boiling petroleum residue fractions and thereby the viscosity of a part of a high boiling fraction is improved such that lighter products are obtained.Typically these reactions are carried out at very high temperatures and with or without the presence of hydrogen under high pressure.

Other methods have been used whereby hydrogen obtained by stripping hydrogen from part of the very refractory residues have been used in a parallel stream employing hydrogen as a reactant gas and the coked product thereafter employed for other purposes or burned. The excess hydrogen may also be used such as in the donor solvent system or elsewhere in the refining cycle used to improve the residues, In my previous work (U.S. Application Serial No. 140, 604), published on November 18, 1981 in Great Britain as U.K.Patent Application 2,075,542, 1 have disclosed a process for hydrotreating petroleum residue materials in which the residue material is contacted with steam and with alkali metal hydrosulfides or the empirical monosulfides or polysulfides and/or the hydrates thereof and mixtures of the foregoing to hydrocrack, hydrogenate, hydrotreat, denitrogenate, and/or demetallize and/or desulfurize the carbonaceous material. According to the process disclosed in my above application, the treatment of petroleum and petroleum residues has been considerably improved based on yield, conversion per pass, space-time velocity, API number, etc., such that the process produces upgraded products of greater value.

In my U.S. Patents 4,366,044 and 4,366,045 to different processes which relate to the treatment of coal, a number of prior art references are also mentioned. The references as cited in these two patents and as found of record therein provide a background information to the alkali sulfide chemistry and also include, in part, references related to the treatment of crude oil or petroleum refining residues. These references show various attempts in the field to improve the residues or treat various carbonaceous materials.

Other prior art, based on U.S. patent literature, has been summarized in M.J. Satrian, Editor, "Hydroprocessing Catalysts for Heavy Oil and Coal", Noyes Data Corporation, Park Ridge, New Jersey, U.S.A. (1982), e.g. pp. 62 to 78. In the U.S. patent literature the following patents have been noted (even though these may have been mentioned inter alia in the above art): U.S.

Patents 3,520,825; 3,775,346; 3,850,840; 4,117,099; 4,119,528 and 4,203,868. This prior art, however, does not disclose the herein described invention, which represents a further discovery and an invention over my process mentioned in published U.K. Application 2,075,542.

Other less relevant prior art is also found, e.g. in Fuel, Vol. 62, Feb. 1983, which issue is devoted exclusively to fundamental studies of catalytic coal and carbon gasifications (cf. Kikuchi et al., "Supported Alkali Catalysts for Steam Gasification of Carbonaceous Residues from Petroleum, ibid., pp. 226 et seq., who report carbon deposits on supported catalysts, generally attempt substantial gasification, and had no cognizance of the role sulfide series of alkali metals play in these reactions).

Still other art has been disclosed by Sikonia et al., "Flexibility of Commercially Available UOP Technology for Conversion of Resid to Distillates", Publication AM-81-46 of National Petroleum Refiners Association (NPRA), 1981, Washington, D.C., U.S.A.; Ritter et al., "Recent Developments in Heavy Oil Cracking Catalysts", Publication Am-81-44, NPRA, 1981; Bartholic et al., "Utilizing Laboratory Equipment in New Residual Oil Development", Publication AM-81-45, NPRA; 1981.

BRIEF DESCRIPTION OF THE PRESENT INVENTION As part of the process disclosed herein, it encompasses a method of using an ebullating bed reactor and supported catalysts as well as a process for treatment of the especially refractory petroleum residues, for example those boiling over 850OF and especially asphaltenes or the ORA fractions, that is a mixture of oil, resin and asphaltenes left over from the petroleum refining and characterized as boiling at 1000'F and higher according to ASTM D-1180 and D-1 160 method. That aspect of the invention represents a further more successful conversion of the especially refractory bottom or residual products.

A further aspect of the invention is to provide a further contribution to the process briefly described immediately above. That aspect includes a more facile and advantageous treatment of petroleum or crude oil, as such, as a starting feed as well as the less completely refined petroleum or crude oil residues, such that when treating crude oil, substantially no residue remains and the starting material is converted substantially completely to lower boiling point products. This result is accomplished because, according to my process, the severe treatment which produces conventional residua is not applied to the crude oil.

By a specified boiling point, as discussed herein, is meant a boiling point such that the products will be substantially completely distilled at that boiling point (according to the defined methods) including up to that temperature, and no substantial amount of residue remains unless so specified.

BRIEF DESCRIPTION OF THE INVENTION The present process distinguishes from the prior art in the specific catalyst species which are being used preferably in a supported form, to attack the extremely refractory petroleum residue components such as boiling over 1,000OF and usefully convert these residue components into highly desirable lighter viscosity products in a highly efficacious manner.

My present process further distinguishes from the prior art in that the reaction is specifically attacking the most refractory components of the residue such as asphaltenes, while at the same time avoiding, due to the discovery of the catalytically aided thermal decomposition by the herein used catalysts, the unfavorable effects of coke formation which may occur if the process is not carried out properly. Moreover, further process advantages are realized by the employment of an ebulllient bed reactor, by carrying out the process continuously and at higher temperatures such as up to 650'C, yet at the same time quenching the conversion product to obtain a desired end product in another reactor(s) in combination with the first reactor.

In accordance with the present process, the ORA fraction obtained from the residues such as from petroleum refining or refining of any other carbonaceous source yielding these fractions are being treated advantageously by the present method. However, with outstanding results the present process is useful for treatment of asphaltenes, e.g. of solvent extracted asphaltenes when these are being treated in one or two stages to further cleave and/or hydrogenate in one or two stages this especially refractory product and thus to improve the overall yield obtainable from a barrel of oil.

This improved process is also characterized by the ability of the specific catalyst to convert the Ramsbottom or Conradsen carbon into usefully hydrogenated products without affecting, to any noticeable degree, the process as practiced herein.

Still further, the present process also provides for especially advantageous catalyst support combinations which can be used such as in an ebuilient bed reactor and produce the conversion products in an especially advantageous manner without the catalyst support combination being affected by the unwanted metal constituents found in petroleum and accumulated predominantly in the ORA fraction.

Still further, it has been found that after the first stage conversion with the active catalysts, a second quench stage may be provided where the cleaved product may be appropriately tailored by a specifically selected catalyst in the second reaction stage to produce the predetermined and/or highly desired product cuts. However, the second stage reaction is interdependent and based on the specific catalyst in the first stage, and based on the properly carried out reaction in the first stage reactor.

Moreover, the present process allows also a recycle of the product not adequately reacted in the first reactor and its subsequent conversion into the desired end product.

Also, it has surprisingly been found that my process can be further improved by the parallel treatment of the various products first recovered from the ebullating or a fluidized bed reactor and split in at least two streams as further described herein. It has been found surprisingly that the split stream treatment procedure (which is a straight through treatment for the first gaseous or vapor phase, i.e. top products), for the bottoms thereof which are obtained by collection in a further separating device such as cyclone, contributes further to the overall conversion of the starting material by this separate treatment of the heavy fractions carried over such as from the ebullating bed.This is due, in part, to the application of different catalyst combinations to overcome some shortcomings of the catalyst compositions when these are used to treat the highly diverse components of the original feed or the recycle thereof. Moreover, the bottoms being pre-treated to a certain degree (e.g. as characterized by lower API values benefit from this parallel treatment.

This separate treatment in at least one parallel stream of a smaller, yet initially upgraded fraction coming over from, e.g. an ebullating bed reactor provides a number of benefits. The catalyst movement is about 2 to 40 inches/second/ 1 0 to 30 minutes of residence time of particle feed with about 1 7 inches/second/20 minutes of residence time of particle of feed being average. (The catalyst movement is defined in accordance with ebullating bed technology.) Additional ebullation is caused by steam and feed being fed into the reactor.At the same time this treatment makes it easier to obtain improved throughput rates, yields, conversions, type of products, etc., for the top products from the first stage such that the overall flexibility of the process for treating a wide variety of feed is remarkably aided by this parallel treatment as disclosed herein, especially allowing the split proportions to be tailored to accomodate the various feeds with different components fractions, as well as the mixed product feeds of various source materials which are especially difficult to treat.

That aspect of the process where the heavy bottom fractions may be recycled but without parallel treatment, the results are outstanding, but further improvement has now been found contrary thereto. That aspect of the process is called a "recycle method". The improved aspect is called a "parallel method". Thus, a separate treatment without recycling is more advantageous in the overall improvement of the products and product yields to be obtained from the wide variety of heavy crude oils, petroleum refining fractions, and the especially refractory component parts. In fact, the parallel method as disclosed herein causes almost an entirely complete recovery of all crude oil components as useful products.

The above process variations are characterized by and their accomplishments made evident from the improved product being obtained from the very poor starting material, the higher hydrogen content of the starting material, the lower viscosity and smaller molecular size of the cleaved product and the amenability of the treated product to further conventional treatment steps.

Inasmuch as the present process accomplished considerable hydrogenation which improves the yields, improves the product, and provides smaller molecules and thus a less viscous product and eliminates and/or minimizes to a very significant degree the presence of free Conradsen or Ramsbottom carbon, the attained end result shows an outstanding achievement in the continuous search for utilizing all fractions of a refinery product which heretofore could not be advantageously upgraded to the degree such as now disclosed herein.

The process as disclosed herein,. i.e. consisting of the "recycle method" and "parallel method", is, thus, especially useful because the petroleum fractions which previously were reacted according to prior art under very severe conditions can be reacted by means of the disclosed catalysts in a less severe, straight through fashion, and either a recycle or in a parallel method. The very refractory bottom products (by conventional practice), constituting a smaller part of the total mixture and which are the hard to treat, are either pre-treated and recycled or, instead of being initially more drastically attacked, e.g. bottom reboiled and/or recycled, can now be treated in a parallel treatment with the necessary degree of severity.This treatment severity for the recycle or parallel method is still markedly and significantly less than that necessary in the prior art, but accomplishes far better results as characterized by more valuable product components, more product recovery and simpler recovery steps of components and other beneficial features further explained herein.

Further, the process as disclosed herein also applies to other products such as Gilsonite (also known as uintaite, a black lustrous asphalt occurring in Utah), pitch (natural and/or refinery product), tars such as from tar pits found in California, asphalts (natural) such as found in asphalt lakes in Trinidad and of the type found in Utah and like products found in nature or obtained after refinery treatments which concentrate oils, resins and asphaltenes in the residues; slurry oil found in refineries which oil is often without a distillation point; bottom oil from coal gasification plants, e.g. solvent extraction products of coal, or coal extracts such as shown by Satrien, above, p. 206, or bottom fractions or retorted shale oils; heavy bottoms from the Lurgi gasifier process which are produced in operation of the SASOL process in South Africa; delayed cocking products obtained during bottom distillation of these products; resins both hard and soft and asphaltenes. It is inadvisable to mix some of the feeds, but that can be determined by experiment at the actual reaction conditions. By heavy crude oil is meant crude oils of an APT number of about less than 22 including the minus number API crudes.

The combination process as disclosed herein provides further that the more volatile fractions from the petroleum and the like hydrocarbon can be treated in a straight through fashion, but the heavier bottoms, that is the heavier components, especially the more refractory components, are most advantageously treated in a recycle or parallel run, depending on the severity of treatment required. For example, the feeds requiring the most severe treatment are best treated by the parallel method. Hence, the method most suitable for the feed is selected which will not impede or detract from the overall, first reactor stream yields, and yet improve the recovered bottom products by a recycle or in a parallel treatment sequence for improved overall yields.

While the "parallel method" aspect has been disclosed with respect to a two reaction train or two stream parallel treatment, equally applicable are further parallel treatments where the refractory products obtained from the second reaction train or stream may be further reacted without recycling.

As a result of the above disclosed combination of steps, the present invention provides for advantageous space-time velocities, yields on an absolute basis, yields per respective parallel pass or combined yields.

