CH663960A5 - METHOD FOR CONVERTING CARBONATED INGREDIENTS TO PRODUCTS OF LOW VISCOSITY AND / OR STRONGER HYDRATED END PRODUCTS. - Google Patents

Description

DESCRIPTION

The present invention relates to the treatment of products such as petroleum and petroleum residues, but in particular of residues having a high boiling point in accordance with ASTM-D-1180 and Dl 160, or residue fractions having no boiling points according to the methods mentioned or destructively distilled to form carbon leaving residues Residue fractions. The invention further relates to the interdependent treatment of the above substances, petroleum and petroleum residues, in order to improve the overall yields and the space-time velocities using somewhat higher pressures, such as pressures up to 10 bar, in order to further increase the output on desired products.

In accordance with the known state of the art, the attempt to process heavy petroleum fractions into high-quality products was only accompanied by limited success. Many of the petroleum components, especially residues, essentially withstood further processing. By expelling the oily constituents from the soil products, for example expelling them by means of steam or the like, or by further processing these soil products, for example

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by hydrogen transfer reactions, partial coking, delayed coking and. Like., These soil fractions were improved and / or further components were obtained therefrom. For example, by using a solvent which is a hydrogen donor, part of the hydrogen in the solvent is transferred to the higher-boiling petroleum residue fractions, and part of the high-boiling fraction is improved in viscosity so that lighter products are obtained. In most cases, these reactions are carried out at very high temperatures and, if appropriate, in the presence of hydrogen under high pressure.

Other methods have also been used in which hydrogen split off from the very heat-resistant residues is used in a parallel stream using hydrogen as the reaction gas and the coking product is subsequently used or burned for other purposes. Excess hydrogen can also be used, for example in the donor solvent system or elsewhere in the refining cycle, to improve the residues.

In a previous work (British patent application No. 2 075 542), the patent owner specified a process for the hydrogenative treatment of petroleum residues, in which the residues are steamed and with alkali metal hydrosulfides or the empirical monosulfides or polysulfides and / or hydrates thereof and mixtures of the aforesaid are brought into contact in order to hydrogenate the hydrogen-containing material, to hydrogenate it, to treat it with hydrogenation, to remove nitrogen and / or to de-metallize and / or to desulphurise. In the process according to the above patent application, the treatment of petroleum and petroleum residues led to a considerable improvement in yield, degree of conversion per run, space-time speed, API number, etc., so that this process provided higher quality products.

U.S. Patent Nos. 4,366,044 and 4,366,045, which relate to various methods of treating coal, also mention a number of earlier references. The references cited in these two patents provide information on alkali sulfide chemistry and include references related to the treatment of crude oil or oil refinery residues. These references show a wide variety of attempts to utilize the residues or to treat various carbonaceous materials.

Another known prior art based on US patent literature was described by MJ Satrian, editor, in "Hydroprocessing Catalysts for Heavy Oil and Coal", Noyes Data Corporation, Park Ridge, New Jersey, USA (1982), for example pages 62 to 78, summarized. U.S. Patent Nos. 3,520,825, 3,775,346, 3,850,840, 4,117,099, 4,119,528 and 4,203,868 have been identified in the U.S. Patent Literature (although these may be considered in the prior art referred to above) . However, this prior art does not disclose the invention described in this application, which relates to a further discovery and to an invention over the method mentioned in published UK application 2,075,542.

Another, less relevant state of the art can also be found, for example, in Fuel, Vol. 62, Feb. 1983, where only fundamental studies of the catalytic gasification of coal and carbon are dealt with (cf. Kiku-chi et al “Supported Alkali Catalysts for Steam Gasification of Carbonaceous Residues from Petroleum, ibid, p. 226ff, which report the deposition of carbon on supported catalysts, strive for essentially complete gasification and had no knowledge of the role of alkali metal sulfides in these reactions).

Further known prior art is from 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, DC, USA- ; 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.

As part of the process dealt with here, the invention comprises a process carried out using a boiling bed reactor and supported catalysts and also a process for treating the particularly heat-resistant petroleum residues with a boiling point above 454 ° C., in particular asphaltenes or OHA fractions, that is to say a mixture of 01, Resin and asphaltenes, which remains as a residue when refining petroleum and has a boiling point of 537 C and higher determined according to the determination method according to ASTM Dl 180 or D-1160. In this regard, the present invention enables more extensive conversion of the particularly heat-resistant bottoms or residues.

Another object of the invention is to make a further contribution to the above-mentioned method. In this regard, the invention allows petroleum or crude oil, but also the less fully refined petroleum or crude oil residues, to be processed more easily and advantageously in such a way that essentially no residue remains when processing crude oil and the starting material used is essentially completely lower in products Boiling point is transferred. This is achieved in view of the fact that the difficult working conditions which lead to the usual residues when processing crude oil are not observed in the method according to the invention.

The boiling points mentioned below are boiling points at which the products (according to the defined methods) essentially completely distill up to and including this temperature and, unless stated otherwise, essentially no residue remains.

The present invention differs from the known prior art with regard to the special catalysts preferably used in the form of supported catalysts, which contain the extremely heat-resistant components of petroleum residues, e.g. B. petroleum residues with a boiling point above 537.78 C, in an extremely effective manner in products of lower viscosity.

The process according to the invention also differs from the known prior art in that the most heat-resistant constituents of the residue, for example asphaltenes, are specifically attacked during the reaction and at the same time the unfavorable influence of coke formation is avoided in view of the catalytically assisted thermal decomposition by the catalysts used here which could occur if the procedure is not carried out correctly. In addition, there are further advantages in terms of the process from the use of a boiling-bed reactor and with regard to the fact that the process is carried out at higher temperatures of, for example, up to 343.33 ° C. with simultaneous quenching of the conversion products, in order to obtain a desired end product in another reactor or in other reactors in combination with the first reactor.

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In accordance with the method according to the invention, the OHA fractions obtained from the residues, for example residues such as are obtained when refining petroleum or any other carbon source supplying these fractions, are advantageously treated according to the method according to the invention. However, the method according to the invention provides outstanding results in the processing of asphaltenes, for example asphaltenes obtained by solvent extraction, if these particularly heat-resistant products are split and / or hydrogenated in at least one stage in order to increase the overall yield which can be achieved from a given amount of oil.

This improved process is also characterized by the ability of the special catalyst to convert the Ramsbottom 15 carbon or the Conradson carbon into usable hydrogenated products without adversely affecting the process described here to some extent.

Furthermore, the method 20 according to the invention also provides for the use of particularly advantageous carrier-catalyst combinations, which can be used, for example, in a boiling-bed reactor and which produce the conversion products in a particularly advantageous manner, without disadvantageously due to those present in petroleum and in particular metal components enriched in OHA fractions.

It has also been found that after the first conversion stage carried out with the active catalysts, a second quenching stage can be provided, in which the split product is more appropriately worked up by a specially selected catalyst to the predetermined and / or the most desirable fractions can. However, the reaction in the second stage depends on the specific catalyst used in the first stage and also on the fact that the reaction is carried out properly in the reactor of the first stage.

The process according to the invention also makes it possible to recycle the product which has not been converted sufficiently in the first reactor and then to convert it into the desired end product.

Surprisingly, it was also found that the process according to the invention can be further improved by splitting the various products initially obtained from the boiling or fluidized bed reactor into at least two partial streams in the manner described later and then processing them in parallel. It was surprisingly found that the partial flow treatment (which is a continuous treatment of the initially occurring gas-vapor phase, i.e. the top products) of the bottom products collected by means of a separating device, such as a cyclone, results in the total conversion of the starting material, for example via that from the Boiling bed reactor discharged heavy fractions, can be improved 55. This is due in part to the use of different catalyst combinations, thereby avoiding inadequacies of the catalysts that occur when these catalysts are used to treat very different components of the original feed or the reflux thereof. In addition, it is advantageous for the sump products pretreated to a certain extent by this parallel treatment (for example with regard to lower API values). 65

This separate treatment in at least one parallel stream of a smaller but initially qualitatively improved fraction, which arrives, for example, from a boiling-bed reactor, offers a number of advantages. The catalyst movement is approximately 5.08 to 101.6 cm / s / 10 to 30 min residence time of the catalyst particles with an average value of approximately 43.18 cm / s / 20 min residence time of the catalyst particles (the catalyst movement is defined in accordance with the boiling bed technology) . Additional boiling bed movement can be achieved by introducing steam into the reactor and by feeding it. At the same time, it becomes easier for the first stage overhead products to achieve higher throughput rates, yields, degrees of conversion, product diversity, etc., so that the flexibility of the process with regard to the processability of a large number of starting materials is considerably improved by the parallel treatment described here, and in particular that Partial streams of the feed materials which contain fractions containing various constituents, can be worked up accordingly, but also mixed products can be worked up which are composed of starting materials which are particularly difficult to treat.

The process, in which the heavy bottom fractions can be recycled without parallel treatment, gives outstanding results, but it was possible to further improve this process. This embodiment of the process is referred to as a "return process". The improved embodiment is referred to as a "parallel process". A separate treatment without recycling is therefore more advantageous with regard to the products and product yields obtainable from a large number of heavy crude oils, fractions of petroleum refinates and in particular the heat-resistant constituents. Indeed, the parallel process described here enables almost all of the crude oil components to be converted into usable products.

The above process variants are characterized by the improved product, which is obtainable from very low-quality starting materials and starting materials with a higher hydrogen content and has a lower viscosity and smaller molecular size as a cleavage product and can be subjected to customary further processing steps.

Since a considerable degree of hydrogenation is achieved in the process according to the invention, the product is improved and, with the creation of smaller molecules, a product of lower viscosity is obtained and the formation of free Conradsen or Ramsbottom carbon is avoided and / or limited to an extremely small extent, so that the end result is an outstanding solution to the problem of ongoing research to utilize all fractions of a refinery product, which could not previously be advantageously achieved to the extent disclosed herein.

The process described here, which consists of the “recycle process” and the “parallel process”, is therefore particularly useful since petroleum fractions, which have hitherto been implemented under very harsh conditions according to the prior art, are used in a direct pass through the disclosed catalysts under less harsh conditions can be implemented either by a return process or a parallel process. The very heat-resistant soil products (the usual way of working), which form the smaller part of the entire mixture and are difficult to treat, are either pretreated and recycled or can be used instead of initially, for example, by re-cooking of the bottom product and / or recycling to be attacked to a greater extent are treated in a parallel treatment to the extent necessary. The strength of treatment in the return procedure

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or in the case of the parallel process is still significantly lower than the treatment intensity required according to the prior art and delivers far better results with regard to product application, with regard to simpler separation steps for the components and other advantageous properties explained later.

The process according to the invention is also applicable to other products such as gilsonite (also known as uintaite, a black shiny asphalt found in Utah), pitch (a natural product and / or refinery product), tars such as those from tar pits in California, asphalts (natural) like that in Asphalt lakes on Trinidad and Utah, asphalt and similar natural products or refinery processes which concentrate oils, resins and asphaltenes in the residues, products available, refinery sludge oil, which often does not have a distillation point, sump oil from coal gasification plants, e.g. products from solvent extraction from coal , or coal extracts of the type specified by Satrien, above, p. 206, or soil fractions of shale oils, heavy soil products from the Lurgi gasification process, which occur in South Africa when working according to the SASOL process, smoldering products such as those obtained when distilling these products, and both hard than soft resins and asphaltenes can also be used. It is recommended to mix some of the starting materials with each other, but this can be determined by tests according to the prevailing reaction conditions. Heavy crude oils are crude oils with an API number of less than about 22, including crude oils with a negative API number.

The combination process disclosed here provides for the more volatile petroleum fractions and similar hydrocarbons to be driven off directly, but the heavier soil products, i.e. the heavier constituents and in particular the more heat-resistant constituents, are advantageously treated according to a return process or a parallel process, depending on the required strength of the treatment or treatment intensity. For example, the starting materials that require the greatest treatment intensity are best treated using the parallel method. The process which is most suitable for the material and does not impair the overall yield of the product from the first reactor and which improves the bottom products obtained by a recycle process or in a parallel treatment sequence and increases the overall yields is thus selected.

Although the «parallel process» has been disclosed in connection with a two-stage reaction sequence or in connection with the treatment of two parallel streams, further parallel treatments can still be used where the heat-resistant products obtained in the second reaction stage or in the second stream can be converted further without recycling .

As a result of the combination of process stages disclosed above, the invention results in advantageous space-time velocities, absolute yields and yields per individual parallel stream and combined yields.

In a parallel treatment, for example, the mixture of particularly heat-resistant products of a heavy bottom fraction gave end products with a K value of about 11.71, i.e. high-molecular paraffins, which is achieved in view of the fact that the top products from the first reactor can be processed more easily, For example, to reduce the bromine number and achieve the like.

In the drawing, FIG. 1 schematically illustrates the reflux process carried out in a boiling bed reactor, and FIG. 2 schematically shows a continuous process in which the various starting materials are converted into useful products by the parallel process.

As shown in FIG. 1, the first stage reactor is a boiling bed reactor 10, which has a reaction vessel 11 equipped on the top with a collecting funnel 15 and a pump 14 for circulating the volatile or liquid products 8 participating in the reaction. The catalyst 9 is a dispersed supported catalyst which is kept in suspension by the medium 8. A medium 8 participating in the reaction and containing the catalyst 9 therefor can be circulated by pumping the medium into the reactor and distributing it within the reactor. Typically, the circulating medium can be supplied in the region of the bottom of the reactor, but it can also be supplied at any other point or at several points in the reactor. On a smaller scale, a stirred tank can be provided as a close approximation to a boiling-bed reactor, provided that the supported catalyst is in a cage or in cages or baskets such as four stainless steel enveloping sieves and these cages or baskets are suspended from a suitable frame which is suspended by a externally arranged motor is driven. By changing the speed of, for example, 20 to 180 min-1, in particular 50 to 150 min-1, the reactions could largely be adapted to the reactions taking place in a boiling-bed reactor. It is also important that there is sufficient contact between steam, catalyst and medium to ensure the desired result.

Circulated fluidized bed reactors in which the supported catalyst is circulated with the medium or fluidized bed reactors can also be used. Similarly, fixed bed reactors can be used in which the medium flows downwards or upwards for the purpose mentioned.

A preheated starting material such as asphaltenes or an OHA cut is introduced via a line 12. Steam 13 can be introduced into the reactor 10 together with the starting material, but can also be distributed to the medium from the bottom of the reactor 10.

The liquids of lower viscosity obtained are discharged via a pipeline 15a of large diameter and, if necessary, suitably cooled in a quenching zone 16 and then introduced into the second stage reactor 17. Although only one reactor is shown in the figure, several reactors connected in series or connected in parallel can also be used. These reactors can contain the catalyst as a fixed bed catalyst or can be gas phase or steam reactors. A column in which the catalyst is supported on trays 17a in the manner shown is typical. For this purpose, other known devices such as trickle bed reactors u. Like. Be used. The bottom product from the second reaction stage is collected in the collecting zone 18. This bottom product is referred to as soil product No. 1 (product from soil No. 1). All or part of this bottom product can be returned to the boiling-bed reactor together with the medium collected in the collecting funnel 15. If necessary, a pump 14a can be used for this. Part of the product can also be branched off and obtained for further processing. The top fraction from the second stage reactor 17 can be heated to reflux in a reflux cooker 19 and the products can be removed from the reactor

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Sem reflux are branched off and the gaseous products can be further processed in a second reactor, which is similar to the reactor 17, but not shown; all of these reactors, unlike the first stage reactor 10 in which the starting material is split, are referred to as second stage reactors. The products from the second stage reactor and other gaseous products can be treated in a further reactor, for example by removing any undesirable constituents such as hydrogen sulfide by passing them through a suitable bath. This can be done in a vessel 20 in which potassium hydroxide has been dissolved.

The gaseous products are then obtained in the usual way.

As shown in Fig. 2, the first stage reactor 110 is a fluidized bed reactor (boiling bed reactor). This reactor can also be a reactor with a suspended stirring basket or stirring cage or a boiling bed reactor (fluidized bed reactor). The stirring cages or baskets are not shown. However, since boiling bed reactors (fluidized bed reactors) and stirred cage or stirred basket reactors are largely similar in terms of operating behavior and mode of operation, these reactors are considered to be equivalent.

The reactor 110 consists of a reaction vessel 111 having a collecting device on the upper side, such as a funnel 115, and a pump 114 for circulating the medium 108 participating in the reaction. The feed or the medium 108 can be petroleum or a boiling below about 454.44 ° C Petroleum residue, but can also be an OHA fraction (oil, resin, asphaltenes) that boils above this temperature, for example above 565.56 C, and therefore has no boiling point. A dispersed supported catalyst 109 is circulated with the medium 108. A sieve 109a prevents the catalyst from flowing upward together with the boiling medium 108. The medium participating in the reaction and containing the catalyst can, if the catalyst remains suspended in the medium, be circulated under certain circumstances by means of the pump 114 (for circulating fluidized-bed reactors), but it is also possible for only one collected in the collecting funnel 115 and from the Catalyst-free liquid phase are circulated. As shown in the drawing, the boiling medium conveyed by pump 114 can be introduced anywhere into the reactor, although it is typically fed to or near the bottom of the reactor.

The feed 108 of the type described above is combined in a single inlet with steam and / or water 113, but individual inlets can be provided for these substances distributed around the circumference of the vessel 111. The heat-resistant petroleum components, which require a longer residence time, are subject to the catalytic reaction and are discharged via the outlet line 115a, which leads directly into a cyclone 116, in which the vaporous products from the liquid products and / or the entrained products and those from the steam products carried along with the bed. More than one cyclone 116 can be used. The bottom products 112a from the cyclone are processed in a separate parallel stream.

The top products from the cyclones 116 can be quenched in the manner illustrated in the drawing in a quenching zone 116 and then immediately implemented in a reaction vessel 117. It may be sufficient not to quench the products before introducing them into the reaction vessel 117 if the quenching can take place therein. The reason for the quenching will be explained later.

After it has been introduced into the vessel 117, a suitable catalyst described later, a supported catalyst or a catalyst without a carrier, is suspended in the manner shown schematically on a dish 117a or by other means. The catalyst effects the conversion in the gas stream or liquid stream to a lighter fraction and a heavier fraction or causes a hydrogenation of the gas stream. These bottoms can be obtained, for example, in a vessel 118 and processed separately in a further parallel flow in the usual manner or in a manner to be described later.

Further reaction vessels, not shown in the drawing, of the type of the reaction vessel 117 can be provided in the same flow path. The final withdrawal of the gases from the reflux cooler 119 or other coolers, not shown, can be carried out in the usual way by cooling and / or low-temperature cooling and / or condensation under pressure and need not be discussed.

It has proven advantageous to further treat the gaseous constituents of the initial distillate obtained immediately after the treatment in the vessel 117 at a temperature of about 130.degree. C. to about 140.degree. C. in a reaction vessel of the type designated 117 in the figure. If more saturated products are desired, a suitably selected and adjusted reflux column 119 can be used to collect these products.

Bottom products 112a of the reaction product separated, for example, in cyclones 116 are generally the more heat-resistant constituents of petroleum or petroleum products. These products 112a, which can be assumed to have been partially implemented in reactor 110 (which results, for example, from the improved API number), are best processed in a first parallel branch. Since these bottoms have been converted to some extent because of the catalyst 109 suspended in the boiling bed, the further implementation is best carried out with another specific catalyst that has been modified, if necessary, to suit the particular feed.

This second stream of feed material thus consists of the bottom products 112a, which are fed to the reaction vessel 130 together with steam, and this stream, which is likewise fed in, is recommended for a catalytic treatment of the still hot product. This feed 112a is more amenable to hydrogenation, which provides further improvements that allow lower viscosity components of the heavier petroleum fraction to be obtained. The bottom products 112a fed can advantageously be treated with a catalyst of different composition in the presence of steam in the same ratio range for steam as for the feed fed to the first reactor 111. Catalysts of various compositions have proven to be more advantageous in the later reactions than the catalyst originally used in the boiling bed and designated 109 there. The use of a different catalyst is due to the partial conversion of the more heat-resistant components in the vessel 111. The catalyst compositions of catalysts for treating the various soil fractions will be discussed further later.

Since another reaction vessel (not shown) can follow the reaction vessel 130, the overhead products obtained in the reflux column 133 can be processed further from the reaction vessel 130. Usually

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The products emerging from the top of the reflux column 133 can be freed of undesired constituents, for example H2S, by washing in a vessel 135 and the gaseous products can then be recovered and reused. The reaction vessel 130 can be used in a similar manner to the reaction vessel 117, that is to say it can be used as an upward or downward flow reactor with a supported catalyst or an unsupported catalyst suspended in a bed or on trays 130a, or as a fixed bed catalyst. A liquid catalyst such as water-containing melts of various alkali metal sulfides that are liquid at the desired temperature can also be used in the downstream reactors in the manner described in U.S. Patent 4,366,044.

If no further reduction in the molecular size of the overhead distillate from the first reactor is desired, this overhead distillate should be in the vapor phase of the reactors immediately after leaving the cyclone 116, which is the bottom products (the heavy portion, the partially converted portion or the unreacted portion to be recycled ) separated from the top distillate representing the product, treated with hydrogenation. Quenching, for example in a quenching device 116a, prevents the split components from condensing again. The hydrating treatment is facilitated by the absence of the heavier fractions separated in cyclone 116. By definition, the hydrogenating treatment takes place without thermal or catalytic cracking by introducing hydrogen into a molecule. Since the hydrogenation reaction is exothermic, the aforementioned quenching device 116a helps to control the temperature and to avoid impairment of the products by excessively high temperatures.

Fixed bed reactors or liquid catalyst reactors can be used for this hydrogenating treatment. At the moment, however, reactors for liquid catalysts at temperatures below 280 C seem to work with the greatest efficiency, since the catalysts solidify above 280 C due to the mutual conversion of the hydrates.

With regard to the degree of hydrogenation, fixed bed reactors are inferior to the reactors operated with liquid catalysts.

The feedstock, typically an OHA fraction consisting of oil, resin and asphaltenes, is a material that boils above 537.78 C at atmospheric pressure. While this is a rough description because the amount of oil, resin and asphaltenes is not necessarily known, it does provide a useful measure of the most heat-resistant component in an oil. It is known that asphaltenes can be extracted from petroleum residues or from the OHA fraction by solvents, and a residue of oil and resin to be treated is obtained. Asphaltenes are, of course, particularly inaccessible to further processing, for example a catalytic or other further processing, and thus represent a fraction which, according to known methods, could only be used by burning or by known coking and splitting off all the hydrogen present.

A useful characterization of asphaltenes is that it is the portion of an asphalt or bitumen that is soluble in carbon disulphide but insoluble in paraffins, for example heptane or paraffin oil, or in ether. Resins can be extracted from OHA fractions using propane. Bitumen is also soluble in carbon disulphide. Carbenes, which form part of a bitumen, are insoluble in carbon tetrachloride, but soluble in carbon disulfide. The oily or soft components of bitumen are also known as malthene or maltene and are soluble in white spirit. Maltenes are compounds soluble in pentane and asphaltenes are compounds insoluble in pentane.

In the broader sense, the above definition includes natural asphalts such as petroleum, mineral pitch, earth pitch, Trinidad pitch and petroleum pitch, and natural hydrocarbon mixtures such as amorphous solid or semi-solid fractions, such as those obtained in the oxidation of residual oils.

Since there is no agreement on a precise definition of these compounds, such as maltenes or asphaltenes or mixtures thereof, mixtures of these compounds are often referred to once and differently in accordance with the known prior art. In addition, the solvents practically used in the extraction and in the precipitation influence the properties of the end product to a greater or lesser extent. For this reason, the asphaltenes extracted with a solvent such as carbon disulfide and precipitated from heptane are not yet regarded as pure compounds because they have no clear melting point, but only a softening point. The softening points of asphaltenes can be up to 204.44 ~ C and more. For this reason, it is useful to define the OHA fraction as a fraction with a boiling point of 537.78 ° C or more, although the temperature limit is chosen arbitrarily and also low temperatures of, for example, 482.22 C can be selected, since not the whole Material can be driven off in the desired manner. A lower temperature of 426.67 C only characterizes a less inaccessible composition.