As an example for a parallel treatment, the products obtained from a mixed heavy bottom fraction of especially refractory products have yielded end products having a K factor of about 11.71, i.e. highly paraffinnic products, which is achieved by the greater ease the tops from the first reactor and subsequently treated, e.g. to reduce the bromine number and the like.

ILLUSTRATION OF THE PROCESS BY A FLOW SHEET AND DESCRIPTION OF THE DRAW INGS In the drawings herein: Figure 1 illustrates schematically the recycle method in an ebullating bed reactor, and Figure 2 shows schematically a continuous process wherein by the parallel method the various source materials are converted into useful products.

As illustrated in Fig. 1, the first stage reactor is an ebullating bed reactor 10. It consists of a reaction vessel 11 with a catch funnel 1 5 at the top and a pump 14 for recirculating the fluid 8 products undergoing the reaction. The catalyst 9 which is in a dispersed supported form is ebullating with the fluid 8. A fluid 8 undergoing the reaction including the catalyst 9 therefor may be made to circulate by the pumping of the fluid into the reactor and appropriately distributing the same. Typically, the circulating fluid might be introduced at the bottom, but introduction may be elsewhere in the reactor at one or more places.On a smaller scale, and as a close approximation of an ebullating bed reactor, a stirred tank is appropriate, provided the supported catalyst material is placed in a cage(s) or baskets such as four stainless steel mesh envelopes and these attached to a suitable frame driven by an external motor. By varying the speed of the rotation, e.g. 20 to 180 rpm, more typically from 50 to 1 50 rpm, reactions very closely approximating those in an ebullient bed reactor are achieved. Again, it is important that adequate steam-catalyst-fluid contact take place to assure the desired result.

Circulatory fluidized bed reactors where the supported catalyst circulates with the fluid or fluidized bed reactors may also be suitable. Similarly, a fixed bed reactor with the fluid downflowing or upflowing may be used for that purpose.

A continuously introduced pre-heated feed charge such as asphaltenes or an ORA cut are introduced via pipe 1 2. Steam 1 3 may be introduced with the feed or it may also be distributed in reactor 10 throughout the fluid from the bottom of the reactor 10.

The obtained lighter viscosity fluids are conducted by a large diameter type or conduit 1 5a and appropriately cooled if needed in a quench zone 16 and then introduced in a second stage reactor 1 7. Although one reactor has been shown in the Figure, a number of reactors in series or in parallel may also be used. These reactors may be fixed catalyst bed, gaseous or vaporphase reactors. A column where the catalyst is shown as supported on the trays 1 7a is typical.

Other well known devices may be employed for this purpose such as trickle bed reactors and the like. The bottoms from the second reaction stage are collected in the collection zone 18. These are classified as No. 1 bottoms. These bottoms may be recirculated entirely or partially into the ebullating reactor with the fluid collected by the ebullating bed funnel conduit 1 5. If necessary, a pump 1 4A may be used for that purpose. Part of the product may also be diverted and recovered for further processing.The top fraction from the second stage reactor 17- may be refluxed via reflux boiler 1 9 and the products may be diverted from this reflux boiler and the gaseous products therefrom further worked up in a second reactor such as depicted for 1 7 but now shown herein, these are all called second stage reactors as distinguished from the first stage reactor 10 where the cleavage of the source material is undergoing. The second stage reactor products and other gaseous products may be treated in a further reactor or such as by bubbling through an appropriate bath to remove any unwanted constituents such as hydrogen sulfide. This may be accomplished in a vessel 20 in which potassium hydroxide has been dissolved.

The gaseous products are thereafter recovered in a conventional manner.

As illustrated in Fig. 2, the first stage reactor 110 is an ebullating bed reactor. It may also be a stirred basket or cage type reactor or a fluidized bed reactor. The rotating, i.e. stirred cages or baskets, have not been shown herein. However, inasmuch as the ebullating bed and stirred cage or basket type reactors are very closely related in their actual behaviour and performance, these will be discussed as being the same.

The reactor 110 consists of reaction vessel 111 with a catch means such as a funnel 11 5 at the top and a pump 114 for recirculating the fluid 108 products undergoing the reaction. The feed or fluid 108 may be petroleum or light petroleum residues boiling below about 850OF, but may also be the ORA (oil, resin, asphaltene) boiling above that temperature, e.g. above + 1,050OF, i.e. having no boiling point. A catalyst 109 which is in a dispersed, supported form is ebullating with the fluid 108. A screen 109a prevents the catalyst from upflowing with the ebullating liquid 108.However, the fluid undergoing the reaction including the catalyst therefor may or may not circulate by the pumping into the reactor through pump 114 because the catalyst may stay suspended in the fluid (for circulating fluid bed reactors) or only a liquid phase freed of the catalyst may be caught in the funnel 11 5. As shown in the drawing, the ebullating fluid pumped through pump 114 may be introduced anywhere in the reactor, although typically it is introduced near or at the bottom thereof.

The feed 108, of the above-described type, is joined in a single introduction port with steam and/or water 11 3 or the introduction ports for either of these may be circumferentially around the vessel 111. The refractory components of petroleum, having a longer residence time, undergo the catalyzed reaction and are carried over by the exit conduit 11 spa. This conduit then leads directly into a cyclone 11 6 which separate the vaporous phase products from the liquid phase and/or entrained products and products carried over with steam from the ebullating bed.

More than one cyclone 11 6 may be used. The cyclone bottom products 11 2a are treated in a separate parallel stream.

Top products coming from the cyclones 11 6 may be quenched in a quenching zone 11 6a as shown in the drawing, and thereafter immediately reacted in a reaction vessel 11 7. The products may not be quenched prior to introduction into reaction vessel 11 7 as quenching may take place therein. The reason for quenching will be explained later.

After the introduction into vessel 11 7, an appropriate catalyst, as further described herein, supported or unsupported, is suspended, such as schematically shown, on tray 11 7a or by other means. The catalyst will cause the reaction in the gaseous or liquid stream to produce again a lighter fraction and a heavier fraction or the gaseous steam may be hydrogenated. These bottom products may be recovered such as in a vessel 118, and further treated in a conventional fashion or, as disclosed further herein, may be treated separately in still another parallel stream.

Although not shown in the drawings, additional reaction vessels in the same stream of a type such as 11 7 may be provided. The final recovery of the gases from reflux condenser 11 9 or other condensers (not shown) may be accomplished in the conventional fashion such as by cooling and/or low temperature chilling and/or pressure condensation and need not be discussed.

It has been found advantageous that the gaseous products of the initial straight through product, after treatment in vessel 117, be further treated at a temperature such as at about 130"C to about 240'C in a reaction vessel of the type as shown in the Figure at 11 7. If more saturated products are desired, a reflux column 119, appropriately selected and adjusted, may be employed to collect these products.

The bottoms 112a, from the reaction product separated such as in the cyclones 116, are generally the more refractory components of the petroleum or petroleum products. These products 112a, although believed to be partially reacted in reactor 110 (e.g. as characterized by the improved API number), are best further treated in a first parallel pass. As these bottom products have reacted to a certain degree due to the catalyst 109 suspended in the ebullating bed, the further reaction is best carried out with another specific catalyst appropriately modified, if necessary, to suit the particular feed.

This second feed stream is thus the bottoms 112a, fed to a reaction vessel 1 30 conjointly with steam, and this co-fed stream lends itself to a catalytic treatment of the still hot product.

The feed 11 2a is more amenable to hydrogenation and thus further improvements are better achieved for obtaining lighter viscosity components of the heavier fractions of petroleum. Feed stream products 11 2a may be advantageously treated with a different catalyst composition in the presence of steam in the same ratio ranges for steam as for the feed fed to the first reactor 111. Various catalyst compositions have been found to be more advantageous in the subsequent reactions in contrast from the catalyst which has been originally used in the ebullating bed reactor and designated as 109 therein. The use of a different catalyst is due to the partial reaction of the more refractory components in vessel 111. Again, the different catalyst composition for treatment of the various bottom'fractions will be further discussed herein.

As the reaction vessel 1 30 may still be followed with another reaction vessel (not shown), the top products recovered in reflux column 133, from reaction vessel 130, may be further treated.

Conventionally, the overhead products from reflex column 1 33 may be scrubbed in vessel 1 35 to remove unwanted constituents, e.g. H2S and the gaseous products thereafter recovered and reused. The reaction vessel 1 30 may be used in a similar manner as reaction vessel 117, i.e. as an upflow or downflow reactor, with a supported or unsupported catalyst suspended on a bed or on trays 130a, or as a fixed bed catalyst. A liquid catalyst, such as hydrated melts of the various alkali sulfides, liquid at the temperature at which these are desired as previously disclosed by me, e.g. in my U.S. Patent 4,366,044 may also be employed in the subsequent reactors.

If no further reduction of the molecular size of the overhead distillate product from the first reactor 110, is desired, the overhead distillate product should be hydrotreated in vapor phase reactors, immediately, as it emerges from the cyclone 11 6 which separates the bottoms (the heavy, partially reacted or uncovered recycle stock), from the overhead distillate which is the product. Quenching such as in quencher 11 6a prevents recondensation of the cleaved components. Hydrotreating is aided by the absence of the heavier fractions (removed by cyclone 11 6). By definition, the hydrotreating is done in the absence of either thermal or catalytic cracking and includes the addition of hydrogen to a molecule.As the hydrogenation reaction is exothermic, the previously described quencher 11 6a helps to control the temperature and prevents product damage if excessive temperatures are encountered.

Fixed bed reactors may be used for this hydrotreating, or liquid catalyst reactors may be used.

However, at present, the liquid catalyst reactor appears most effective at temperatures below 280"C, due to the solidification of the catalyst above 280"C, as a result of the interconversion of the hydrates.

The fixed bed reactors are less efficient in terms of the degree of hydrotreating achieved from the liquid reactors.

Turning now to the description of the feed, typically an ORA fraction, that is oil, resin, asphaltene fraction, is described as one which boils at atmospheric pressure at a temperature above 1000"F. Although this is a rough description because the amount of oil, resin and asphaltene are not necessarily ascertained, it is a convenient measure for this the most refractory component in an oil. It is known that asphaltenes can be solvent extracted from petroleum residues of from the ORA fraction to allow the oil and resin residue to be further treated.

Asphaltenes, of course, are especially intractable to further treatment such as by catalytic or other means, and thus constitute a fraction which can only be usefully burned or coked by the prior art methods to strip all available hydrogen therefrom.

A convenient characterization of the asphaltene is that it is that portion of the asphalt or bitumen which is soluble in carbon disulfide, but insoluble in paraffins, e.g. heptane, paraffin oil, or in ether. Resins from the ORA fraction may be extracted with propane. Bitumens are also soluble in carbon disulfide. Carbenes, which are consitituents of bitumen, are insoluble in carbon tetrachloride but soluble in carbon disulfide. Further, the oily or soft constituent of bitumen is also named malthenes or maltenes. These are soluble in petroleum spirits. Malthenes are pentane soluble compounds and asphaltenes are pentane insoluble compounds.

Still further and in a broader sense, the natural asphalts such as petrolene, mineral pitch, earth pitch, Trinidad pitch, petroleum pitch, and native mixtures of hydrocarbons such as amorphous solid or semi-solid fractions produced by oxidation of residual oils are included within the above definition.

Inasmuch as there is no agreement on the exact definition of these compounds such as malthenes or asphaltness, mixtures are often reported as one or the other in the prior art.

Moreover, the solvents used and the extracted and precipitation techniques practiced affect to a lesser or greater degree the end product properties. For this reason, the solvent extracted asphaltenes such as carbon disulfide extracted asphaltenes precipitated from heptane are still not considered pure compounds as these have no specific melting points but only softening points. Asphaltene softening points may be up to 400OF and higher. As a result, a convenient measure is to define the ORA fraction as one boiling at 1000OF and higher, although this temperature limit is arbitrary and lower temperatures such as 900OF may be selected because all of the material may not be desirably stripped away. Hence, a lower temperature of 800OF merely characterizes a less intractable composition.