For example, an asphaltene with a high softening point obtained by solvent extraction has a softening temperature of approximately 132.22 ° C., a density of approximately 1.1149 at 15.56 ° C. and a viscosity of 4060 P at 135 ° C. The density at 135 ° C. is 1.026 , with which the viscosity is 3957 St or 395 700 cSt (the viscosity in St is obtained by dividing the viscosity given in P by the density at the specified temperature) The viscosity of the same asphaltene, which has a high softening point, is 148.89 C. 877 P and 86,000 cSt considering the density of 1,016 measured at 148.89 "C. The viscosity of this asphaltene at 162.78 ° C is 261.5 P or 26 000 cSt considering the specific weight of 1.006 at 162.78 'C.

The analysis for the above asphaltene obtained by solvent extraction can be found in the following table.

Analysis of asphalting

Metals

carbon

84.59%

Fe

360 ppm

hydrogen

8.80%

Ni

147 ppm

nitrogen

0.82%

V

490 ppm

sulfur

5.52%

N / A

497 ppm

ash

0.27%

K

4 ppm

humidity

0.0%

oxygen

-

total

100.00%

Calorific value of the asphaltenes: 5.1629 kWh (17 627 B.T.U.)

Although the above-mentioned asphaltenes can be considered representative, various other asphaltenes can have different characteristics depending on their origin.

From the analysis above it can be seen that these product

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contain considerable amounts of metals. The amount of metals depends on the origin of the material and can be up to 6000 ppm vanadium and is usually around 600 ppm. Nickel and other metallic components can be present in amounts up to the latter amount. These metallic components therefore also impair the processability of the residue by customary processes for processing petroleum residues.

On the basis of various analyzes, the hydrogen content of an OHA fraction can be in the range from 13.5 to about 7% by weight or less, but this is not an exact identification here either. The hydrogen content of an oil residue boiling above 537.78 ° C is about 12.5% and less. There is also “free” carbon (as Conradsen carbon) in abundant quantities, for example up to 45% by weight. Free carbon is defined as Conradsen carbon or Ramsbottom carbon, but these analytical values are not identical because different methods are used to define “free” carbon and this carbon is actually not “free”. In the OHA fraction, the Conradsen carbon content can be up to 40% by weight or more. In any case, the carbon residue can be made accessible for conversion by the process according to the invention.

Every cracking, be it thermal cracking or catalytic cracking, produces smaller molecules than they were contained in the starting mixture to be cracked. A smaller molecule has a boiling point that is at a lower temperature than that of a larger molecule. A hydrating treatment does not seem to result in cracking and the formation of smaller molecules. The temperature range of the boiling points appears to be the same before and after the hydrating treatment.

The hydrogenating treatment effects, for example, the hydrogenation of olefins to paraffins and the removal of diolefins and triolefins in the product by hydrogenating these diolefins and triolefins to monoolefins or to paraffins. Aromatics are converted into naphthenes by hydrogenation treatment. Real paraffins do not appear to be involved in the implementation of the hydrating treatment.

The sulfur content and nitrogen content of the product can also be reduced by a hydrating treatment. However, if the asphaltenes containing most of the nitrogen have not previously been cracked, the nitrogen of these asphaltenes is not amenable to being split off by the hydrogenation treatment. The extent of the sulfur and nitrogen cleavage is equivalent to the hydrogenation treatment of an unsaturated bond (where there is no hydrogen at the point where hydrogen is to be introduced into the molecule). The sulfur and nitrogen compounds which can be converted into more saturated compounds by hydrogenating treatment with the elimination of ammonia and / or hydrogen sulfide include those which have exposed chemical and chemical atoms in their exposed position and which can be replaced by hydrogen in a thermodynamically favorable exchange reaction.

The bromine number is a measure of the degree of unsaturation of the product. The lighter fractions of unsaturated products have the highest bromine numbers. By hydrogenating the product, the bromine number is reduced to acceptable levels, for example 40 to 50, for a product with an initial boiling point of 204.44 ° C (which product was obtained by distillation within the product range of the methods disclosed above).

As already mentioned above, the bottom products 31 can form a further separate parallel branch and can be used as feed material, but these products can likewise be recycled since the further implementation only causes further complications. If, however, the soil products 31 are difficult to split, they can be subjected to a further operation of the type shown for the loading 12a and in particular in FIG. 2.

Since the majority of the top products from the reaction vessel 11 are treated in a single straight pass, the bottom fractions 12a collected, for example, in the cyclone 16 and the 15 fractions collected in the 15 collecting vessel 18, which, however, make up a smaller proportion depending on the composition of the petroleum. These soil products can thus be treated in smaller amounts, taking into account the proportions of the constituents 20 in the petroleum product to be worked up. Accordingly, a wide variety of mixtures of a wide variety of petroleum can be beneficially treated depending on the constituents of the soil product 12a. The size of the reactor 30 can be dimensioned in accordance with the proportion of petroleum in soil product components, for example those with a boiling temperature of 537.78 ° C. and above.

Steam is fed into reactor 110 or reactor 130 at a rate such that the desired product dictates to some extent the amount of steam used. The steam can be used in an amount of about 50 to 100 and even 130% by weight of the product obtained. The amount of steam can also be expressed on a lower limit basis 35 such that, due to the lack of steam, coking does not occur to any appreciable extent, but would occur in reactors 110 and 130 if there was insufficient steam . The lower limit of the amount of steam supplied must therefore be 40 such that any significant coking is avoided in the reactors 110, 117, 130 etc. Additional steam can be supplied, for example for the reactors 117, 130, etc. Low-tension steam, typically exhaust steam, is used as steam. 45 The temperature in the reactor can generally be up to 650 ° C, although the most advantageous working temperature is below 450 ° C, for example below 425 ° C. In the following reaction vessels, for example in reaction vessels 117, etc., however, lower temperatures can be maintained at which a more hydrogenated product, such as a product having a lower bromine number, appears to be produced. However, treating the more heat constituent in reactor 130 may require the same temperature as in reactor 110. 55 Quenching can be done at a temperature of 390 ° C or below. The third stage reactor, not shown, can have a temperature in the range of 170 ° C and below, etc. down to 125 ° C (without steam condensing). Lower temperatures are desirable because the hydrogenation is exothermic and most of the hydrogenation is done in the downstream reactors, while most of the cracking is done in reactor 110 and less cracking is done in reactor 130.

65 The pressure in vessel 111 and downstream thereof can be kept at atmospheric pressure. Although pressures below atmospheric pressure are possible, they are not so useful. If the pressure is between about 6.67 and about

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10.01 bar (about 100 psig or less than 150 psig), the reaction seems to proceed better, although pressures up to 25 bar can also be used. When working at higher pressures, a pressure between 4 and 10.01 bar seems to be the most suitable for unexplained reasons. When working under higher pressures, steam condensation must be avoided. Although the exact reason is unknown and the reaction mechanism is still unclear, the hydrogenation in this area appears to be proceeding better, but increasing the pressure affects the economy due to the more complicated setup and the higher capital investment.

Furthermore, instead of separating the liquid and / or entrained products in cyclones such as the cyclones 116 shown in the drawing, of which at least one is provided, another separation device can also be provided in order to treat the petroleum constituents in parallel ways; such other devices can centrifugal separators, separating drums or the like. be.

Various crude oils of any origin and viscosity can be subjected to the treatment. Provided they are sufficiently liquid at the treatment temperature, these crude oils can be treated and provide the outstanding results stated herein. Residues and soil products from atmospheric pressure refining can also be treated. These starting materials are thus identified as products having a boiling point less than about 454.44 ° C, however their constituents or their reaction products coming from reactor 110 which have a boiling point above 454.44 ° C are best after one Processed in parallel.

The further processing in the second parallel reaction path provides products with unusually improved viscosity, improved hydrogen content and a better chemical structure, for example products with a higher content of paraffins with a high K value or aromatic polycyclic compounds, which subsequently also produce desired, olefinically less unsaturated cyclic compounds and linear or branched paraffins can be processed. The gaseous products obtained in the course of the process according to the invention and the conversion of the unsaturated products into more saturated products make a further contribution to the process described above by the patentee.

In order to reduce the bromine number and to reduce the nitrogen and sulfur content of the hydrogenating treatment of the products obtained by this process, it must be taken into account that the hydrogenating treatment at the lower temperature does not "crack" the product to any appreciable extent with the formation of small molecules. which would reduce the boiling temperature range. Hydrogenating treatment of the product would further improve quality, e.g. by cracking (determined by the lower boiling temperature range of the products of smaller average molecular size), since paraffins obviously cannot be cracked in this process.

The hydrogenating treatment should therefore form the final stage of the process described here and should only be provided when the desired molecular size has been reached in the course of the “cracking stages” of this process. Thus, feed materials formed from residues from vacuum distillation are initially cracked to a smaller molecular size under less harsh conditions and then the heat-resistant residues, etc. the

Feed stream 112a, cracked under the required appropriately harsh conditions that do not adversely affect the products of the main stream. The product obtained has a lower boiling temperature range. This reduction in molecular size is accompanied by a reduction in the sulfur content and the nitrogen content of the top product, based on the feedstock.

The smaller molecules forming the “cracking products” do not contain as large a quantity of nitrogen as the starting material. Nitrogen is usually contained in the sheared "asphalt tail" of the oil residues. Metals bound via the nitrogen to the methylene groups of the porphyrin skeleton of the asphaltenes are accordingly separated in the parallel reaction path. The metals and the nitrogen thus remain in the unreacted central porphyrin skeleton of the asphaltenes and are therefore not contained in the overhead product of the distillation, or are contained only in very small amounts, which was obtained from the vacuum residues containing the asphaltenes, the asphaltenes relative to the asphaltene content of the Crude oil from which these vacuum residues originated have been enriched. The metals and nitrogen are present in increased concentration in bottom product no. 1, for example 112a, which does not provide overhead distillate. The sulfur content in the asphalt content of the crude oil or in the residue is not necessarily increased. The resin content of the residue can contain up to 5% by weight of sulfur.

The catalysts, which are sometimes also referred to as reagents, are used in a form which brings about the desired results, namely adequate cleavage and / or sufficient cleavage of the bonds within the first reactor and the further parallel-flow reactors, or in the form finally desired. These results are obtained by first appropriately subjecting the product treated first to a strong but not overly strong treatment with the catalysts described herein, which are described in connection with the results obtained, and then to pretreating the (in reactor 110) heat-resistant components in the parallel branch are treated accordingly. Mixtures of the catalysts are considered to be extremely desirable for crude oils and their residues or for mixtures of different fractions of crude oils.

A suitable, highly effective catalyst for reactor 110 is referred to as catalyst A and is prepared in the following manner. 1 mol of potassium hydroxide is dissolved either in ethanol, methanol, preferably in ethanol, or in a mixture of ethanol and methanol or, which is less advantageous because of the lower solubility, in 1-propanol or 1-butanol. The solubility of the catalyst is lower in the latter two, which is why larger amounts of alkanol must be used, which then has to be separated off again. The alkanols can be absolute alkanols, but these can be of a similar nature to 95% ethanol. The potassium hydroxide dissolved in this solution is then reacted with hydrogen sulfide passed through this solution. After complete saturation, the catalyst-alcohol mixture is separated off and the alcohol is separated off by suction in a vacuum. The residue is the catalyst. In the case of a solvent mixture, typically 1 mole of potassium hydroxide can be dissolved in 200 ml of ethanol and 130 ml of methanol. Potassium hydroxide is typically used analytically pure in the form of cookies (about 86% KOH), absolute ethanol or 95% ethanol and absolute methanol. As mentioned above, the proportions of ethanol and methanol in the mixture can be varied5

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the. The solution is evaporated in vacuo until no residual alkanol can be driven off. These catalysts are very suitable for splitting and in particular for splitting the heat-resistant constituents and are used in suitable amounts together with other catalysts or even alone in order to bring about the desired splitting.