For example, a high softening point solvent extracted asphaltene will have a softening temperature of about 270OF, a specific gravity at 60"F of about 1.1149 and a viscosity at about 275OF of about 4060 poises. The specific gravity at 275OF is 1.026, and thus the viscosity is 3,957 strokes or 395,700 cST (centistokes) (stokes are obtained by dividing the poises by a specific gravity at the indicated temperature). Viscosity at 300OF of the same high softening point asphaltene is 877 poises with a specific gravity at 300OF of 1.016, and 86,000 cST.

Viscosity at 325OF for this asphaltene is 261.5 poises, and specific gravity at 335OF is 1.006 giving 26,000 cST.

The analysis for the above solvent extracted asphaltene is found in the following table.

ASPHALTENE ANALYSIS Metals Carbon 84.59% Fe 360 ppm Hydrogen 8.80% Ni 147 ppm Nitrogen 0.82% V 490 ppm Sulfur 5.52% Na 497 ppm Ash 0.27% K 4 ppm Moisture 0.0% Oxygen Total 100.00% B.T.U. content of asphaltene: 1 7,627 Although the above asphaltenes may be considered as representative, various other asphaltenes, depending on the source, may have different characteristics.

Based on the above product analysis, it is seen that these products contain considerable amounts of metals. These amounts vary based on the source of the material and may range up to 6,000 parts per million (ppm) of vanadium, but typically up to about 600 ppm. Nickel and other constituents may also be present up to about the last named amount. Consequently, these metals also affect the ability of the residue to be treated by conventional methods of petroleum residue treatment.

Based on the various analyses, typically the hydrogen content of the ORA fraction may range from 13.5% to about 7%, and lower by weight, but again this is not a precise characterization.

In a petroleum residue boiling over 1,000OF, the hydrogen content will be about 12.5% and lower. A considerable precentage of "free" carbon is also found (as Conradsen carbon), e.g. up to 45% by weight. The free carbon is defined as Conradsen carbon or Ramsbottom carbon, but these analyses are not identical because different methods are used to define the the "free" carbon which, in fact, may not be "free". In the ORA fraction, Conradsen carbon may range up to 40 + %, by weight. In any event, the carbon residue is amenable to conversion according to the process as disclosed herein.

Any cracking, either thermal or catalytic, produces smaller molecules than were present in the feedstock being cracked. A smaller molecule will have a lower temperature boiling point than a larger molecule. Hydrotreating appears not to produce cracking (and thereby form smaller molecules). The boiling point temperature range seems to be approximately the same before and after hydrotreating.

Hydrotreating is effective, for example, in the hydrogenation of olefins to paraffins, and in the elimination of diolefins and triolefins in the product by hydrogenating these diolefins and triolefins in the product by hydrogenating these diolefins and triolefins to either mono-olefins or paraffins. Aromatics are also converted to naphthenes by hydrotreating. True paraffins appear not to show a reaction during hydrotreating.

Hydrotreating may also reduce the sulfur and nitrogen content of the product. However, if the asphaltenes, which contain most of the nitrogen, have not previously been cracked, the nitrogen of said asphaltenes is not available for removal by hydrotreating. The degree of sulfur removal and nitrogen removal is equivalent to the hydrotreating of an unsaturated bond (with no hydrogen present at the position at which the hydrogen is to be inserted into the molecule).

Those sulfur and nitrogen atoms, in an exposed position in the chemical structure of the product, and which nitrogen or sulfur atoms can be replaced by hydrogen, in a thermodynamically favorable exchange, with the formation of ammonia of hydrogen sulfide and the formation of more saturated molecule of the product, are the sulfur and nitrogen compounds that may be removed by hydrotreating.

The bromine number is a measure of the degree of unsaturation of the product. The lighter unsaturated product cuts will have the highest bromine numbers. The hydrotreating of the product will bring the bromine number down to acceptable levels, e.g. a bromine number of 40-50 for the initial boiling point of 400OF product (obtained by distillation within that range of the product by the method(s) disclosed above).

As previously mentioned, the bottom products 31 may form a still separate parallel path and may be used as a feed, but these can likewise be recycled because the further reaction pass only introduced further complications. Nevertheless, if the bottom products 31 are difficult to treat, by a further split, these may be subjected to still another treatment stream of the type as shown for feed 12a, and specifically in Fig. 2 herein.

Inasmuch as most of the overhead products from reaction vessel 11 are treated in one path straight through, the bottom fractions 1 2a such as collected in cyclone 1 6 or in the collection vessel 18, constitute products, but of a smaller proportion depending on the component composition of the petroleum. These bottom products may be thus treated in smaller vessels based on the proportion of the components of the petroleum products being treated. Accordingly, various mixtures of petroleum can now be usefully treated depending on the bottom product 1 2a components. The size of the reactor 30 may range in proportion to the bottom components in the petroleum such as those boiling at a temperature of 1000oF and above.

Steam is being introduced in the reactors 110 or 130 at a rate such that the product sought to be obtained dictates, to a certain extent, the amount of steam being used. The steam may be from about 50 and up to 100 to 130% of the product recovered, by weight. Steam may also be expressed on a basis defining a lower limit, namely such that coking does not occur to any substantial degree due to absence of steam, but which would occur in the reactors 110 or 1 30 if insufficient steam were present. Thus as a lower limit, the amount of stream introduced must be such that any significant coking is avoided in reactors 110, 117, 130, etc. Additional steam may be supplemented such as for reactors 117, 130, etc. Steam is low pressure, typically waste steam.

The temperature in the reactor generally may be up to 650"C, although the most advantageous operating temperature is below 450"C, such as below 425"C. The temperature, however, in the subsequent reaction vessels, e.g. 117, etc., may be lower, and lower temperatures seem to produce a more hydrogenated product, such as characterized by a lower bromine number. However, the more refractory component treatment in reactor 1 30 may require the same temperature as in reactor 110.

The quenching may be carried out at a temperature of 390"C and lower. The third stage reactor, not shown of 390"C and lower. The third stage reactor, not shown herein, may have a temperature as low as in the range of 170"C and lower, i.e. down to 125"C (but without condensation of steam). Lower temperatures are desirable because hydrogenation is exothermic and most of the hydrogenation is carried out in the downstream reactors, whereas most of the cracking is carried out in reactor 110, and to a lesser degree in reactor 1 30.

The pressure in the vessel 111 and downstream thereof may be maintained at an atmospheric pressure. Although subatmospheric pressures are possible, these are not as convenient. If the pressure is between about 100 psig or below 1 50 psig (although up to 25 atm. may be used), it seems that the reaction may be improved. When higher pressures are used, between 60 psig and below 1 50 psig for some unknown reason suggest the best range. Steam condensation must be avoided at the higher pressures. Although the exact reason is unknown and the conversion is not understood, it seems to indicate a better hydrogenation in that range, but the increase in pressure carries with it a certain trade-off due to more complicated equipment and more capital investment.

Further, instead of cyclone separation such as depicted in the drawing herein, i.e. one or more of the cyclone 116 being employed, other separation means may be employed to recover the liquid and/or entrained products and thus to accomplish the parallel path treatment of the petroleum components; such other means may be centrifuge separators, knock-out drums or the like.

With respect to the materials which may be subjected to the treatment, these are various crude oils from whatever source and of whatever viscosity. These crude oils, as long as these are adequately liquid at the treatment temperature, may be treated and produce the outstanding results as disclosed herein. Atmospheric residues and bottoms at atmospheric conditions obtained from conventional petroleum refining may also be included. These source materials thus are characterized as products which have a boiling fraction of less than about 850"F.

However, their component parts or reaction products (from reactor 110) which have a boiling point above 850OF are best treated in a parallel path as disclosed herein.

The further treatment in the second parallel reaction path will produce products of outstanding and improved viscosity, improved hydrogen content and improved chemical structure (e.g.

products are of a higher paraffinnic content with high K values, or aromatic polycyclic compounds which can still thereafter subsequently be treated to produce the desirable, less olefinically unsaturated cyclic compounds and linear or branched paraffinnic products). Hence, the gaseous products obtained by practicing the process herein and the conversion of the unsaturated products to more saturated products is a further contribution to my process previously described.

In hydrotreating the products of this process, in order to reduce the bromine number and to further reduce the nitrogen and sulfur content of said products, it is necessary to keep in mind that the lower temperature hydrotreating aspect does not appear to "crack" the product to any significant degree and thereby form smaller molecules, which in turn would reduce the temperature of the boiling point range. Hydrotreating of the product will render further upgrading, i.e. by cracking (as determined by a lower temperature boiling point range such as the formation of smaller average molecular sizes), more difficult due to the apparent inability in this process to crack paraffins.

The hydrotreating should, therefore, be the final part of the process as disclosed herein and not be used until the desired molecular size has been achieved by the "cracking" aspect of this process. Thus the vacuum residua feedstocks are initially cracked into smaller molecular size, under less severe conditions, and then the refractory residues, i.e. feed stream 11 2a, under the necessary but appropriately severe conditions which do not affect the mainstream products. The resultant product has a lowered temperature boiling point range. Accompanying this molecular size reduction is a reduction in the sulfur and nitrogen content of the top product compared to that of the feedstock.

The resulting smaller molecules which form the "cracked" products do not contain the high amounts of nitrogen of the feedstock. Nitrogen is normally concentrated in the heavy "asphaltene" end of the petroleum residua. Metals are bound through the nitrogen bond to the methyene groupings of the porphyrin structure of the asphaltenes, and are thus usefully separated in the parallel reaction path. Therefore, the metals and the nitrogen remain with the unreacted central porphyrin structure of the asphaltenes and are either not present or present in very limited quantity in the overhead distillate product derived from the vacuum residua containing the asphaltenes, which have been concentrated relative to the asphaltene content of the crude which formed said vacuum residua.The metals and the nitrogen are more concentrated in the No. 1 bottoms, e.g. 112a, which do not form overhead distillate product.

The sulfur content is not necessarily concentrated in the asphaltene portion of the petroleum crude or residuum. The resin portion of the residue may contain as much as 5%, by weight, of sulfur.

With respect to the catalysts which are sometimes also called reagents, these are employed in the form such that these produce the desired results, namely sufficient cleavage and/or scision of the bonds in the first and further parallel stream reactors or the hydrogenation ultimately desired. The results are obtained by first appropriately severely, but not overly severely, treating the first treated product with the further herein described catalysts, and these will be described with reference to the results achieved. Secondly, the pre-treated (in reactor 110) refractory components are thereafter appropriately severely treated in the parallel path. Mixtures of the catalysts are also suggested as highly desirable for the crude oils and their residues, or for mixtures of crude oils based on the various crude oil component fractions.

A suitable severe catalyst for reactor 110 is designated as Catalyst A, and it is prepared as follows. A mole of potassium hydroxide is dissolved in either ethanol, methanol, particularly best in ethanol, or an ethanol-methanol mixture or less advantageously, because of lesser solubility therein, in 1-propanol or 1-butanol. Solubility of the catalyst product is lower in these last two and larger amounts of the alkanol must be used and subsequently separated. The alkanols may be absolute alkanols, although these may be such as 95% ethanol. The potassium hydroxide is dissolved in this solution and is then reacted with hydrogen sulfide bubbled through the solution. After thorough saturation, the catalyst-alcohol mixture is recovered and the alcohol separated by vacuum aspiration. The residue is the catalyst.For a mixture of solvents, typically one mole of potassium hydroxide is dissolved in 200cc of ethanol and 1 30cc of methanol.

Typically, analytical reagent grade pellets of potassium hydroxide (about 86% KOH), absolute ethanol or 95% ethanol and absolute methanol are used. As mentioned above, the proportions of an ethanol-methanol mixture may be varied. The solution is evaporated under vacuum until no more residual alkanol can be removed. These catalysts appear to be best suited for cleavage, and especially for cleavage of the refractor components and, therefore, are used in the appropriate proportions (or solely) with others to achieve the desired cleavage.