The catalyst can be used in the first reaction stage without a carrier, but the catalyst for the first reaction stage or for subsequent reaction stages is preferably applied to the carrier and calcined.

The catalyst supports used were mainly used to enlarge the catalyst surface. The catalyst supports are spinels such as chromium spinel (CrO) and most advantageously porous metals such as stainless steel of the existing AISI grades and the like. The latter are obtained by sintering powder-metallurgical particles of uniform and very small particle size or by leaching out leachable components from thin, for example 3 18 mm thick sheet metal, which creates interconnected channels. Other metallic supports are, for example, those made by sintering very thin, approximately 0.2 to 5 mm thick wires and cutting them to a length of, for example, 2 to 5 mm. Other carriers are, for example, alumina with pore sizes of 5 to 35 nm and even up to 100 nm, but it may be necessary to protect such carriers in a manner to be described later. Although the treatment in downstream reactors may place lower demands on the properties of a support, the treatment in the reactor of the first reaction stage of the process according to the invention is best carried out with a strong, inert support, such as a support made of a porous metal, the pore size of which is up to about 350 nm and more can reach, so that the metallic carrier can have 10% metal and 90% pores based on its volume, although the metal can also claim about 25% by volume.

In all reactions, the catalyst is reacted in the absence of atmospheric oxygen and thus in the absence of oxygen. Likewise, the catalyst is deposited on the support in the absence of oxygen and volatile is stripped from the support in the absence of oxygen.

After the volatile constituents have been driven off, the catalyst and the carrier are heated to a suitable temperature, for example a temperature of 320 to 450 ° C. or even 560 ° C. The catalyst adheres firmly to the support and can therefore also be used in a spinning cage reactor (also called a spinning basket reactor), a boiling-bed reactor or a fluidized-bed reactor.

If the support is severely attacked by the catalyst, as is the case, for example, in the first stage reactor with alumina as the support, the following procedure is used. The above catalyst is largely evaporated to dryness and then dissolved in glycerin, whereupon the glycerin-catalyst mixture is deposited on a carrier, for example alumina. Other carriers less resistant to attack by the catalyst, such as carriers formed by molecular sieves, are treated in a similar manner. These molecular sieves can typically be of the Y and X type, for example YL-82, low sodium (available from Union Carbide, Danbury, CT, or comparable carriers from Mobil Oil, New York, NY). However, the molecular sieves act as carriers for the catalyst, thus increasing the contact area of the catalyst.

After being deposited on the support, the glycerin-catalyst mixture is slowly heated to a temperature of, for example, 560 ° C., which expels the volatile.

It is also possible to apply the glycerol to the support first and then to heat it to about 200 ° C., 5 then to apply the catalyst to the cooled support and finally to heat it to the desired temperature.

The reaction in the first reactor can be carried out at a higher temperature, for example a temperature in the range from about 320 to about 450 ° C., even though temperatures up to 560 ° C. and even up to 650 ° C. were used. For asphaltenes, the preferred temperature range is from about 360 to about 430 ° C, and it appears that a very good working temperature range is between 390 and 425 ° C.

Since in connection with these catalysts the reaction must be carried out in the presence of steam to facilitate hydrogenation, hydrogenating cracking, etc., steam in an amount of at least about 20% by weight of the feed, for example the OHA fraction , a crude oil, the residues, etc. of the reactor. On the other hand, it is also possible to increase or decrease the amount of water supplied at the working temperature in the form of steam, depending on the degree of hydrogenation (to a certain extent hydrogenation can also take place in the first reactor). If more hydrogenation is desired, more steam is added, but the amount of steam typically does not exceed about 85% by weight of the feed, whether or not 130% by weight of steam is added based on the hydrogenation product obtained If a gaseous fraction is formed, it is converted into a liquid equivalent. In other words, the amount of water to be used is determined by subtracting the hydrogen content of the feed from the hydrogen content of the desired product and the difference obtained by the factor Is multiplied by 9 (since the molecular weight of hydrogen is 1/9 of the molecular weight of water). Typically, 40 is introduced into the first reactor in an excess of 30% water.

If water, for example in the form of steam, is not introduced into the reactor or the water supply is interrupted for any reason, coking can occur, with carbon being deposited or being produced in a process similar to the catalytic thermal cracking process , however, in this case the catalyst acts as a thermal cracking catalyst and this with some advantage (since the catalytic thermal cracking takes place at 50 a fairly low temperature of, for example, 320 ° C.) but with a considerably lower efficiency than it does in the presence of steam Hydrogenation catalyst and / or cracking catalyst shows. The aforementioned minor coking can thus be taken as a lower, though less desirable, lower limit.

Insufficient steam supply or intermittent steam supply also causes the formation of particularly heavy products in reactor 110. It is important that steam is always appropriately introduced into the reactor and thoroughly distributed (without leaving any steam-free and / or unreacted space).

It must be mentioned, however, that excessive steam supply obviously interferes with a faultless reaction process, because partially converted products are entrained.

If carbon separates for any reason, it usually happens in hot places like

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the heated reactor walls or the catalyst support. The reactor 10 is therefore preferably operated adiabatically. Carbon deposits on the catalyst can be removed and converted into usable products by exposing the catalyst to the action of the steam for a certain period of time without the addition of further starting material, so that the catalyst is then usable again and for the production of the desired product or desired fraction can be used. Batch feeding of hydrogen sulfide or sulfur can be useful in general and especially for low sulfur feedstocks such as sweet crude oils.

It has also been found that when the working temperature is increased to about 320 to 420 ° C., for example when using the above catalyst, depending on the composition of the feed, for example on account of its asphaltene content, an exothermic reaction takes place, for example at 440 CC can use. The exothermic reaction can cause the temperature to rise to 600 ° C., but this temperature rise depends on the amount of steam supplied. A larger amount of steam results in lighter carbon compounds. Excessive temperatures are not desirable, which is why temperatures below 440 ° C. are preferred.

In the reactor 17 shown in FIG. 1 or in the second reactor 117 shown in FIG. 2, where the further reactions take place, the products coming from the first reactor rapidly increase to about 250 ° C. in the presence of a catalyst, usually to temperatures between 250 and 390 C cooled. The quality of the product is increased to 430 ° C., preferably 425 ° C. when the working temperatures are increased, but there is no increase in the degree of conversion above 390 ° C. The cooling in the quencher 16a takes place to such an extent that steam does not condense and the reaction is disrupted. Of course, the light tail, which are hydrogenated products, is not condensed. The catalyst in the second reactor 17 or 117 is preferably also a supported catalyst, but the catalyst can also be used without a carrier.

A typical catalyst for the reaction in the second stage, that is to say in reactor 17 or 117, is a less potent catalyst B. This catalyst is produced in that technically pure or analytically pure potassium hydroxide, which contains about 86% potassium hydroxide, in absolute or 95% iges ethanol or methanol (preferably ethanol) and the resulting solution is saturated with hydrogen sulfide without evaporation of the alkanol, entrained alcohol is collected in a reflux vessel. Other vessels suitable for carrying out the reaction can be operated with a downward flow in order to be able to capture the hydrogen sulfide. As soon as the last vessel containing KOH indicates the start of a reaction, the reaction in all upstream vessels is stopped.

If the reaction is carried out in a further reactor or in further reactors 17 or 117, that is to say in second-stage reactors, the advantages which can be achieved by the process are that the reaction products from the first-stage reactor are immediately at about 300 ° C., however preferably 250 C, are quenched in the presence of the catalyst and then the products of the first stage are reacted in the second stage. For this purpose, it has proven to be particularly advantageous to support the catalyst on a suitable support. The supports can be the same as in the first stage, but in any case must be inert under the reaction conditions prevailing in the special reactor (for example 17 in FIG. 1 or 117 and 130 in FIG. 2). The second stage reactors 17 and 117 can be operated as fluidized bed reactors (circulating fluidized bed, partially circulated fluidized bed or closed fluidized bed), as fixed bed reactors or as fluidized bed reactors. The reactor 30 can also be a boiling bed reactor or a fluidized bed reactor or a circulating fluidized bed reactor.

It has been found to be acceptable for the second stage reactors, for example for reactors 17 or 117, to use generally available supports such as alumina, fixed zeolite type aluminum silicates, i.e. molecular sieve type, in which the sodium or potassium is replaced by ammonium. Zeolites (10 and 13) of type X and Y are suitable. Y-type zeolite molecular sieves are preferred, and of these, molecular sieves with a low sodium content of less than about 1.0% Na20 are particularly desirable. The molar ratio of silica to alumina in these zeolites is greater than approximately 3: 1, for example approximately 5: 1, etc., Na20 being present in an amount of approximately 0.2% by weight. These zeolites are commercially available in powder form or in spherical form or in the form of cylindrical or differently shaped extrusions, etc., the suitable dimension being, for example, 3 mm for extrusions or spheres. Although it has been assumed that these supports are poisoned or destroyed by alkali metals in the procedure described below, these supports are useful in spite of the reagents deposited thereon from the described alkali metal sulfides. These supports can also be used in the first stage reactor 10.

Other zeolites are the potassium type ELZ-L zeolite of the type described in U.S. Patent 3,216,789 and the silicalite material described in U.S. Patent 4,061,724. The latter material has a pore diameter of about 0.6 nm. Other carriers are, for example, those described in GB-PS 1 178 186, that is to say carriers with a very low sodium content of less than 0.7% by weight, for example ELZ-fì- 6, or ELZ-E-6, E-8 or E-10. Other carriers are very low sodium, mordenites and erionites available through exchange for ammonia. Of the above-mentioned molecular sieves, the low-sodium molecular sieves which are subjected to ion exchange with ammonia, for example 0, are available under the trade name LZ-Y82 from Linde Division, Union Carbide Corporation, Danbury, CT, Mobil Oil Corporation, New York, NY and other sources of supply. 15% by weight, the Y-type molecular sieves are preferred. In any case, the stability and service life of these molecular sieves used as supports are determined under the reaction conditions, the behavior of the supports in the second stage reactor being determined.

The supports for the second stage are made in the following manner. The zeolite extrusions, which have a low sodium content and have been subjected to an ion exchange for ammonium ions and can be present, for example, as a powder or as shaped bodies such as cylinders, saddles, stars, rings, balls, etc., the powders or extrusions having dimensions of approximately 3.18 to 3 , 97 or 4.76 mm, are treated with glycerol or similar polyhydroxyalkanes, such as partially reacted polyhydroxy compounds including hexahydroxyalkanes, by first impregnating these moldings in a sealed reactor. Then, if, for example when using glycerol, the powders or moldings are heated from room temperature to 265 to 280 ° C. and even up to 560 ° C. and the resulting decomposition products are driven off, one runs

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favorable, if unknown reaction. The support thus reacted is then sieved, drawn off and cooled in a tightly closed container when the temperature has been brought to 560 ° C.

The carrier described above is impregnated (empirically) in the cold state with a catalyst of the general formula K2S1.5; this catalyst is acceptable, but not outstanding for splitting. Although the catalyst is described by an empirical formula, the behavior of the catalyst is determined by the manufacturing method chosen. Accordingly, a wide variety of catalysts can be illustrated by the same empirical formula, however, very different properties are observed on these catalysts and different products are obtained when different starting materials are treated. This catalyst, referred to as catalyst C, is dissolved by dissolving 6 moles of KOH in 4.5 to 7.5 moles of H20 without external heating, whereupon a small amount of an alkanol, for example 2 to 2.5 ml of methanol or ethanol, is used per mole of KOH. is added. 4 moles of elemental sulfur are then added to the solution thus obtained, which reacts exothermically. Elemental sulfur is then added to obtain a sulfide of the desired empirical formula, which can range from K2S to K2S5 but can also range from K2S] j to K2S25, depending on which end product is desired. For a higher gas content in the product, catalysts are used which are less saturated with sulfur. For a higher liquid content in the product, more saturated catalysts are used with sulfur.