Catalyst may be unsupported in the first stage reaction, but most advantageously for the first or subsequent stages it is deposited on supports and calcined.

Turning now to the catalyst supports which have been employed, these have been employed mostly for the purpose to obtain increased surface area. The catalyst supports are spinels and such as chromite spinel (CrO) and, most advantageously, porous metal, i.e. stainless steel of the available AISI grades, and the like. The last are obtained by sintering very fine sized uniform, powdered metallurgy particles or are produced as thin plates obtained by leaching out leachable constituents in the thin, (e.g. one eighth of an inch) metal plate, providing thereby intercommunicating passages. Other metal supports are such as are obtained by sintering very fine wires, about 0.2 to 5 mm thick, and cutting these to length, e.g. 2 to 5 mm.Still other supports are such as alumina with sizes of the pores ranging from 50"A to 350 and even up to 1,000An, but these may need to be protected as further explained herein. Although the treatment in subsequent reactors may be less demanding based on support characteristics, the treatment in the original first stage reactor in accordance with this invention is best carried out with a strong, inert support such as the porous metal supports which have a size range of the pores, e.g. up to 3,500"A and larger, i.e. the metal may be from 10% metal and 90% the pores, by volume, although metal may be up to about 25%, by volume.

In all the reactions the catalyst is allowed to react with the exclusion of atmospheric oxygen and thus in absence of oxygen. Similarly, the deposition of the catalyst on the support is in absence of oxygen as is the driving off of the volatiles from the support.

The catalyst and the support, after the volatiles have been driven off, are heated to an appropriate temperature such as between 320 and up to 450 or even up to 560. The catalyst tightly adheres to the support and may thus be used such as in a spinning-cage (also called spinning basket) reactor, ebullating bed or fluidized bed reactor.

If the support is unduly attacked by the catalyst, such as alumina in the first stage reactor, then the following method is used. The above catalyst is evaporated to considerable dryness, dissolved in glycerol, and the gylcerol-catalyst mixture deposited such as on an alumina support.

Other less resistant supports are treated similarly. Typically these molecular sieves may be of the Y and X, e.g. YL-82, type, with low sodium content (available from Union Carbide, Danbury, CT. or comparable supports from Mobil Oil, New York, NY). The molecular sieves function, however, as supports for the catalyst, i.e. to increase the contact area for the catalyst.

The glycerol catalyst mixture after depositing on the reagent is then progressively heated such as up to 560"C with the volatiles being driven off.

Glycerol may also be first deposited, heated up to about 200'C. and then the catalyst deposited thereon (after the support has been cooled), and then heated to the desired temperature.

The reaction in the first reactor may be at a higher temperature, and may range from about 320"C to about 450"C although temperatures up to 560 have been used, even up to 650"C.

For asphaltenes, the preferred temperature range is from about 360"C to about 430'C; it appears that between 390"C and 425"C is a very good operating range.

Inasmuch as for these catalysts the reaction must at all times be conducted in the presence of steam to facilitate the hydrogenation, hydrocracking, etc., steam is used in a ratio such that it is, at minimum, about 27%, by weight, based on the weight of feed such as the ORA fraction, crude, the residues, etc., charged to the above reactors. Conversely, the amount of water charged in the form of steam at the operating temperature may be increased or diminished based on the degree of hydrogenation desired (which also may take place in the first reactor to a certain degree). If more hydrogenation is sought to be achieved, more steam is being introduced, but typically steam does not exceed about 85% weight percent of feed, although 130%, by weight, may be used based on the hydrogenated product being obtained, i.e.

withdrawn (if gaseous fraction is being produced then it is converted to a liquid equivalent).

Stated on another basis, the amount of water used is determined by subtracting the hydrogen content of the feedstock from the hydrogen content of the desired product, on weight basis and multiplying the amount by 9 (as water is 1/9 by weight of hydrogen). Typically, up to a 30% excess is injected in the first reactor.

If water is not being introduced in the reactor, such as in the form of steam, or is interrupted for one reason or another, then cooking is apt to occur; thus carbon is being deposited or generated by a process somewhat similar to catalytic thermal cracking, but in this event the catalyst acts a thermal cracking catalyst, albeit with some advantage (because this catalytic thermal cracking is at a fairly low temperature, e.g. 320"C), but vastly less efficiently than when it functions in the presence of steam as a hydrogenation and/or cleavage catalyst. Hence, as previously mentioned slight cooking may also be taken as a lower, although less desirable, limit.

Intermittent or insufficient steam introduction will also cause production of especially heavy product in the reactor 110. It is important that steam is introduced at all times, in a proper manner in the reactor and thoroughly dispersed (without any steam and/or reagent free space).

Nevertheless, it must be mentioned that excessive amounts of steam also prevent the reaction from being carried out appropriately, apparently by unduly entraining the partially converted products.

If carbon is being laid down for one reason or another, typically it is on the hot spots such as heated reactor walls or catalyst support. Hence, the reactor 10 is preferably operated adiabatically. Carbon deposits on the catalysts can be driven off, that is, converted back into useful product by exposure to steam for a period of time without introducing additional feed, after which the catalyst is useful again and can be used for the production of the desired product cut. Intermittent introduction of hydrogen sulfide or sulfur may be helpful in general and for low sulfide content feedstocks, e.g. sweet crude oils.

Still further, it has been found that if the temperature, such as with the above catalyst, is increased to about 320"C to 420"C, depending on the feedstock composition (for example its asphaltene composition) an exothermic reaction may take place, e.g. at 440"C an exothermic reaction sets in. The exothermic reaction may reach temperatures up to 600"C, but it also depends on the amount of steam being introduced. More steam would tend to produce lighter carbon products. Excessive temperatures are not desired, and temperatures below 440"C are preferred.

In the second stage reactor in Fig. 1, 17, or 11 7 in Fig. 2 where further reactions take place, advantageously the products from the first reactor are rapidly cooled to about 250"C and in the presence of a catalyst, generally between temperatures of 250"C and 390"C. Quality of product is increased when the process is operated at temperatures up to about 430"C, preferably 425"C, but the conversion apparently will not increase after 390"C. Cooling in quencher 1 6a is at such a rate that steam does not condense and interfere with the reaction. The light ends, of course, that is hydrogenated products, will not be condensed.

The catalyst in the second reactor 1 7 or 11 7 again is preferably a supported catalyst but it can be an unsupported catalyst, and the treatment of the vaporous or gaseous products. A typical catalyst for the second stage reaction, i.e. reactor 1 7 or 11 7 is a less severe Catalyst B.

This catalyst is produced by dissolving a technical or analytical grade of a potassium hydroxide which is approximately 86% potassium hydroxide in absolute or 95% ethanol or methanol (preferably ethanol) and saturated with hydrogen sulfide but without boiling off the alkanol; collecting and trapping any alkanol given off in a downstream vessel. Other vessels in which the reaction takes place may be further downstream to catch the hydrogen sulfide. When the last vessel containing KOH shows a reaction, the reaction is stopped in all upstream vessels.

If the reaction is carried out in a further reactor(s) 1 7 or 11 7, i.e. second stage reactors, the advantages of the process reside in the combination obtained by the immediate quenching of the reaction products from the first stage reactor to about 300"C but preferably 250"C, in the presence of catalyst, and then conducting these reactions of the first stage products in the second stage. For this purpose, it has been found especially advantageous to support the catalyst on a suitable support. These supports may be the same as in the first stage, but in any event these supports must be inert under the reaction conditions in the particular reactor, e.g.

Fig. 1, 17, or Fig. 2, 117 and 130. These second stage reactors 1 7 or 117 may be used as fluidized bed (circulatory fluidized bed, partially circulating or confined fluidized bed), fixed bed or liquid bed reactors. Reactor 30 may alsp be an ebullating bed reactor, or a fluidized bed or circulating fluid bed reactor.

It has been found acceptable for the second stage reactors, e.g. 1 7 or 11 7, to use the supports of a type commonly available such as alumina-alumina silicates of a fixed zeolite type, i.e. molecular sieve type, with sodium or potassium in the zeolite exchanged with ammonia.

Type X and Y zeolites (10 and 13) are suitable. Type Y molecular sieve zeolites are preferred; of these, the low sodium ratio sieves are especially desirable (i.e. about less than 1.0% Na2O). The molar ratio of silica to alumina of these is about greater than 23 to 1; about 5 to 1, etc.; Na2O is about 0.2 weight percent. These are available such as from commercial sources, in forms such as powder spheres, cylindrical and other extrudates, etc., of suitable size such as 1/8 of inch extrudates or spheres. Although these have been alleged to be poisoned or destroyed by alkali metals, as worked up by the below-described procedure, these supports are useful despite the deposition thereon of the herein described alkali sulfide reagents. These supports may also be used in the first stage reactor 10.

Other zeolites are ELZ-L zeolite of the potassium type as described in U.S. Patent 3,216,789, and silicalite material as described in U.S. Patent 4,061,724. The last has a pore dimension of about 6 Angstrom units. Other supports are such as those described in British Patent 1 , 1 78,1 86, i.e. the very low sodium type less than 0.7 percent, by weight, e.g. ELZ- -6, or ELZ-E-6, E-8, or E-10. Other supports are mordenites and erionites with very low sodium content obtained by ammonia exchange and of the calcined type. Of the above molecular sieves, the type Y very low sodium, e.g. 0.1 5, by weight, ammonia exchanged supports available under Trademark LZ-Y82 from sources such as Linde Division, Union Carbide Corporation, Danbury, CT, Mobil Oil Corporation, New York, NY, and other sources are preferred.In any event, the stability and durability of these molecular sieves used as supports are tested under the reaction conditions and are established by the performance in the second stage reactor.

The preparation procedure for the second stage supports is as follows. The low sodium ammonium exchanged zeolite extrudates, such as powders, or of shapes such as cylinders, saddles, stars, rings, spheres, etc., of powder, or extrudates of about 1/8 to 5/32 or 3/1 6 inch size are treated with glycerol or like polyhydroxy alkane compounds, such as partially reacted polyhydroxy compounds including up to hexahydric alkanes, by first impregnating these shapes in a reactor which is kept closed. Thereafter, e.g. when using glycerol, by heating and removing decomposition products from these powders, or shapes, from room temperature up to 265' to 280' and even up to 560"C, an appropriate, but unknown, reaction takes place.The thus reacted support is then screened, drained, and cooled in a closed and tightly sealed container if the temperature has been brought up to 560"C.

When cold, the above-described support is then impregnated with a reagent-catalyst of the general formula K2S1.5 (empirical); this catalyst is acceptable, but it is not outstanding for cleavage. Although the catalyst is designated by an empirical formula, a particular preparative method determines the catalyst behavior. Consequently, while various catalysts may be represented by the same formula (empirical), vastly different properties are observed and different feed stocks being treated will produce different products. This catalyst, designated as Catalyst C, is obtained by dissolving 6 moles of KOH in 4 1/2 to 7 1/2 moles of H20 without external heat being applied, and thereafter with a small amount of alkanol, e.g. 2 to 2.5 cc of methanol or ethanol being added per mole of KOH.Then 4 moles of elemental sulfur are added to the foregoing solution which react exothermically. Thereafter, an appropriate amount of sulfur is added for adjusting the catalyst to the desired sulfur level by addition of supplemental sulfur to form the empirical sulfide, i.e. from K2S but K2S,, to K2S25, including up to K2S5 is useful, depending on the desired product cut. For more gas in the product, less sulfur saturated species are used. For more liquids in the product, more sulfur saturated species are used.

Another catalyst, Catalyst D, is prepared as follows. One mole of KOH is dissolved in 1.0 moles of water with vigorous stirring. Then 2 ml of methanol or ethanol are added immediately after KOH has dissolved. Immediately thereafter 2/3 moles of elemental sulfur are added and are allowed to react by a vigorous reaction. The catalyst is adjusted to the desired empirical sulfur content by adding appropriate amounts of sulfur by further stirring, e.g. one quarter of 2/3 moles of sulfur adds 0.5 to the empirical sulfur content of K2A; i.e. 1/4 of 2/3 moles dissolved sulfur gives K2S1.5; 1/2 of 2/3 moles gives K,S20, etc., including other appropriate fractions. Thus the catalyst may range from K2St 1 to K2S2.5 or even up to K2S5. The lower sulfur content species, e.g.K2S1.5, are good cleavage catalysts when admixed with Catalyst A, and is a good catalyst in general for the less severe treatment of crudes. This catalyst is also a hydrogenation catalyst. Catalyst A is generally admixed from 3% to 25% but typically less than 10%, by weight, to Catalyst D.

When the catalyst has been thus prepared, it is vacuum evaporated to a flowing slurry. It is then poured over the glycerol treated, cooled extrudate as described above (i.e. if the support had been heated up to 300"C. or higher), and under very low vacuum, agitated and aspirated until dry. Then the catalyst is further screened when dry and introduced immediately in the second stage reactor which has been purged of air oxygen when used as hydrogenation catalyst.

As another method for protecting the support, if the glycerol treated support is heated between 260'C to a decomposition point (indicated by appreciable slowing down of a liquid condensate being collected), then the above described catalyst slurry is added and the vessel is covered and heated up to at least 440"C, including up to 560"C. In any event, the catalyst is calcined above the temperature at which the catalyst is used in the process.

Still another method is to mix the glycerol, e.g. about 88 ml of glycerol, to about the one mole (K basis) of the catalyst, in solution, or by admixing the above catalysts or mixtures thereof to glycerol. Then the catalyst-glycerol mixture is heated to drive off water and/or alcohol leaving a glycerol solution of the catalyst. Temperature is brought up to 190"C for the foregoing. The mixture is then poured over the support and with agitation brought up to at least 450"C and even up to 560"C. This supported catalyst gives a very unpleasant odor. It must be prepared under well isolated conditions.

In use for a gallon-sized first stage reactor in conjunction with a second stage reactor, up to about 2/3 of mole of supported catalyst (K basis) is charged to the second reactor. As an example, alumina supported K2S 5 (empirical) catalyst may be charged to the second stage reactor.

Another catalyst, Catalyst E, is a nonsupported or supported catalyst capable of decreasing the molecular size of the product in a first stage reactor (or used in a further second stage reaction).

Catalyst E is obtained by adding a dried KHS powder or slurry in appropriate mole or weight percent increments (based on the desired size of the product) to any of the above-described reagent mixtures A, B, C or D. Either unsupported or supported forms of the catalyst may be used. That is from 1/5 to 1/3 moles on molar basis of K of KHS is added to e.g. K2S (empirical) sulfide, or to K2S15 (empirical), and the molecular size of the product is decreased by these additions of KHS.

When the process is run with the thus supported catalyst in the second stage reactor, appropriate adjustments may be made, e.g. K2S1.1 or K2S15 give more hydrogenation, and K2S2 gives larger molecules (also more distillate, less gases). These reactions are run in a temperature range from 113"C to 440"C. Similar catalyst adjustments may be made in other reactors, e.g.

when more than one second stage reactor 1 7 or 11 7 is used.

In any event, the first stage reaction, however, is carried out with the specified reagent to accomplish the desired degree of cleavage of the refractory, intractable initial source material components, e.g. residues of crude oil, the ORA fraction, and especially asphaltenes. In the parallel treatment then these not yet reacted intractable compounds are then treated catalytically to accomplish the desired degree of cleavage. The total process combination in the further stages, that is second, third, fourth, etc., stages, depends on the specified first stage reaction and removed portion of unreacted constituents and is thus interdependent.

The amount of catalyst deposited on the support is from about 4M of catalyst (K basis) to about 0.5 M or as low as 0.1 M (K basis) per 500 cc of support. On another basis, not identical basis, the amount of catalyst is about 20 grams per 300cc of support, but it may range from 3 grams or 5 grams of reagent per 1 OOcc of support to about 25 grams to 30 grams per 1 OOcc of support.

The vanadium metal may be entirely removed from the petroleum feedstock and the heaviest product may contain essentially all of the vanadium. If run with the presently supported catalysts, the vanadium is analyzed in very small quantities or as non-detectable in the No. 2 bottoms, i.e. product 1 8. The removal of the nickel is aided if some of the reagent be present in the hydrosulfide form. There is no reaction between the alkali metal sulfides and nickel sulfides but there is a solubility reaction when alkali metal hydrosulfide and nickel sulfides are present.

Nickel (and iron) form complexes like ferrites with the alkali metal sulfides-hydrosulfides. These complexes are immediately hydrolyzed in liquid water to form the precipitates of iron or nickel hydroxides. In liquid water, the vanadium complex with the catalyst is highly water soluble and water stable. Iron is normally present in the residue, after distillation range determinations, in amounts between 3 and 5 ppm, but the amount depends also on the amounts in the initial stage.

In general, the catalysts for the second and further stage reactions herein are the hydrosu Ifides and sulfides, that is, monosulfides and polysulfides of the Group IA elements of the Periodic Table other than hydrogen prepared from the alkanol solution as mentioned above.

Although for the stated purpose sodium, potassium, rubidium and lithium may be used, far and away the most advantageous are sodium and potassium. Of these two, potassium is preferred.

Although rubidium compounds appear to be acceptable, rubidium, the same as lithium, is not cost-advantageous. However, for the first stage reactor, rubidium may be very advantageous in a blend of rubidium, potassium and sodium, in the following proportions: 14% rubidium, 26% potassium, and 60% sodium sulfides, i.e. the various species thereof, on basis of the elemental metal, by weight. The ratio ranges for the preceding mixture are 1:1.5 -- 2.5:3.5 - 4.5, respectively, but these compositions must be prepared in the manner as defined according to the procedure described for Catalyst A. The catalysts used are typically used as the hydrates, but a small portion of the catalyst is apparently in the form of alkanolate (the hydrate analogue), i.e.

up to about 15% but typically less than 10% or even less than 5%, by weight. Alkali metal thionates are also present as transitory intermediates. Hydrates (and alkonolates) of these compounds are very complex and undergo a number of transitions during the start-up and as these achieve the reaction conditions. It appears that the presence of mixed hydrates and alkanolates are necessary for the outstanding results. No attempt has been made to elucidate the nature of these transitions or the actual structure at the reaction conditions for the sulfides, hydrates, alkanolates, thionates or the mixtures of each.It is sufficient to indicate, however, that charged catalyst can be mixture of a number of hydrates and/or alkanolates or a eutectic mixture of various hydrates and/or alkanolates, but in any event the specific cleavage catalyst must be used on the refractory compounds.

Similarly, during the reaction, as there is interconversion of the sulfur-containing forms of the sulfides, no attempt has been made to characterize this interconversion. However, as mentioned before, if the first stage feed contains sizeable amounts of the refractory components, i.e.

asphaltenes of various sizes, resins, and oils such as slurry oils, i.e. more than 50%, by weight, and up, but including up to 95% of a fraction boiling above + 1000OF, the first stage reaction requires at least a portion of the specific catalyst, e.g. Catalyst A, as defined above. As the percentage of the more refractory compounds in the feed decrease, less of Catalyst A needs to be used. Similarly for feed 12a, Catalyst A is again recommended. It may also be added to other catalyst compositions depending on the desired bond scission (as characterized by API number, viscosity, etc.), as that composition appears best to achieve bond scission or cleavage without excessive gas production.For the No. 1 bottoms, i.e. the feed stream 12a, the preferred reaction vessel or the stirred cage or basket reactor, further a fluid bed reactor and still further a fixed bed reactor. Another catalyst for feedstream 1 2a is Catalyst B with 5%, by weight (on elemental K basis) of KHS admixed thereto or replaced in the same proportions including up to 95% of Catalyst A. In general, the order of the catalysts for bond scissions or cleavage of the refractory compounds is as follows: Catalyst A (ethanol prepared species); Catalyst A (methanol prepared species); Catalyst A (methanol-ethanol prepared species); Catalyst A + B (mixture 50% or more Catalyst A on K basis moles); Catalyst D; Catalyst D (about 1/3 moles of KHS added); Catalyst B, and lastly Catalyst C.Of course, increasing the amount of KHS additions to the catalysts will increase the cracking ability.

The order of the catalysts for hydrogenation is as follows: 1 /2 Catalyst A + KOH on molar basis, then add the other half of A (very exothermic hydrogenation reactions); Catalyst C with KHS mixture (based on needed degree of cleavage); Catalyst B + A (50% or more of Catalyst B on K mole basis), and Catalyst B.

It must be remembered that in the above discussion, the refractory nature of the carbonaceous source materials and their components is that assigned to these according to the prior art.

By comparison, the present provess is carried out at lower temperatures and lower pressures and is capable of converting all of the prior art refractory materials with great advantage and facility.

However, as the carbonaceous source materials and their components have a different "refractory" nature according to the present process, e.g. the paraffins are the most refractory while the asphaltenes and aromatics are not; the invention resides in the ability to treat, in a proper sequence, according to this process, with the proper catalysts, all the components of these carbonaceous materials as characterized by heretofore unachievable yields when treating these carbonacous source materials.

In the following Examples, various reactions are described. There is no intent to limit the invention by the Examples but merely to illustrate its applicability.

EXAMPLE 1 A high softening point asphaltene 270OF as described below and of a solvent extracted type was treated with the following reagent to obtain product A. The catalyst was Catalyst A previously described. When the product from the first stage treated asphaltenes were reacted in a second stage, the product was identified as B. The second stage catalyst was the same as in the first stage. Both catalyst compositions were unsupported.

Residue Distillate Distillate Blend of A + B Feed A B A+B 600OF Gravity, 'API @60tF -4.6 31.3 36.7 36.0 8.7 Kin. Visc. @210 F, cSt -* - - 0.94 58.0 Con. Carbon Res.Wt% 39.5 - - 0.20 16.4 Aniline Point, "C - - - 44.6 FIA, vol% Aromatics - - - 71.5 Olefins Saturates - - - 28.5 Bromine No. - - - 53 Carbon, wt% 84.24 - - 83.77 84.81 Hydrogen, wt% 8.50 - - 12.52 9.95 Nitrogen, wt% 0.75 - - 0.10 0.58 Sulfur, wt% 6.19 - - 2.31 4.46 Ash, wt% 0.30 - - 0.08 0.20 Moisture, wt% Nil - - 0.05 Nil Oxygen, wt% 0.02 - - 1.17 0.00 Nickel, ppm(w) 71 - - - 33 Vanadium, ppm(w) 174 - - - 160 Iron, ppm(w) 151 - - - 24 Heptane Insoluble 16.7 - - - - (IP Method) *270 F-Softening Point ≈Aromatics and olefins were not clearly separated in the column probably due to a heavy tail above 600OF.

The above data clearly indicate the considerable improvement in the viscosity as well as the gravity of the products, the dramatic increase of the hydrogen content and the considerable removal of the metals present from the later fractions.

The following examples show the results obtained. The feedstock charge material was solventextracted asphaltenes of the type identified above and in Example I herein.

All of the runs were made as batch-process runs, in a stir-tank reactor. The stir-tank reactor has an inside volume of 6.24" diameter and 10" height. The stir-tank reactor is fitted with an agitator and a steam sparger.

A steam generator, directly connected to the city water supply, forms steam at 40 Ib/sq. in.

pressure. The steam passes through 3/8" inside diameter lines to the sparger and is at atmopsheric pressure. However, the reactor may operate from 1/2 atm. to about 5 atm. or even higher as previously discussed.

The sparger is approximately 3+" in diameter and has a series of sparger holes, all of the holes direct the steam upwardly. The sparger is located on the bottom of the reactor.