Another catalyst, referred to as catalyst D, is produced as follows. 1 mol of KOH is dissolved in 1.0 mol of water with vigorous stirring. Immediately after the KOH has dissolved, 2 ml of methanol or ethanol are added. Immediately afterwards, 0.66 mol of elemental sulfur is added, whereupon the sulfur is allowed to react in the course of a more violent reaction. The catalyst is brought to the desired empirical sulfur content with vigorous stirring by adding a suitable amount of sulfur, for example 1/4 of 2/3 mol of sulfur adding 0.5 of the empirical sulfur content of K2S, ie 1/4 of 2/3 mol dissolved sulfur K2Si, 5, 1/2 of 2/3 mol of the dissolved sulfur K2S2, o etc., including other suitable fractions. The composition of the catalyst can range from K2Su to K2S2.5 or even to K2S5. The lower sulfur catalysts, when mixed with Catalyst A, are good cleavage catalysts and are generally good catalysts for a less harsh treatment of crude oils. This catalyst is also a good hydrogenation catalyst. The catalyst A is generally added to the catalyst D in an amount of 3 to 25% by weight, usually in an amount of less than 10% by weight.

Once the catalyst has been prepared in this manner, it is evaporated in vacuo to a flowable slurry which is then poured over the glycerin treated and cooled as described above (if the support has been heated to 300 ° C or higher) is), stirring in a high vacuum and finally suctioned to dryness. The catalyst is further sieved in the dry state and then immediately introduced into the second stage reactor which, if the catalyst is used as the hydrogenation catalyst, has been flushed out to displace atmospheric oxygen.

Another method of protecting a glycerol-treated and to a temperature between 260 ° C and a decomposition point (which is characterized by considerably slowing the formation of a trapped liquid

To recognize condensate) treated carrier consists in adding the catalyst slurry described above, closing the vessel and heating to a temperature of at least 440 ° C up to and including 560 ° C. In any case, the catalyst is calcined above its use temperature in the process.

Another method is to add the glycerin, for example about 88 ml of glycerin, to about 1 mole (potassium based) of the dissolved catalyst or to add the above catalysts or mixtures thereof to the glycerin. The catalyst-glycerin mixture is then heated to drive off water and / or alkanol, leaving a solution of the catalyst. The temperature is raised up to 190 ° C to achieve this. The mixture is then poured over the carrier and the whole is heated to a temperature of at least 450 ° C to 560 ° C. This catalyst has a very unpleasant smell and must be manufactured with a good finish

When used in a first stage reactor with a capacity of one gallon (3.7851) together with a second stage reactor, the second reactor is charged with about 2/3 mole of supported catalyst (potassium based), e.g. the second reactor can be fed into the reactor Step an alumina catalyst of empirical composition K2S1j5 are introduced.

Another catalyst referred to as catalyst E is an unsupported catalyst or a supported catalyst that is capable of reducing the molecular size of the product in a first stage reactor (or that can be used in a further second reaction stage). Catalyst E is obtained by powdering dry KHS or a slurry thereof in appropriate amounts (mole% or% by weight) of any of the above reagent mixtures (catalysts) A, B, based on the desired size of the product , C or D is added. The catalyst can be used without a carrier or with a carrier. Thus 1/5 to 1/3 mol, based on mol K of KHS, is added to, for example, K2S (empirically) or to K2Sj 5 (empirically), with the addition of KHS reducing the molecular size of the product.

If the process is carried out in the second stage reactor with this supported catalyst, suitable settings can be made, since, for example, K2Si, i or K2S1> 5 is more hydrogenated and K2S2 provides larger molecules and thus also more distillate and fewer gases. These reactions are carried out in a temperature range of 113 to 440 ° C. Similar catalyst settings can be made in other reactors, for example when more than one second stage reactor 17 or 117 is used.

In any case, the reaction in the first stage is carried out with the reagent indicated in order to cleave the heat-resistant, poorly degradable constituents of the starting material, for example crude oil residues, the OHA fraction and in particular asphaltenes, to the desired extent. In the case of the parallel treatment, the as yet unreacted compounds which are difficult to decompose are catalytically treated in order to achieve the desired degree of cleavage. The overall process combination in the further stages, that is to say the second, third, fourth, etc. stage, depends on the stated implementation in the first stage and the separated portion of unreacted constituents and is therefore dependent on it.

The amount of catalyst deposited on the support is from about 4 moles of catalyst (based on potassium) to about 0.5 mole or even only 0.1 mole (based on potassium)

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500 ml carrier. On a different but not identical basis, the amount of catalyst is about 20 g per 300 ml of support, but the amount of catalyst can range from 3 g or 5 g of reagent per 100 ml of support to about 25 to 30 g per 100 ml of support.

The vanadium can be removed entirely from the crude oil to be worked up, the heaviest products containing essentially all of the vanadium. When working with the supported catalysts described here, vanadium is present analytically in bottom product no. 2, that is to say in product 18, only in a very small amount or in undetectable amounts. The removal of nickel is promoted when part of the reagent is in the form of hydrosulfide. A reaction between the alkali metal sulfides and nickel sulfides does not take place, but a solubility reaction is observed when alkali metal hydrosulfide and nickel sulfides are present. Nickel (and iron) form complexes similar to ferrites with the alkali metal sulfides and hydrosulfides. These complexes are immediately hydrolyzed in liquid water to insoluble iron hydroxide or nickel hydroxide, which precipitates out. The complex of vanadium and catalyst is extremely soluble and stable in liquid water. Iron is contained in the residue, according to the distillation range determinations, in amounts between 3 and 5 ppm, but this amount also depends on the amounts in the starting product.

In general, the catalysts for the reactions in the second and further stages are the hydrosulfides and the sulfides, that is to say monosulfides and polysulfides of the elements of group IA of the periodic system of the elements which are different from hydrogen and, as mentioned above, are prepared from alkanol solutions were. While sodium, potassium, rubidium and lithium can be used for the stated purpose, sodium and potassium are by far the most advantageous, of which both are particularly preferred. Although rubidium compounds are useful, they are not inexpensive like lithium compounds. In the first stage reactor, however, rubidium can be present in a mixture of 40% rubidium sulfides, 26% potassium sulfides and 60% sodium sulfides, that is to say in the form of the various types of sulfides and calculated as metallic elements in% by weight. The range of ratios for the aforementioned mixtures is 1: 1.5 to 2.5: 3.5 to 4.5, but these mixtures must be prepared in the manner described for Catalyst A. The typically used catalysts are used in the form of the hydrates, however, a small part of the catalyst, etc. about up to 15% by weight and typically less than 10% by weight or even less than 5% by weight, obviously in the form of an alkanolate (the analogue of the hydrate). Alkali metal ethonates are also available as transition intermediates. Hydrates (and alkanolates) of these compounds are very complex and initially undergo a number of conversions depending on the prevailing reaction conditions. It appears that the presence of mixed hydrates and alkanolates is required for the superior results. No attempt has been made to elucidate the nature of these conversions or the actual structure under the reaction conditions for the sulfides, hydrates, alkanolates, thionates or their mixtures. It is sufficient to point out, however, that a catalyst used can be a mixture of a number of hydrates and / or alkanoates or a eutectic mixture of different hydrates and / or alkanolates, in which case, however, the specific cracking catalyst is used for the heat-resistant compounds got to.

No attempt was made at the

To identify reaction to mutual conversion of the various sulfur-containing sulfides. If, however, as mentioned before, the loading of the first stage heat-resistant components, etc. Contains asphaltenes of various molecular sizes, resins and oils such as mud oils, in significant amounts of more than 50% by weight and, including a fraction with a boiling point of more than 537.78 ° C, of up to 95% by weight Implementation in the first stage requires at least part of the specific catalyst defined above, for example catalyst A. The lower the proportion of the feed in more heat-resistant compounds, the less catalyst A is required. Catalyst A is also recommended for feed I2a. Depending on the desired molecular breakdown (which is characterized by the API number, the viscosity, etc.), this catalyst can also be combined with catalysts of a different composition, since this is the best way of effecting the molecular breakdown without generating excessive amounts of gas. For the bottom product No. 1, ie the feed 12a, a boiling-bed reactor or a reactor with a stirred cage or a stirred basket is used as the reaction vessel, but a fluidized-bed reactor and a fixed-bed reactor can also preferably be used. Another catalyst for the feed 12a is catalyst B, to which KHS has been added in an amount of 5% by weight (based on elemental potassium) or which has been replaced by up to 95% of catalyst A, taking into account the same merge proportions. Regarding the suitability of the catalysts for the molecular cleavage of the heat-resistant compounds, the following order generally applies: Catalyst A (catalysts made using ethanol), Catalyst A (catalysts made using methanol), Catalyst A (made using methanol and ethanol) Catalysts), mixture of catalyst A and catalyst B (contains at least 50% of catalyst A based on the moles of potassium), catalyst D, catalyst E (with the addition of about 1 mol of KHS), catalyst B and finally catalyst C. Of course the cracking capacity is increased by increasing the amount of KHS added to the catalysts.

The order of catalysts suitable for hydrogenation is as follows: 0.5 tie. Catalyst A + KOH calculated in moles and then added another 0.5 tin. on catalyst A (very exothermic hydrogenation reaction), then catalyst C with admixed KHS (considering the required degree of cleavage), catalyst B + A (50% or more of catalyst B based on the moles of potassium) and catalyst B.

It should be remembered that in the discussions above, the heat resistance of the carbonaceous starting materials and their components is that which is assigned to these components according to the known prior art. In comparison with the known prior art, the method according to the invention makes it possible to convert all of the materials designated according to the prior art as heat-resistant at lower temperatures and lower pressures and without difficulty and thus in an extremely advantageous manner. However, since the carbon-containing starting materials and their constituents prove to be differently “heat-resistant” in the process according to the invention and, for example, the paraffins are the most heat-resistant and the asphaltenes and aromatics are less heat-resistant, the process according to the invention is carried out in such a way that in a suitable sequence with the Suitable catalysts treated all those components of those carbonaceous materials

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which have so far provided inadequate yields of usable products.

Various reactions are described in the following examples. The examples are not intended to restrict the invention, rather the examples are only intended to explain the applicability of the invention.

example 1

An asphaltene of the type shown below, which had a high softening point of 132.22 ° C and was obtained by solvent extraction, was treated with the following reagent to obtain the product A. The catalyst was Catalyst A described above. When the product obtained by treating the asphaltenes in the first stage was reacted in a second stage, a product labeled B was obtained. The two catalysts were used unsupported.

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distillate

distillate

mixture

Residue

kung

from A + B

A

B

A + B

315.6 ° C

spec. Weight "API 15.56" C

-4.6

31.3

36.7

36.0

8.7

Kin. Visc. 98.9 ° C, cSt

_ +

-

-

0.94

58.0

Conradsen carbon wt%

39.5

-

-

0.20

16.4

Aniline Point, C

-

-

44.6

-

FIA,% by volume

Aromatics§

-

-

71.5

-

01efine§

-

-

Saturated

-

-

-

28.5

-

Bromine number

-

-

-

53

-

Carbon,% by weight

84.24

-

-

83.77

84.81

Hydrogen,% by weight

8.50

-

-

12.52

9.95

Nitrogen,% by weight

0.75

-

-

0.10

0.58

Sulfur, wt%

6.19

-

-

2.31

4.46

Ash,% by weight

0.30

_

_

0.08

0.20

Moisture,% by weight

no

-

-

0.05

no

Oxygen,% by weight

0.02

-

-

1.17

0.00

Nickel, ppm (w)

71

_

33

Vanadium, ppm (w)

174

-

-

-

160

Iron, ppm (w)

151

24th

Insoluble in heptane (IP method)

16.7

-

-

-

+132.2 ° C softening point

§ Aromatics and olefins were probably not clearly separated in the column because of a heavy tail above 315.6 ° C.

The above data clearly show the improvement which can be achieved with regard to the viscosity and the specific weight of the products, the considerable increase in the hydrogen content and the remarkable reduction in the metal content from the later fractions.

The following examples show the results obtained. Asphaltenes of the type specified above and in Example 1 obtained by solvent extraction were used as the feed material.