An agitator is provided for the reactor when unsupported catalyst is used. The motor is mounted directly above the reactor. A seal seals the area through which the agitator rod passes into the reactor. The agitator is of twin circles connected by angled, curved blades. The agitator may be replaced by baskets holding supported catalyst as further described herein. Four baskets containing supported catalyst are mounted on the agitator shaft. The baskets plus the agitator shaft has a total diameter of almost 6.25". The baskets are 6" high and are approximately i" thick. The unmounted basket is a i" deep rectangle.

The top of the reactor contains the seal through which the agitator shaft turns, the riser, which exits overhead distillate from the reactor, a pressure relief line, which consists of a valve which opens above 30lbs/sq.in. pressure and vents the contents of the reactor to a hood. This pressure relief fitting is also used to fill the reactor with solid feedstock charge.

Usually, two thermocouples are fitted into the top of the reactor. One thermocouple measures the temperature at the bottom half of the reactor and the upper thermocouple measures the temperature in the upper half of the reactor.

The riser consists of a line approximately 9" high and having an inside diameter of approximately 3/4".

The second stage reactor is a tube reactor having an inside diameter of 1 +" and is 12" long.

The capacity of this reactor is 347.55cc. This reactor is fitted with three wrap-around heaters. A thermocouple controls each of the heaters through the controllers, which are mounted on a portable stand.

Gases and vapors passing the up-flow second stage reactor are then conducted downward through a 16" glass bubble condenser. This condenser is not water cooled.

The first condenser is mounted vertically and the bottom of the condenser holds a 500 cc collecting flask. The flask is normally maintained at 240"C by a mantel type heater. The bottom of the flask has a stop cock for collecting product.

A second condenser rises from the above flask and is parallel with the first condenser. The second condenser is not cooled by water. The second condenser is also a glass condenser with bubble type cooling areas.

The downward slanting tube from the second condenser connects to a water-cooled condenser. This water-cooled condenser is mounted vertically and is approximately 18" long; it fits into the top of an unheated 500cc flask which is fitted with a stop-cock at the bottom. A parallel vertical water-cooled condenser rises from the second fitting in this flask. Another watercooled condenser is fitted directly above this condenser.

The top water-cooled condenser is fitted with a 12" long line which angles upwardly. This line has a diameter between i" and 3/4". This line connects to an ice cooler.

The ice cooler is a twin wall vessel, ordinarily used to trap vapors before they can enter a vacuum pump. The center container contains a water-ice mixture while the gases and vapors pass through the external container section. The gases and vapors enter at the bottom of the vessel and exit at the top of the vessel. The bottom of the vessel contains a 50cc collector. The collector is fitted with a stopcock, for removing product.

The remaining gases and vapors are sent to another cooler, similar to the ice cooler. This cooler is cooled by a mixture of solid carbon dioxide and 2-propanol. The product is again collected in a 50cc vessel below the cooler and this vessel is fitted with a stopcock.

The remaining gases and vapors are then washed in a solution of potassium hydroxide. The solution contains 6 moles of KOH dissolved in 360 ml of water. The gases are then measured by passage through a wet test meter. After this measurement, samples are periodically collected and the uncollected gases are vented to the hood.

For the following runs, 1 300 grams of the solid asphaltenes are weighed out and crushed to size to charge the reactor. Liquid or solid catalyst protected from oxygen is then added to the reactor, previously purged, e.g. with helium. About 40 grams of theoretical anhydrous Catalyst A is charged to the reactor.

The secondary reactor is charged, again with the same precaution, usually with about 300cc of a supported catalyst. The secondary reactor is initially heated, in order to drive out the water content of both the zeolite support and the catalyst.

After the water content of the secondary reactor space has been reduced by bringing the temperature of the second rector to above 300"C, the primary reactor is heated.

Solvent-extracted asphaltenes having melting points of either 200 or 400OF were used. The melting point determines the particular form of the asphaltenes.

After temperature adequate to melt the asphaltenes were reached in the primary reactor, the agitator was turned on. Normally the agitator is initially operated at approximately 30 to 60 rpm.

Steam is normally introduced when the primary reactor reaches a temperature of 220'C. By this time, the second stage reactor should have reached or leveled off to 424"C.

Helium is normally sparged through the sparger prior to the introduction of steam to the system in order to keep the sparger holes clear and the system free of oxygen. The helium is sparged at approximately 200 cc/minute.

EXAMPLE II 1 300 grams of solvent-extracted asphaltenes were reduced in size so that they could pass through the 1" opening in the top of the reactor. The asphaltenes were not heated but were charged to the reactor in solid form. After the charge, the reactor was sparged with helium.

The catalyst used was another version of Catalyst A prepared as follows. To the previously described initial solution of KOH was added a solution of one mole KOH dissolved in 30 cc of H20 and then the solution mixture saturated with hydrogen sulfide. The solution separates in two layers about 1/3 top layer and 2/3 bottom layer. The layers are separated and dried and then the two proportions reblended. The reblending may be in the same proportions as obtained (as it was in this Example), or the proportions of the two catalysts may be varied. The catalyst may also, upon reblending, be dissolved or dispersed for deposition on a support. On a theoretically anhydrous condition, the weight of charged catalyst was approximately 40 grams.

The second stage reactor had been charged with zeolite supported catalyst during apparatus assembly. The second stage reactor contained approximately 300 grams of support and catalyst.

The zeolite support was L2-Y82 and the catalyst was catalyst D, i.e. K2S,5 (empirical), to enhance the hydrogenation of the cleaved product. The start-up procedure requires that the secondary reactor be brought to at least 175"C before the primary reactor is heated.

The first stage reactor was then heated. Only the bottom one-half of the reactor is heated, the top half of the reactor is not heated. At 220 C, a small amount of steam was added to the reactor through the bottom sparger.

At approximately 320"C, in the bottom of the first stage reactor, there began a steady but slow production of hydrocarbon product, which was condensed in the flask below the twin water-cooled condensers. However, this product was much heavier than the product obtained at process-temperatures, in the 390'C to 424"C range.

When the bottom of the first stage reactor reached 360"C, there was a considerable improvement in the rate of product production. The reaction became exothermic and rose rapidly and leveled off at about 415"C. This temperature was maintained from that time forward in the bottom of the reactor. The top of the reactor had reached 360"C.

The temperatures in the second stage reactors are in the 220'C to 460"C range.

When the contents of the first stage reactor were in contact with the agitator, the process ran uniformly at about 415"C in the first stage reactor and with variations between 440 and 460"C in the second stage reactor.

The amount of steam was estimated at approximately 20 cc of water converted to steam/minute. At the end of the run, the top and bottom temperatures in the primary reactor were allowed to rise to 440"C.

Catalyst A variation above (as described above) gave almost no gas through the wet test meter. The amount of gas was less than 6 liters.

The bulk of product was the No. 2 bottoms, collected below the water-cooled condensers.

This product, when combined with the product collected below the water-ice trap totalled 458 grams. This product had an API number of 23 (sp.gr. @60 F 0.9158).

The No. 1 bottoms totalled 33 grams and had a gravity of 0.96587 (API number 15) @60 F.

The No. 1 bottoms were collected below the air-cooled condensers.

The amount of bottoms collected below the dry ice-2-propanol cold trap measured 44 cc in the calibrated trap. However, when collected, only 28 cc were obtained due to the evaporation of these light ends. The API gravity of these light ends was 81@ -- 10"C (sp.gr.

@ - 1OC = 0.6553). Due to rapid evaporation this gravity is very imprecise.

An extraordinarily light coke was formed and formed 2" thick layers in the reactor. This coke measured 1800 cc but had a weight of 51 3 grams.

The dead space below the agitator causes a delayed coking operation below the true reaction zone.

EXAMPLE 111 This example was carried out in a similar manner to that of Example II with the exception of a different form of the catalyst and a more rapid start-up heating of the first stage reactor.

Catalyst A for this example was a single layer catalyst and did not require the layer separation during the drying phase that was used for the catalyst used in Example II above. The sustained temperature of this run was 420"C.

In the 42 minutes of this run, after achieving process-temperature (at about 420"C), the catalyst, approximately 40 grams on a theoretically anhydrous basis, converted the following bottoms from the initial 1 300 gram solvent-extracted asphaltene charge: a) No. 1 bottoms, collected below the air-cooled condensers, totalled 29 grams of a hydrocarbon, having an API gravity of 11.5 (sp. gr. of 0.9895 @60"F).

b) No. 2 bottoms, collected below the water-cooled condensers and the water-ice condensers, totalling 398 grams of hydrocarbon, having an API gravity of 29 (sp.gr. @60"F = 0.8816).

c) No. 5 bottoms, collected below the dry ice-2-propanol cdld trap, totalled 63 cc with a gravity of 83 at ambient temperatures. A substantial part of the No. 5 bottoms were lost in determining this API gravity.

d) A total of 1 35 liters of gas were produced. The gases were measured following the alkali hydroxide wash of the gas-vapors (following the dry ice-2-propanol cold trap). These gases were not collected for analysis, but the average analysis of similar runs produced approximately 5% (volume percent) non-hydrocarbon gases, such as hydrogen, carbon monoxide and carbon dioxide. The remaining hydrocarbon gases have an average molecular weight of 49; on this basis a total weight of 280 grams of hydrocarbon can be assigned to the gas obtained.

e) The same light coke as formed in Example II was observed in the reactor following this run.

The weight of the coke was 489 grams.

The accounted for weights are: No. 1 Bottoms 29.0 grms No. 2 Bottoms 398.0 grams No. 5 Bottoms 41.5 grams (cold trap) Gases 280.0 grams Coke 489.0 grams Total 1,238.0 grams Accountability = 95.23% Conversion = Nos. 1,2,5, bottoms + gas/feedstock charge = 57.57% For both Examples II and Ill the supported catalyst of the second reactor was completely clean and free of pitch, carbon, etc.

The total accountability for products obtained by Example II was 1,045.7/1300 = 80.42%.

The conversion, i.e. the Nos. 1,2,5 bottoms + gases = 532.7/1300 = 40.97%.

The principal difference between the two Examples was the much higher gas production in Example Ill.

In these two examples, the bottoms completely separated from the water, condensed from the steam and no emulsion was formed.

EXAMPLE IV In this Example, a blend of catalyst was used, i.e. about 2/3 of the catalyst was that described of Example II, but not separated, 1/3 catalyst of Example Ill (K basis). The various proportions may be changed, including the proportions of the catalyst layers in Example II. The catalyst was unsupported and was about 40 grams on a theoretically an hydros basis.

The reactor reached a temperature of 420 during this run.

The products obtained during this run were: No. 1 Bottoms 195.9 grams (sp.gr. 0.9793 @ 60 F or API 13) No. 2 Bottoms 309.5 grams (sp.gr. = 0.86 @ 60"F or API 33) No. 5 Bottoms 28.0 grams Gas 276.4 grams (133 1iters X 0.95/22.4 X 49 = 276.4 grams) Coke 473.0 grams Total 1,282.7 grams Accountability = 1,282.7/1300 = 98.67% Conversion (Nos. 1,2,5 bottoms + gas = 809.7 grams/1300 grams) = 62.28%.

It was apparent that the coke formation was from the 480 cc of space below the agitator and some of this space is also the sparger.

EXAMPLE V The principal difference between this Example and the previous Examples was the use of a supported catalyst instead of unsupported catalyst being added prior to the beginning of the run. The catalyst was supported on stainless steel sintered mesh in four baskets. The stainless steel was 1 /8" thick and had been cut into 1 /8" strips which were in turn cross-cut for 3/16" sizes. The support size was therefore 1/8" X 1/8" X 3/16". The catalyst was the same catalyst as used in Example II.

The supported catalyst was placed in 1 /4" X 2.5" X 6" baskets, 4 baskets were used and the baskets were supported and turned by the agitator shaft. The baskets became the agitator. A mesh held the supported catalyst in place and the wire mesh baskets were supported by a frame.

With the same catalyst as in Example II (and considerably less of the catalyst in the supported form) the amount of gas produced decreased from 1 35 liters to 90 liters. Most of this gas was produced during the end of the run when the temperatures rose to 440"C.

The speed of the agitator which spun the baskets was initially 60 rpm and later in the run was increased to 1 20 rpm.

Gas 187.0 grams (90 liters of gas X 0.95/22.4 X 49 = 187.0) No. 1 bottoms 407.0 grams No. 2 bottoms 368.0 grams No. 5 bottoms 43.0 grams Coke 199.5 grams Total 1,204.5 grams The feedstock charge of the solvent-extracted asphaltenes was 1 290 grams. Accountability = 1204.5 grams/1290 grams = 93%. Conversion = Nos. 1,2, and 5 bottoms + gas / 1290 = 77.90%.

The API number of the combined No. 2 and No. 5 bottoms was 32.0 @ 60OF. Of the No. 1 bottoms, a division was made between the part which was liquid at 200"C and that which was not liquid at 200"C. The liquid portion had a calculated API number of 5 and this portion constituted 228 grams of the 407 grams of total No. 1 bottoms, or 56.03%. The remainder appear to be slightly upgraded forms of the solvent extracted asphaltenes. Recalculation of the conversion based on the liquid portion @ 200'C of the No. 1 bottoms give a conversion of 64%.

In all these conversions and accountability estimates, the gas produced is calculated at 95% hydrocarbon, of a hydrocarbon of an average molecular weight of 49.

It is evident that, when the feedstock charge is below the sparger, a competitive process is operational. It involves a decreased temperature catalytic thermal cracking with an improved threshold limit for thermal cracking, i.e. when steam is not present along with the catalyst.

When the catalyst is supported, there is no coke formation if steam, the asphaltene feed, and the catalyst are in intimate contact with each other. If insufficient steam, or no steam reaches the asphaltene feed, then the reaction turns into a catalytically aided thermal cracking. Although carried out at lower temperature then normal thermal cracking, at about atmospheric pressure, and at about a rate 10 times faster than normal thermal cracking, the catalytic hydrocrackinghydrogenation is vastly more desirable because of the high yields, high space velocities, desirable, adjustable product composition and the reaction conditions.

EXAMPLE VI In accordance with the above described procedure as it concerns Example II, a mixed residue feed containing various types of residues designated as FHC-353 was used.

Analysis for the feed is as follows: weight % of below 1000OF cut 8.6% API number 6.6 specific gravity at 60"F 1.0243 sulfur, wt. % 3.91 nitrogen, wt % 0.478 hydrogen, wt % 10.35 carbon, wt % 84.72 oxygen, wt % 0.471 Total elements: 99.9% oils, wt % 29.6% resins, wt % 57.8 asphaltenes, wt % 1 2.6 Recovery, wt % 100% Ramsbottom carbon, wt % 21 % Metals: vanadium 228 ppm nickel 52 ppm iron 14 ppm sodium 4 ppm viscosity 1430 cts at 100"C 250 cts at 135C pour point 130'F The residue feed was gradually heated up so as not to cause coking until the mixed residue became pumpable or reached the reaction temperature of about 425 C, at which temperature the reaction was conducted in the presence of catalyst and steam. Total feed sample was 15,650 grams, and recovery was 98.1% of the feed charge based on the various condensates obtained.

A boiling point redetermination of this feed residue according to ASTM D-1 160 and D-1 180 was also carried out on a 200cc sample of the feed in accordance with the described method and the data obtained were as follows (reported as temperatures observed during the determination): Sample % "C temp. at "C temp. Corrected temp.Pressure in cc Recovery top of column of pot (D-1160; D-1 1 80) column-mm. of Hg 40 20 240 275 975 4.2 50 25 253 286 810 4.2 60 30 263 295 835 4.6 70 35 273 304 858 4.2 80 40 283 310 882 4.2 90 45 290 316 4.2 100 50 297 324 907 4.2 110 55 305 330 120 60 312 338 938 4.2 130 65 320 344 953 4.2 140 70 326 351 4.2 160 80 319 347 990 2.0 170 85 330 2.0 180 90 339 364 1030 2.0 188 94 347 1050 2.0 347 t90 99 351 380 1090 1.0 This product was converted in a stirred basket reactor having a supported catalyst. The support was a Y-82 zeolite from Union Carbide, Danbury, CT in the form of a 1/8 inch extrudates, e.g.

flat shapes. The catalyst has an empirical formula of K2S1 5, i.e. prepared as catalyst D above, with 5%, by weight, of Catalyst A added thereto. The amount of catalyst was 20.1 6 grams per 220cc of the described support.

Initial conversion of product for first product (top product from reactor 110) was 71 % (for combined product except the very volatile, dry ice and iso-propanol cold trap materials and gases). This conversion also does not include the distilled material in the feed recoverable below 1050OF. The top products plus recovered bottoms were accounted for and 98.1% of initial feed charge, the remainder apprently being the gas and volatile distillates.

A 1,392 test sample containing 7 grams of H20 of top products of reactor 110 and 117 except for the condensates mentioned above had the following component distribution: Initial boiling point to 360OF 71.6 grams or 5.24 weight % 360OF to 650 F 203.4 grams or 14.89 weight % 650OF to 1000OF 1084.1 grams or 79.36 weight % Total distilled: 1359.1 grams.

The product of reactor 110 and 117 had an initial boiling point to 360'F product which had the following characteristics: API number 54.5 specific gravity .7608 at 60'F viscosity 0.72 cts at 100"F 0.568 cts at 150 F ash content .009%, by weight nitrogen content 0.05% sulfur content 1.0%.

The fraction boiling from 360OF has the following characteristics: API number 34.2 specific gravity 0.8540 at 60"F water 0.002%, by weight pour point - 20"F viscosity 2.57 cts at 100OF 1.61 cts at 160"F sulfur 2.18% ash 0.01%, by weight.

The fraction boiling to 650 F to endpoint had the following characteristics: API number 17.8 specific gravity 0.9478 at 60"F pour point + 115 F viscosity 37.34 cts at 150 F 12.05 cts at 210OF carbon content 85.03%, by weight hydrogen 11.55%, by weight nitrogen content .26%, by weight Conradson carbon 2.56%, by weight.

The recovered residue, i.e. remaining after + 1050 F stage had been reached had: vanadium 5.6 ppm iron 48.6 ppm sodium 8.2 ppm potassium 22.3 ppm nickel 0.48 ppm.

For the vacuum gas oil fraction, i.e. 650 F and above fraction, the characteristics were as follows.

Initial boiling point 625OF and recovery as follows: 10% at 732'F 20% at 782'F 30% at 832OF 40% at 878OF 50% at 894'F 60% at 928OF 70% at 948OF 80% at 976"F 90% at 1080"F end point at 1050"F at 94.4% recovery.

The K factor for completely treated top products from reactor 110 for the 650OF plus fraction (hydrogenation as a last step) was 11.71.

The bottom from the first stage reactor 110 recovered as 11 2a were rerun on a batch basis when sufficient amount were recovered. The catalyst for feed 11 2a was Catalyst A supported on a steel mesh particles of a size of 1/8 X 1/4 X 3/16, of total volume of 200cc which contained 15.7 grams of catalyst. Total sample run was 1286 grams. Total recovery was 750.7 grams (which included also 8 grams from the dry-ice trap but no gas). About 29% of FHC-353 feed were recovered as bottoms. Feed analysis of 11 2a bottoms was: carbon 84.78%, by weight hydrogen 8.74%, by weight ash 0.98%, by weight sulfur 4.64%, by weight Conradson carbon 40.57%, by weight About 58.4% of feed was converted.On a continuous basis higher conversion would be reached, but on batch basis the dead space in the reactor underneath the spinning cage or basket did not allow complete reaction space, as well as the holdup in transfer lines prevent completion of the reaction.

Combined recovered products had the following characeristics: carbon 82.24%, by weight hydrogen 11.95%, by weight oxygen 0.64%, by weight water 2.8%, by weight.

The 360 to 650OF boiling point fraction has the following characteristics: API number 27.2 pour point -1SF viscosity 8.533 cts at 1OO'F 4.028 cts at 150'F nitrogen .17%, by weight water 2.8%, by weight The fraction boiling 650"C to endpoint had the following characteristics: API number 17.4 specific gravity .944 pour point + 95"F viscosity 21.54 cts at 150e F 7.989 cts at 210"F carbon 84.7% hydrogen 11.27% oxygen .66% Conradson carbon 2.2%, by weight nitrogen 0.27%, by weight vanadium 3.5 ppm iron 7.6 ppm sodium 1.6 ppm potassium 10.3 ppm nickel .45 ppm.

Distillation residue: carbon 85.22%, by weight hydrogen 11.06%, by weight oxygen .49%, by weight sodium 3.07 ppm vanadium N.D.

nickel .344 ppm iron 4.6 ppm potassium 27.4 ppm.

The 650OF + product (from feed 11 2a) has the following disillation characteristics for a 191 gram charge of which 1 63 grams were overhead; it is a 1 75 ml sample at 60OF. The recovered fractions were as follows: 10% at 650OF 20% at 775OF 30% at 830OF 40% at 870OF 50% at 907OF 80% at 1030"F 85% at 1050OF This distillation and residue product had the following characteristics: API number of 20.0 at 60 F and 16.4, respectively.Moreover, the low metal content and the fairly high API number make this residue a good feed, e.g. as a catalytic cracker feed.

While the residue was about 15%, it would constitute the bottoms in the reactor vessel, e.g.

130, and these again may be subjected to further treatment, e.g. in a further parallel pass, i.e.

the same as for feed 11 2a. The analysis of the reaction vessel residue product was: carbon 83.99%, by weight hydrogen 6.88%, by weight oxygen 2.9%, by weight water .05%, by weight Based on the combined 58.37% (recovered product divided by weight of feed stock with accountability of 83%) conversion (rerun of bottom products from reactor 110 feed 112a) and the top products of reactor 110 but not counting in for the last the dry ice trap liquids and not the recovered gases for the top products of reactor 110 and not the supposed distillate products (supposedly still in FHC-353 feed distilling below + 1050OF), the total conversion is 87.9%; ; however, if gases and dry ice cold trap products are counted on an average molecular weight basis of 49 and the volume converted to weight, including the actual distillate content of FHC 353, then the actual recovery is even higher.

The yields in the above examples have been calculated on a basis to render the interpretation of these yields consistent with the conventional practice. The conversion on a space-time velocity may also be calculated, i.e. the starting material treated per unit volume of the catalyst for that starting material per a given unit of time.

The yields are conventionally calculated on the basis of the total amount of starting material used and converted with a fraction of a component boiling at a boiling point of less than + 1,050OF included. The boiling points are determined by the previously described methods, which are the conventional methods used in the petroleum industry. In petroleum industry practice the limit of + 1050OF is used to describe the very refractory crude or distillation product components according to the conventional methods.

The yields may also be based on a proportion of the components converted to a lighter viscosity product such as on a degree of API product basis. The residues such as collected by the cyclones 11 6 and used as feed 11 2a are calculated into the yields as additions to the initial amount of starting material treated and recovered. The residue from vessel 1 30 boiling at or over 1,050OF is then subtracted from the overall collected material provided this residue is not further treated.

In some practice the yields are based on the treatment of residues having certain boiling points, namely residues boiling more than at 650"C and less than 850"C or like limits, and then to consider only the convertible fraction by conventional processes (whatever that process is for the particular refiner). Inasmuch as that basis is not useful for evaluating the present process because the present process converts nearly all the previously unconvertible residues, it is believed best to characterize the products which are recovered in the straight through run and the first and second (if necessary) parallel run products distillable at less than 1050OF being additions.That is after the first run products and second or third parallel run products are added together including any further reaction vessel, e.g. 117, etc., including all condensed products, etc., and these are combined and a distillation is carried out as described and the residues boiling above + 1050OF are subtracted for the conversion calculations, the yields may then be determined.