All experiments were carried out in batches in a stirred container, the inside diameter of which was 15.85 cm and the height of which was 25.4 cm. The stirred tank was equipped with a stirrer and a steam distributor. Steam was generated at a pressure of 2.67 bar (40 lb / sq.in.) In a steam boiler directly connected to the city water supply. The steam reaches the steam distributor via pipes with an inner diameter of 9.53 mm and atmospheric pressure. However, the reactor can also be operated at a pressure of 0.5 bar to about 5 bar or, as mentioned above, at a higher pressure. The steam distributor has a diameter of approximately 8.89 cm and several outlet holes that direct the steam upwards. The distributor is located at the bottom of the reactor.

If a unsupported catalyst is used, a stirrer is provided for the reactor. The motor is located directly above the reactor. The area in which the stirrer shaft is led into the reactor is sealed by a seal. The stirrer consists of twin circles which are connected by angled, curved blades 55. The stirrer can be replaced by baskets which are loaded with supported catalyst in a manner to be described later. Four baskets loaded with supported catalyst are arranged on the agitator shaft. The baskets together with the stirrer shaft result in a total diameter of almost 15.88 cm. The baskets are approximately 15.24 cm high and approximately 12.7 mm thick. The non-hanging basket is a 12.7 mm deep rectangle.

The top of the reactor has a seal in which the stirrer shaft rotates, the riser pipe which De-65 stillatically discharges from the head of the reactor, a pressure blow-off pipe which consists of a valve which operates at a pressure of more than 2.00 bar ( 30 lb / sq.in.) Opens and the reactor contents are blown off into a hood. This pressure relief

663 960

16

direction is also used to load the reactor with a fixed charge.

As a rule, two thermocouples are used in the cover of the reactor. One thermocouple measures the temperature in the lower half of the reactor and the other thermocouple 5 measures the temperature in the upper half of the reactor.

The riser is approximately 22.86 cm high and has an inside diameter of 1.91 cm.

The second stage reactor is a 30.48 cm long tube reactor with an inside diameter of about 3.81 cm. The capacity of this reactor is 347.5 cm3. This reactor is equipped with three radiators wrapped around the reactor. The temperature of each radiator is controlled by a thermocouple via> 5 controllers, which are arranged on a portable control panel.

Gases flowing through the second stage reactor from bottom to top are then passed through a 40.64 cm diameter wash bottle 20 which is not cooled with water.

The first wash bottle is arranged vertically and equipped with a 500 ml collecting flask on the underside. The collecting flask is usually kept at 240 ° C by a jacket heater. The bottom of the catch flask is fitted with a tap to draw off the product.

A second cooler (wash bottle) protrudes from the top of the collecting flask and runs parallel to the first cooler (wash bottle). The second cooler is not cooled with water. The second cooler is also a glass cooler with spherical cooling areas.

The pipe that slopes downward from the second cooler leads to a water-cooled cooler. This water-cooled cooler is arranged vertically, about 45.72 cm long and inserted from above into an unheated 500 ml flask, which is equipped with a drain cock on the underside. A water-cooled cooler leads parallel and vertically from the second fitting of this piston. Another water-cooled cooler is placed directly on this piston.

The uppermost water-cooled cooler is equipped with a 30.38 cm pipe running upwards. This line has a diameter between 12.7 and 19.05 mm. This line leads to a cooler cooled with ice.

This ice-cooled cooler is a double-walled vessel, as is usually used to collect vapors before they enter a vacuum pump. The inner container contains a mixture of water and ice, whereas gases and vapors flow through the outer jacket of the vessel. The gases and vapors enter the bottom of the vessel and exit the top of the vessel. The bottom of the container has a collecting compartment with a content of 50 ml. The collecting compartment is equipped with a drain tap for draining 55 product.

The remaining gases and vapors are sent to another cooler, which is similar to the ice-cooled cooler. This cooler is cooled with a mixture of carbon dioxide snow and 2-propanol. Here too, the product 60 is collected in a collecting compartment which is provided below the cooler and is equipped with a drain tap.

The remaining gases and vapors are then washed with potassium hydroxide solution, which contains 6 mol KOH dissolved in 360 ml water 65. The amount of gas is then determined by passing the gases through a wet gas meter. After this measurement, samples are taken at intervals and the remaining gases are led into the fume cupboard.

For the following experiments, 1300 g of solid asphaltenes are weighed in and comminuted to a grain size suitable for charging the reactor. With protection against oxygen, a liquid or solid catalyst is introduced into the reactor which has previously been purged using helium, for example. About 40 g of the theoretically anhydrous catalyst A are introduced into the reactor.

The second stage reactor is usually charged with about 300 ml of a supported catalyst under the same precautionary measures. The second stage reactor is initially heated to drive off the water from both the zeolite support and the catalyst.

After the water has been largely removed from the reactor space by heating the second stage reactor to above 300 ° C., the first stage reactor is heated.

Asphaltenes obtained by solvent extraction with a melting point of 93.33 ° C. or 204.44 JC were used. The melting point determines the respective form of the asphaltenes.

After the temperature in the first stage reactor required to melt the asphaltene was reached, the stirrer was switched on. As a rule, the stirrer is initially operated at a speed of approximately 30 to 60 min-1.

Steam is generally supplied when a temperature of 220 ° C. has been reached in the first stage reactor. At this point, the reactor of the second stage should have reached a temperature of 424 C or have leveled off at this temperature.

Helium is normally supplied via the steam manifold before steam is introduced into the system to keep the manifold openings clear and to keep the system free of oxygen. The helium is supplied in an amount of about 200 ml / min.

Example 2

1300 g of asphaltenes obtained by solvent extraction were comminuted to such an extent that the particles could be passed through an opening of 25.4 mm in diameter provided at the top of the reactor. The asphaltenes were not heated, but introduced into the reactor in solid form. The reactor was flushed with helium after loading.

The catalyst used was a variant of catalyst A prepared in the following manner. A solution of 1 mol of KOH in 30 ml of H20 was added to the solution of KOH described above, and the resulting solution was saturated with hydrogen sulfide. The solution separated into two layers, the upper one third and the lower two thirds of the total liquid level. The two layers were separated and dried, and the two residues obtained were mixed again. The mixing of the residues can take place in the same ratio as they were obtained (and as was the case in this example), but the quantitative ratio of the two catalysts can be changed. The catalyst obtained by mixing the residues can also be dissolved or dispersed and then deposited on a support. The weight of the catalyst used was, for example, 40 g in the theoretically anhydrous state.

The second stage reactor was charged with a zeolite-deposited catalyst during assembly of the device. The second stage reactor contained about 300 g of carrier and catalyst. The supporting zeolite was L2-Y82 and the catalytic converter

17th

663 960

was the catalyst D, K2S1j5 (empirical), to promote the hydrogenation of the fission product. To start up the plant, it was necessary to heat the second stage reactor to at least 175 ° C. before heating the first stage reactor.

The first stage reactor was then heated. Only the lower half of the reactor was heated, but the upper half of the reactor was not heated. At 220 ° C, a small amount of steam was introduced into the reactor via the steam distributor located on the bottom.

At a temperature of about 220 ° C at the bottom of the first reactor, the production of a hydrocarbon product started slowly and steadily, which was condensed in the flask below the water-cooled twin cooler. However, this product was significantly heavier than the product obtained in the range of 390 to 424 ° C of the working temperature.

As soon as a temperature of 360 ° C. was reached in the bottom of the first stage reactor, there was a considerable improvement in the production speed for the product. The reaction became exothermic so that the temperature rose rapidly and settled to about 415 ° C. This temperature was then maintained in the bottom of the reactor. The temperature at the top of the reactor had reached 360 ° C.

The temperatures in the second stage reactors ranged from 220 to 460 ° C.

When the contents of the first stage reactor were in contact with the stirrer, the process in the first stage reactor proceeded smoothly at about 415 ° C with the temperatures in the second stage reactor fluctuating between 440 and 460 ° C.

The amount of steam corresponded to approximately 20 ml of water transferred to steam per minute. Towards the end of the experiment, the temperatures at the top and bottom of the first stage reactor were allowed to rise to 440 ° C.

The variant of catalyst A described above produced almost no gas in accordance with the measured value of the wet gas meter. The amount of gas was less than 61.

The main part of the product was formed by the bottom product No. 2, which was collected below the water-cooled cooler. Together with the product collected in the water and ice-cooled collecting compartment, this product gave a total of 458 g of product with an API number of 23 (density = 0.9158 at 15.56 ° C.).

Bottom product No. 1 weighed a total of 33 g and had a density of 0.96587 and an API number of 15 at 15.56 ° C. Bottom product No. 1 was collected below the air-cooled cooler.

44 ml of bottoms were collected under the calibrated trap cooled with dry ice and 2-propanol, but only 28 ml could be drawn off because the light tail evaporated. The API number of this light tail was 81 at -10 ° C (density 0.6553 at -10 ° C). Since the evaporation process is rapid, the value given for the density is very inaccurate.

The result was an unusually light coke that separated in the reactor in the form of 50 mm thick layers. The volume of this coke was 1800 ml, but the weight of this coke was only 513 g.

The dead space below the stirrer causes delayed coking below the real reaction zone.

Example 3

This example was operated in a similar manner as in Example 2, but with a different catalyst

Form used and the first stage reactor was heated more quickly.

The catalyst A used for this example provided only a single catalyst layer during its preparation, so that the separation of the two layers formed in the above example 2 was not necessary during drying. A temperature of 420 ° C was maintained in this experiment.

During the 42-minute test period of this test, approximately 40 g of theoretically anhydrous catalyst, after reaching the working temperature of approximately 420 ° C., converted a feed of 1300 g of asphaltenes obtained by solvent extraction into the following soil products:

a) Bottom product No. 1, which was collected below the air-cooled cooler and delivered a total of 29 g of hydrocarbons with an API number of 11.5 (d = 0.9895 at 15.56 ° C).

b) Bottom product No. 2, which was collected below the water-cooled cooler and the water and ice-cooled cooler and which gave 398 g of hydrocarbons with an API number of 29 (d = 0.8816 at 15.56 ° C.).

c) Bottom product No. 5, which was collected under the trap cooled with dry ice and 2-propanol in an amount of 63 ml and had an API number of 83 at room temperature. A significant portion of Soil Product No. 5 was lost in determining the API number.

d) A total of 1351 gas was generated. The amount of gas was measured behind the alkali hydroxide wash of the gas-steam mixture after the trap cooled with dry ice and 2-propanol. These gases were not analyzed, but the average of the gas analysis of similar experiments showed that about 5 mol% of gases other than hydrocarbons such as hydrogen, carbon monoxide and carbon dioxide are produced. The remaining gaseous hydrocarbons have an average molecular weight of 49, so that about 280 g of gaseous hydrocarbons were obtained.

e) In this experiment, the light coke also formed in the experiment according to Example 2 was also produced. The weight of the coke was 489 g.

The products were obtained in the following quantities:

Soil product No. 1 29.0 g Soil product No. 2 398j0 g Soil product No. 5 41.5 g (cooled trap)

Gases 280.0 g

Coke 489.0 g total 1 238.0 g

Eligibility = 95.23%.

Yield = (soil products No. 1, 2 and 5 + gases / feed) • 100 = 57.57%.

In both the example 2 and example 3, the supported catalyst in the second stage reactor was completely clean and free of pitch, carbon, etc.

The total creditability for products obtained according to Example 2 was (1,045.7 / 1,300) - 100 = 80.42%). The yield was (bottoms No. 1, 2 and 5 + gases) (532.7 / 1 300) • 100 = 40.97%.

The fundamental difference between Examples 2 and 3 lies in the significantly higher gas production in Example 3.

In both examples the bottoms separated from the water and condensed the bottoms from the steam, leaving no emulsion.

5

10th

15

20th

25th

30th

35

40

45

50

55

60

65

663 960

18th

Example 4

In this example, a catalyst mixture was used which consisted of approximately 2/3 of the non-separated catalyst according to Example 2 and approximately 1/3 of the catalyst according to Example 3 (potassium base). The various proportions, including the proportions of the catalyst layers in Example 2, can be changed. The unsupported catalyst was used in an amount of about 40 g, calculated as theoretically anhydrous.

The reactor reached a temperature of 420 ° C during this experiment.