For best practice, the gaseous hydrocarbon products in the feed distillable at less than 1050OF, the low temperature condensibles and gas are included. Gas can be included on a calculated basis such as average molecular weight calculated for the volume collected, converted to a weight percent to provide for accountability of all products.

According to the above description and pursuant to the Example VI as disclosed herein, it is believed that the specific treatment of the split or separated fractions from petroleum crude, low boiling point residues, up to 850'F, or high boiling point residues up to and over 1050OF, in a straight through fashion at low pressures and low temperatures allows the conversion of almost the entire crude oil or petroleum or other hydrocarbon values into usable products if a second or further treatment of the bottom fractions is practiced with the specifically selected catalysts.

Heretofore considerable effort had to be devoted to the conversion of the intractable portions of crude oil or petroleum residues into useful components, but the yields have been low and numerous recycles had to be conducted to achieve acceptable yields. If it is remembered that about up to 50% of these intractable portions had to be subjected to the very severe treatment, the present parallel method, which provides considerably higher yields and better end products, high space-time velocity, etc., while utilizing poorer starting product or separately treating the split product at less drastic conditions, then it is evident that the present parallel method stands out as a great contribution in the continuous search for obtaining all useful products from a given volume of crude oil, or its residue.

Claims (27)

1. In a process for converting carbonaceous starting materials, such as crude oil, residues thereof and the like as starting materials having refractory components therein, without addition of hydrogen gas as a reactant, to obtain products of lower viscosity and/or end products that are more hydrogenated, said converting being in presence of an alkali metal sulfide catalyst, the improvement comprising: a) reacting, in a first reaction zone in the presence of added water in form of steam and a catalyst, at least one carbonaceous starting material of a heavy crude oil, a natural asphaltic material, a natural tar, a pitch, a Gilsonite, a slurry oil, a solvent extracted asphaltene; a pitch or a tar derived from coal; a petroleum residue; oil, resin and asphaltene mixtures; or an oil-resinasphaltene fraction of a distillate having a boiling point of up to 850OF + and higher; a distillation residue having no boiling point below destructive distillation of same; a coal oil extract; a bottom fraction of retorted shale oil; a heavy bottom from coal gasifiers of SASOL type; a delayed coking product distillation bottom or mixtures of the foregoing, said catalyst comprising said alkali metal sulfide catalyst and an additional cleavage catalyst of at least 5% for pre-treating in said first reaction zone the refractory components in said starting material, said cleavage catalyst being in admixture with said alkali metal sulfide catalyst for treating said refractory components in said starting material; wherein said cleavage catalyst is a supported or unsupported catalyst composition comprising: 1) a first solution of an alkali metal hydroxide dissolved in methanol, ethanol, 1 -propanol or 1-butanol or mixtures of these alkanols, or 2) a second solution of said alkali metal hydroxide-alkanol as defined in 1) above to which water dissolved alkali metal hydroxide has been added and wherein said alkali metal hydroxide, on a mole basis, in said solutions is from 0.5:1 to 1:0.5, said first or second solution being saturated with hydrogen sulfide, such that in either solution.

i) a single phase solution forms, or ii) a two phase solution forms, said catalyst composition being said single phase solution of (i), said two phase solution of (ii), each of the phases of (ii), taken individually, mixtures of the phases of (ii) with each other, a mixture of each of the individual phases of (ii) with the single phase solution of (i), or a mixture of the two phases of (ii) with each other taken with the single phase of (i); b) recovering a top reaction product from said first reaction zone, including gases as vaporous or gaseous products and separating bottom products; c) recycling bottom products or separating as bottom products in a parallel stream from the products from said first reaction zone, said products being liquid products, entrained liquid products, or partially pretreated refractory components in said starting materials;; d) reacting in presence of initially introduced steam, or added steam, the separated bottom products in the presence of a said cleavage catalyst as defined in (a) above and said alkali metal sulfide catalyst, and e) recovering the products produced in step (d).

2. The process as defined in claim 1 wherein in step (a) and step (d) the catalysts are metal supported catalysts.

3. The process as defined in claim 1 wherein the products of step c) are recycled instead of separately treated.

4. The process as defined in claim 1 wherein the top vaporous or gaseous reaction products of step (b) or step (d) are reacted further, directly, or after cooling of same, in at least one additional reaction zone, for hydrogenation of same, in presence of a supported catalyst and added steam, wherein the supported catalyst comprises an alkali metal hydrosulfide, sulfide, polysulfide, a hydrate of a sulfide, a hydrate of a polysulfide, or mixtures thereof such that the bromine number is reduced compared to the bromine number in products of step (b) or step (d).

5. The process as defined in claims 1 and 3 wherein the top vaporous or gaseous reaction products of step (b) are reacted further, directly, or after cooling of same, in at least one additional reaction zone, for hydrogenation of same, in presence of a supported catalyst and added steam, wherein the supported catalyst comprises and alkali metal hydrosulfide, sulfide, polysulfide, a hydrate of a sulfide, a hydrate of a polysulfide, or mixtures thereof such that the bromine number is reduced compared to the bromine number in products of step (b).

6. The process as defined in claims 1 to 5 wherein the catalyst in step (a) or step (d) is a supported catalyst and comprises 1/ Catalyst A, or 2/ Catalyst A up to 95%, on K mole basis of Catalyst A, in admixture with Catalysts B, C, D, or E, or mixtures thereof, each of said catalysts being prepared as defined herein:: Wherein Catalyst A is prepared by dissolving a mole of potassium hydroxide in an ethanol, methanol, an ethanol-methanol mixture, 1-propanol, or 1-butanol, reacting the potassium hydroxide solution with hydrogen sulfide bubbled through the solution, recovering the catalystalcohol mixture and separating said alcohol from said solution;; Wherein Catalyst B is prepared by dissolving a technical or analytical grade of a potassium hydroxide of approximately 86% potassium hydroxide, in ethanol or methanol and saturated with hydrogen sulfide in a series of vessels by introducing in a first vessel said hydrogen sulfide, but without boiling off the alkanol, collecting and trapping any alkanol given off by an exothermic reaction between said potassium hydroxide and hydrogen sulfide in a downstream vessel, and stopping the addition of hydrogen sulfide reaction when the last vessel containing KOH shows a reaction; Wherein Catalyst C is obtained by dissolving about 6 moles of KOH in about 4 1/2 to 7 1/2 moles of H20 without external heat being applied, and thereafter adding a small amount of alkanol, of about from 2 to 2.5 cc of methanol or ethanol per mole of KOH; adding about 4 moles of elemental sulfur; adding an appropriate amount of sulfur for adjusting the catalyst to the desired sulfur level by addition of supplemental sulfur to form the empirical sulfide, from about K2S,1 to K2S5; Wherein Catalyst D is obtained by dissolving one mole of KOH in 1.0 moles of water, adding immediately 2 ml of methanol or ethanol after KOH has dissolved; adding 2/3 moles of elemental sulfur; adjusting the catalyst to the desired empirical sulfur content by adding sulfur by further stirring, said catalyst ranging from K2S1, to K2S5; Wherein Catalyst E is obtained by adding a dried KHS powder or slurry to each of the abovedescribed catalysts or mixtures of catalysts A, B, C or D; said addition being from about 1/5 to 1/3 moles on molar basis of K of KHS to K2S (empirical) sulfide to K2S25 (empirical), on molar basis.

7. The process as defined in claim 1 wherein the catalyst in step (a) or step (d) is a supported catalyst and the support is a porous metal, chromite spinel, zeolite or an alumina.

8. The process as defined in claim 1 wherein the catalyst in step (a) or step (d) is a porous metal supported catalyst and the said porous metal is a stainless steel of up to 35% metal by volume.

9. The process as defined in claim 6 wherein the carbonaceous starting material reacted in step (a) is a heavy crude oil, or a residue having no boiling point, or an asphaltene and the supported catalyst is at least 75% weight percent Catalyst A in combination with Catalyst D deposited on a catalyst support.

10. The process as defined in claim 9 wherein Catalyst A is from 5%, by weight, to.80% by weight, and balance is Catalyst D.

11. The process as defined in claim 6 wherein in step (d) the catalyst is a porous metal supported Catalyst A and wherein the reaction is at a temperature from 360"C to 440"C.

1 2. The process as defined in claim 1 wherein the reaction is carried out at a pressure from subatmospheric to less than 1 50 psi at a temperature from about 160"C to about 6000"C, in presence of a porous metal supported catalyst.

1 3. The process as defined in claim 6 wherein the catalyst for the further reaction is a Catalyst A composition with KHS as an admixed component of up to 15%, by weight, and the resultant catalyst product is deposited on a porous stainless steel support.

1 4. The process wherein the support for a catalyst for the reactions as defined in claim 1 is protected and wherein said support is a porous metal, chromite spinel, alumina, a zeolite, or mixed supports.

1 5. The process as defined in claim 1 4 wherein the catalyst is supported and the support is alumina of a pore size from 6A" up to 13,000An, preferably 350An to 900An, said alumina being protected from an attack by said catalyst by depositing said catalyst in an admixture with glycerol and calcining it with exclusion of oxygen, or a polyhydric alkanol of up to six carbon atoms, depositing on said support said glycerol or polyhydric alkanol, heating said support up to 200"C, cooling said support, and depositing then said catalyst thereon and thereafter heating said catalyst up to about 560"C.

1 6. The process as defined in claim 1 wherein the carbonaceous material in a second reaction zone is a residue of a first reaction zone collected in a cyclone zone held at 390"C.

1 7. The process as defined in claim 1 wherein after recovery of the product from step (c), part of said liquid product is split and reacted further in at least one additional second reaction zone (d), and a bottom product from zone (d) reaction is further reacted in still another reaction zone.

1 8. The process as defined in claim 1 wherein the reaction in step (a) and step (d) is carried out adiabatically at a temperature up to 560"C.

1 9. The process as defined in claim 1 wherein the reaction products of a crude oil residue are reacted in said zones immediately after step (a) with quenching of the top reaction products from step (a), and while a top reaction product of step (d) is further reacted, with quenching, in a further reaction zone for hydrogenation of same.

20. The process as defined in claim 1 wherein the reaction in step (a) or step (d) is with said catalyst supported on a porous metal, in an ebullating bed reactor, fixed bed reactor, liquid bed reactor, or fluidized bed reactor.

21. The process as defined in claim 6 wherein the reaction in step (a) is in an ebullating bed reactor with the catalyst supported on zeolites wherein said catalyst comprises 25% of Catalyst A and 75% of Catalyst B.

22. The process as defined in claim 6 wherein the reaction in step (d) is in an ebullating bed reactor and said catalyst is supported on a zeolite as a support and said catalyst is entirely Catalyst A.

23. The process as defined in claim 6 wherein the reaction in step (a) is in a spinning basket or ebullating bed reactor, and the catalyst is a mixture of Catalyst A and Catalyst B, C and D or mixtures of the latter and the amount of A is from 3% to 95%, by weight, said catalyst being supported on a zeolite, porous metal, chromite spinel, or alumina.

24. The process as defined in claim 6 wherein the reaction in step (d) is in a spinning basket or ebullating bed reactor and the catalyst is a mixture of Catalyst A and Catalysts B, C and D or mixtures of the latter and the amount of A is from 50 to 100%, by weight.

25. The process as defined in claim 24, wherein Catalyst A is from 50 to 95%, by weight.

26. The process as defined in claim 6, wherein the top product of steps (a) and (d) reactions are further reacted before these are hydrogenated.

27. The process as defined in claim 26, wherein the top products of steps (a) and (d) are hydrogenated with Catalyst A with KOH added thereto; Catalyst C with KHS admixed thereto; Catalysts B and A, with 50% or more of Catalyst B on K mole basis, or Catalyst B.