The products obtained in this experiment were as follows:

Soil product No. 1

Soil product No. 2

Bottom product No. 5 gas

Total coke

195.8 g (d = 0.9793 at 15.56 ° or API number 13)

309.5 g (d = 0.86 at 15.56 C or API number 33)

28.0 g

276.4 g (133 1-0.95 / 22.4-49 = 276.4 g)

473.0 g 1,282.7 g

Eligibility = (1 282.7 / 1 300) • 100 = 98.67%.

Yield (in soil products No. 1, 2 and 5 and gases) = (809.7 g / 1 300 g) • 100 = 62.28%.

The coke was created in a volume of 480 ml below the stirrer and part of this space is also under the steam distributor.

Example 5

The main difference between this example and the previous examples was the use of a supported catalyst instead of the unsupported catalyst added before the start of the experiment. The catalyst was supported in four baskets on a sintered stainless steel screen. The stainless steel was 3.175 mm thick and was cut into 3.175 mm strips, which in turn were transversely divided into 4.763 mm pieces. The dimensions of the carrier were therefore 3.175 mm • 3.175 mm ■ 4.763 mm. The catalyst used was the same as in Example 2.

The supported catalyst was introduced into baskets measuring 6.35 mm • 63.5 mm • 152.4 mm, four baskets being connected to the stirrer shaft so that the baskets formed the stirrer when the stirrer shaft was rotated. The supported catalyst was held in place by a sieve and the wire mesh baskets were supported by a frame.

When using the same catalyst as in Example 2, the amount of gas generated was reduced from 135 l to 90 l when using a substantially smaller amount of the catalyst in the form supported on the support. Most of this amount of gas emerged towards the end of the experiment when the temperature rose to 440 ° C.

The speed of the stirrer shaft, which carried the baskets, was initially 60 min-1 and was increased to 120 min-1 during the experiment.

gas

The feed consisting of asphaltenes obtained by solvent extraction weighed 1 290 g.

Eligibility = (1 204.5 g / 1 290 g) • 100 = 93%.

Yield = [(soil products No. 1, 2 and 5 + gas) 5 1 290] • 100 = 77.90%.

The API number of the combined soil products No. 2 and No. 5 was 32.0 at 15.56 C. The soil product No. 1 was separated into a portion which was liquid at 200 ° C. and a portion which was not liquid at 200 ° C. The liquid io fraction had a calculated API number of 5 and weighed 228 g, representing 56.03% of the total amount of 407 g of soil product No. 1. The rest was apparently made up of somewhat upgraded forms of asphaltenes obtained by solvent extraction. A new calculation carried out on the basis of the 15 portion of soil product No. 1 that is liquid at 200 ° C. gave a yield or a degree of conversion of 64%.

All of these yield and eligibility calculations were based on the assumption that 20 95% of the gas was formed by a hydrocarbon with an average molecular weight of 49.

It is clear if the feed below the steam distributor changes, a competing process takes place in which catalytic thermal cracking takes place at a lower temperature and the threshold for thermal cracking, i.e. if steam is not present at the same time as the catalyst, is more favorable . If the catalyst is a supported catalyst, a coke is formed when steam, the asphaltene feed and the catalyst are in intimate contact with one another. If an insufficient amount of steam or no steam at all reaches the asphaltenes, the reaction degenerates into catalytically assisted thermal cracking. The catalytic hydrogenating cracking and hydration is much more desirable because of the high yields, the high space velocities and the desirable adjustability of the composition of the product obtained and also because of the reaction conditions, although it is at a lower temperature than the normal thermal cracking at about Atmospheric pressure and about 10 times faster than normal thermal cracking.

45

Example 6

In the procedure described in Example 2 described above, a feed, designated FHC-353, was used from a mixture of residues with different compositions.

Feed analysis:

Soil product No. 1 Soil product No. 2 Soil product No. 5 Total coke

187.0 g (90 1 gas 0.95 / 22.4 • 49 = 187.0 g)

407.0 g 368.0 g 43.0 g 199.5 g 1 204.5 g

50

Caught under 537.78 GC

fraction

API number

Density at 15.56 "C 55 sulfur, wt .-%

Nitrogen,% by weight hydrogen,% by weight carbon,% by weight oxygen,% by weight

60

Total elements: oils,% by weight

Resins,% by weight

Asphaltenes,% by weight

65

Output,% by weight

Ramsbottom carbon,

% By weight

8.6% 6.6

1.0243 3.91 0.478 10.35 84.72 0.471

99.9% 29.6 57.8 12.6

100%

21%

19th

663 960

Metals:

Vanadium

nickel

iron

sodium

viscosity

Pour point

228 52 14 4

ppm ppm ppm ppm

1430 cSt at 27.78 C 250 cSt at 121.11 C 54.44 C

The feed was gradually heated so as not to cause coking as long as the mixture of residues was not yet pumpable or had not yet reached the reaction temperature of about 425 ° C. At this temperature the reaction was carried out in the presence of the catalyst and steam. The feed weighed a total of 15,650 g and, based on the various condensates obtained, the output was 98.1% of the feed.

The boiling point of the feed consisting of residues defined by ASTM D-l 160 and D-l 180 was determined on a sample of 200 ml of the feed in accordance with the standard mentioned. The data obtained were as follows (given as the temperatures observed during the determination).

sample

Ausbrin

temperature

Temperature corrected

Pressure in ml on the colon in the bladder

temperature

Column in

in%

head in C

(D-l 160;

mbar

in C

D-l 180)

40

20th

240

275

975

5.60

50

25th

253

286

810

5.60

60

30th

263

295

835

5.60

70

35

273

304

858

5.60

80

40

283

310

882

5.60

90

45

290

316

5.60

100

50

297

324

907

5.60

110

55

305

330

120

60

312

338

938

5.60

130

65

320

344

953

5.60

140

70

326

351

5.60

160

80

319

347

990

2.66

170

85

330

2.66

180

90

339

364

1030

2.66

188

94

347

1050

2.66

347

190

99

351

380

1090

1.33

This product was converted using a supported catalyst in a stirred reactor in which the catalyst was arranged in baskets. The carrier was a Y-82 zeolite from Union Garbide, Danbury, CT, and was in the form of extrusions, for example flat shaped bodies, of 3.175 mm. The catalyst had the empirical formula K2SL5, etc. Catalyst D prepared above, to which 5% by weight of Catalyst A had been added. The amount of the catalyst was 20.16 g per 220 ml of the carrier described.

The initial conversion of the feed to the first product (overhead from reactor 110) was 71% (the combined products without the very volatile materials and gases trapped in the dry ice and isopropanol cooled trap and gases). This conversion also does not contain those components of the feed that could be distilled off at a temperature below 565.56 ° C. The overheads, including the bottoms separated, were calculated to be 98.1% of the original feed, the remainder being obviously gases and volatile distillates.

The product from reactors 110 and 117 started to boil at 182.22 C and had the following properties:

API number 54.5

Density at 15.56; C 0.7608

Viscosity 0.72 cSt at 37.78 ° C

0.568 cSt at 65.56 ° C 45 ash content 0.009% by weight

Nitrogen content 0.05% by weight

Sulfur content 1.0% by weight

The fraction boiling between 182.22 and 337.78 ° C had the following properties:

API number 34.2

Density 0.8540 at 15.56 ° C

Water 0.002% by weight

Pour point - 28.89 ° C

55 Viscosity 2.57 cSt at 37.78 ° C

1.61 cSt at 71.11 ° C sulfur 2.18% by weight

Ash 0.01% by weight

A sample containing 7 g H? 0 of 1392 g of the overhead products from reactors 110 and 117 (which, however, did not contain the condensates mentioned above) had the following boiling analysis:

Initial boiling point up to 182.22 nC: 71.6 g or 5.24

% By weight 65

182.22 to 343.33 C: 203.4 g or 14.89% by weight

343.33 to 537.78 ° C: 1084.1 g or 79.36% by weight

Total distillate: 1359.1 g

The fraction boiling between 343.33 ° C and the boiling end had the following properties:

API number 17.8

Density 0.9478 at 15.56 ° C

Pour point 46.11 ° C

Viscosity 37.34 cSt at 65.56 ° C

12.05 cSt at 98.89 ° C carbon content 85.03% by weight

Hydrogen content 11.55% by weight

663 960

20th

Conradson carbon nitrogen content

0.26% by weight 2.56% by weight

The residue obtained at temperatures above 565.56 C contained:

Vanadium 5.6 ppm

Iron 48.6 ppm

Sodium 8.2 ppm

Potassium 22.3 ppm

Nickel 0.48 ppm

The properties of the gas oil fraction obtained by vacuum distillation, i.e. the fraction with a boiling point of 343.33 C and above, were as follows:

Start boiling at 344.44 3C and apply the following: 10% to 388.89 C 20% to 416.67 C 30% to 453.33 C 40% to 470 C 50% to 478.89 C 60% to 497.78 C 70% to 508.89 C 80% to 524.44 C 90% to 547.78 C Boiling end 565.56 C with an output of 94.4%.

The K factor of the top fraction passing above reactor 343.33 C (hydrogenation as the last stage) was 11.71.

The bottom products obtained from the first stage reactor 110 at 212a were recycled in batches when a sufficient amount had been withdrawn. The catalyst for feed 212a was Catalyst A deposited on steel sieve particles measuring 3.175 mm • 6.35 mm • 4.76 mm and contained 15.7 g of catalyst with a total volume of 200 ml. A total of 1286 g sample was used. The total output was 750.7 g (including 8 g of product from the dry ice trap without gas). About 29% of the FHC-353 feed was obtained as bottoms. Feed 112a analysis (bottoms) was as follows:

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 48.4% of the feed was converted. Continuous work would result in a higher degree of conversion, but in batch work, because of the dead space in the reactor underneath the surrounding cages or baskets, no continuous reaction space is permitted and also because of delays in the connecting lines, a completion of the reaction is prevented.

The combined products obtained have the following properties:

Carbon 82.24% by weight

Hydrogen 11.95% by weight

Oxygen 0.64% by weight

Water 2.8% by weight

viscosity

Nitrogen water

8.533 cSt at 37.78 ° C 4.028 cSt at 65.56 ° C 0.17% by weight 2.8% by weight

The fraction having a boiling range from 343.33 C to the boiling end had the following properties: API number 17.4

Density 0.9444

io pour point 35 ° C

Viscosity 21.54 cSt at 65.56 ° C

7.989 cSt at 98.89 ° C carbon 84.7% by weight

Hydrogen ll, 27% by weight

15 oxygen 0.66% by weight

Conradson carbon 2.2% by weight

Nitrogen 0.27% by weight

Vanadium 3.5 ppm

Iron 7.6 ppm

20 sodium 1.6 ppm

Potassium 10.3 ppm

Nickel 0.45 ppm

Distillation residue: 25 carbon hydrogen oxygen sodium vanadium 30 nickel iron potassium

85.22% by weight 11.06% by weight 0.49% by weight 3.07 ppm not determined 0.344 ppm 4.6 ppm 27.4 ppm

The product passing above 343.33 C (from 35 Bischickung 112a) had the following boiling analysis for a sample of 191 g, from which 163 g were withdrawn overhead; it was a 175 ml sample at 15.56 C. The fractions collected were as follows: 10% to 343.3 C.

40 20% to 412.78 C

30% to 443.33 C

40% to 465.56 C

50% to 486.11 C

45

80% 85%

to until

554.44 565.56

This distillation product had an API-50 number of 20.0 at 15.56 C and the residue product had an API number of 16.4. The low metal content and the rather high API number make this residue a good feed, for example for catalytic cracking.

The residue, which makes up about 15%, forms the bottom product in the reaction vessel, for example 130, and can be subjected to further treatment, for example in a further parallel stream, for example the same as for the feed 112a. The analysis of the residue in the reaction vessel was as follows:

60 carbon 83.99% by weight

Hydrogen 6.88% by weight

Oxygen 2.9% by weight

Water 0.05% by weight

The boiling range from 182.22 to 343.33 C - 65 de fraction had the following properties:

API number 27.2

Floating point - 26.11 C.

Based on the 58.37% conversion (product obtained divided by the weight of the feed which can be counted with 83%) (return of the bottom products from the reactor 110, feed 112a) and on the top products

of the reactor 110 is the total conversion, without taking into account the liquids in the last dry ice trap and the gases drawn off with the overhead products of the reactor 110 and also the putatively distillable products (which are presumably still contained in the feed FHC-353 and below 565.56 ° C distill), 87.9%, but the actual yield is even higher if the gases and the products trapped in the dry ice trap with an average molecular weight of 49 and their volume are converted to weight, including the actual FHC-353 content of distillate will.

The yields given in the examples above were calculated so that they can be interpreted in a manner which corresponds to normal practice. The conversion can also be calculated in the form of a space-time velocity, that is to say as the starting material converted per unit volume of the catalyst per unit time.

The yields are usually based on the total amount of starting material used, which was reacted together with a fraction of a component which has a boiling point of more than 565.56 ° C. The boiling points were determined according to the methods described above, which are common in the petroleum industry. In the petroleum industry, the limit value of 565.56 ° C is used to describe the very heat-resistant components of crude oils and distillation products using conventional methods.

The yields can also be based on a proportion of those constituents which have been converted to products of lower viscosity, for example products based on the API number. The residues collected, for example, in the cyclones 116 and used as feed 112a are included in the yields as additives to the initial amount of treated and obtained starting material. The residue from the vessel 130 boiling at 565.56 ° C. or above is then, if it is not further treated, subtracted from the total collected material.

Occasionally, the yields on the treatment of residues possessing certain boiling points, etc. To obtain residues with a boiling point of more than 343.33 C and less than 454.44 ° C or between similar limits and then only to consider the fraction which can be converted by customary methods (whichever method is used in the special refining device). Such a calculation method is not useful for the evaluation of the method according to the invention, since almost all of the previously non-convertible residues are converted in the method according to the invention

663 960

to characterize products obtained in the first and second (if necessary) parallel stream boiling below 565.56 ° C. as additional amounts. I.e. that after the first stage products have been combined with the products of the first or second parallel stream and with the products from any other reaction vessel, e.g. 117 etc., and also with all condensates etc., the whole is distilled in the manner described above and the residues boiling above 565.56 ° C. are subtracted and the conversion is calculated without taking these residues into account, whereupon the yields are determined.

In practice, it is best to consider the gaseous hydrocarbons that boil below 565.56 ° C in the feed, also taking into account the low temperature condensables and gases. Gas can be taken into account arithmetically by calculating an average molecular weight from the gas volume collected and then converting the value obtained into weight percent so that the gas quantity can be credited to all products.

According to the above description, it can be assumed, based on Example 6, that by the specific treatment of those separated from crude oil, residues with a low boiling point up to 454.44 ° C or residues with a high boiling point of 565.56 ° C or more or separated fractions in a pass at low pressures and low temperatures, almost all of the crude oil or other hydrocarbonaceous materials can be converted into usable products if the bottom products are subjected to a second or a further treatment by means of specifically selected catalysts.

So far, considerable efforts have had to be made to convert the hard-to-convertible proportions of crude oil or petroleum residues into usable components, but the yields have been low and numerous recirculations have been required in order to achieve acceptable yields. If one takes into account that up to 50% of these parts, which are difficult to convert, had to be subjected to the very thorough treatment, it becomes clear that the parallel process according to the invention, which is carried out using inferior starting materials or by treating the cleavage products separately under less severe conditions, and considerably higher Yields and better quality end products and a high space-time speed, etc., makes a remarkable contribution to the constant effort to obtain all usable products from a given volume of crude oil or its residue.

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2 sheets of drawings

Claims (27)

663 960 2nd PATENT CLAIMS 1. A process for converting carbon-containing starting materials containing heat-resistant constituents to products of low viscosity and / or more hydrogenated end products in the presence of a catalyst without the addition of gaseous hydrogen, characterized in that (a) that at least one carbon-containing starting material is reacted in a first reaction zone in the presence of water added in the form of steam and a catalyst, the catalyst containing an alkali metal sulfide catalyst and at least 5% of an additional cleavage catalyst in order to remove those present in the starting material in the first reaction zone pretreating heat-resistant ingredients, and wherein the cracking catalyst is mixed with the alkali metal sulfide catalyst to treat the heat-resistant ingredients in the raw material, and wherein the cracking catalyst is a supported catalyst or a supported catalyst from a solution (A) of an alkali metal hydroxide in methanol, ethanol, 1-propanol or 1-butanol or mixtures of these alkanols or a solution (A) to which alkali metal hydroxide dissolved in water has been added and in which the ratio of solution (A) to added alkali metal hydroxide, on a molar basis, is from 0.5: 1 to 1: 0 , 5 is, and this solution is saturated with hydrogen sulfide in such a way that in each solution (i) a single phase is formed or (ii) a two-phase solution is formed, the catalyst comprising the aforementioned single-phase solution according to (i), the two-phase solution according to (ii), one of the phases according to (ii), mixtures of the phases according to (ii), a mixture of each of the individual phases according to (ii) with the single-phase solution (i) or a mixture of the two phases according to (ii) with one another and with the single-phase solution (i), (b) a top product is withdrawn as a reaction product from the first reaction zone and contains gases as vaporous or gaseous products, and bottom products are separated off, (c) that bottom products are recycled or separated as bottom products in a parallel stream from products from the first reaction zone, these products being liquid products or partially pretreated heat-resistant constituents of the starting material, (d) in the presence of initially supplied steam or steam, the separated bottom products are reacted in the presence of a cracking catalyst of the type defined in (a) above and the alkali metal sulfide catalyst and (e) that the products obtained in stage (d) are obtained. 2. The method according to claim 1, characterized in that the catalysts used in stage (a) and stage (d) are supported catalysts with a metallic support. 3. The method according to claim 1, characterized in that the products from step (c) are recycled instead of being worked up separately. 4. The method according to claim 1, characterized in that the vaporous or gaseous reaction products (top products) from stage (b) or stage (d) directly or after cooling in at least one further reaction zone in the presence of a supported catalyst and added steam for the purpose of hydrogenation are, wherein the supported catalyst contains an alkali metal hydrosulfide, sulfide, polysulfide, sulfide hydrate or polysulfide hydrate or mixtures thereof, so that the bromine number is reduced compared to the bromine number of the products from step (b) or step (d). 5. The method according to claim 1 or 3, characterized in that the vaporous or gaseous products from stage (b) are reacted directly or after cooling in the presence of a supported catalyst and added steam in at least one further reaction zone for the purpose of hydrogenation, the supported catalyst being a Alkali metal hydrosulfide, sulfide, polysulfide, sulfide hydrate or polysulfide hydrate or mixtures thereof, so that the bromine number is reduced compared to the bromine number in products from step (b). 6. The method according to claim 1 or 5, characterized in that the catalyst used in stage (a) or in stage (d) is a supported catalyst and 1 / a catalyst A or 2 / the catalyst A in an amount of up to 95%, based on mole K in catalyst A, together with a catalyst B, C, D or E or mixtures thereof, and in that catalyst A is prepared by adding a mole of potassium hydroxide in ethanol, methanol, a mixture of ethanol and methanol, 1- Dissolved propanol or 1-butanol and the potassium hydroxide solution is reacted with hydrogen sulfide by passing hydrogen sulfide through the solution and the mixture of catalyst and alcohol is obtained and the alcohol is driven off from the solution, that the catalyst B is prepared by dissolving technical or analytically pure potassium hydroxide containing about 86% potassium hydroxide in ethanol or methanol and saturating the solution in a series connection of vessels by introducing hydrogen sulfide into a first vessel without boiling the alkanol with hydrogen sulfide an exothermic reaction between the potassium hydroxide and the hydrogen sulfide driven off alkanol is collected and trapped in a downstream vessel and the supply of hydrogen sulfide is stopped as soon as a reaction appears in the last vessel containing KOH, that the catalyst C is prepared by dissolving about 6 moles of KOH in 4.5 to 7.5 moles of H20 without external heat, then adding a small amount of an alkanol of, for example, 2 to 2.5 ml of methanol or ethanol per mole of KOH is that about 4 moles of elemental sulfur are then added in order to adjust the desired sulfur content of the catalyst by adding this additional sulfur and to obtain a sulfide of the empirical formula from K2Sy to K2S5, that the catalyst D is prepared by dissolving 1.0 mol of KOH in 1.0 mol of water, adding 2 ml of methanol or ethanol immediately after dissolving the KOH, then adding 0.6 mol of elemental sulfur and then the catalyst by adding sulfur with further stirring to the desired empirical sulfur content, which corresponds to an empirical formula in the range of K2Si, i to K2S5, and that the catalyst E is prepared by dry KHS powder or a slurry thereof from each of the above described catalysts or mixtures of catalysts A, B, C or D is added, the addition being carried out in an amount of 1/5 to 1/3 mol, based on moles K in KHS, in order, on a mole basis, of a sulfide the empirical formula K2S to get to the empirical formula K2S2.5. 7. The method according to claim 1, characterized in that the catalyst used in step (a) or in step (d) is a supported catalyst and the support is a porous metal, a chromite spinel, a zeolite or an alumina. 8. The method according to claim 1, characterized in that the catalyst used in stage (a) or in stage (d) 10th 15 20th 25th 30th 35 40 45 50 55 60 65 3rd 663 960 Tor is a supported catalyst with a porous metal as a support and the porous metal is a stainless steel with up to 35 vol .-% metal. 9. The method according to claim 6, characterized in that the plastic-containing starting material reacted in step (a) is a heavy crude oil or a residue not having a boiling point or an asphaltene and the supported catalyst is a catalyst A in combination with at least 75% by weight having catalyst, which is deposited on a catalyst support. 10. The method according to claim 9, characterized in that the catalyst A is present in an amount of 5 to 80 wt .-% and the rest of the catalyst D is formed. 11. The method according to claim 6, characterized in that the catalyst used in step (d) is a supported catalyst made of catalyst A supported on porous metal and the reaction is carried out at a temperature of 360 to 440 ° C. 12. The method according to claim 1, characterized in that the reaction at a pressure in the range from subatmospheric to a pressure of less than 10 bar (150 psi) and at a temperature of 160 to 600 C in the presence of a porous metal supported catalyst is made. 13. The method according to claim 6, characterized in that the catalyst used in the further reaction is of the composition of catalyst A, the KHS is mixed in an amount of up to 15 wt .-% and the catalyst obtained on a support made of porous stainless Steel is deposited. 14. The method according to claim 1, characterized in that the carrier is a porous metal, a Chromitspi-nell, an alumina, a zeolite or a mixed carrier, which carrier is protected against attack by the catalyst. 15. The method according to claim 14, characterized in that the catalyst is a supported catalyst and the support is an alumina with a pore size of 0.6 nm to 1300 nm, preferably 35 nm to 90 nm, the alumina thereby preventing an attack by the Protected catalyst, that the catalyst is deposited as a mixture with glycerol or a polyhydric alkanol having up to 6 carbon atoms and then calcined in the absence of oxygen, or that the glycerol or the polyhydric alkanol is deposited on the carrier, the carrier heated to 200 C, which Cooled carrier and then the catalyst is deposited on the carrier, whereupon the catalyst is heated to a temperature of about 560 ° G. 16. The method according to claim 1, characterized in that the carbon-containing material in a second reaction zone is formed by a residue of a first reaction zone, which was collected in a cyclone zone kept at 390 ° C. 17. The method according to claim 1, characterized in that after recovering the product from step (c), part of this liquid product is split and reacted further in at least one further second reaction zone (d) and in that a bottom product from the reaction zone (d) in at least one other reaction zone. 18. The method according to claim 1, characterized in that the reaction in stage (a) and in stage (d) is carried out adiabatically at a temperature up to 560 ° C. 19. The method according to claim 1, characterized in that the reaction products of a crude oil residue are reacted in said zones immediately after stage (a) with quenching the reaction products from stage (a) leaving as top products and that a reaction product leaving stage (d ) is reacted with quenching in a further reaction zone for the hydrogenation of this product. 20. The method according to claim 1, characterized in that in step (a) or in step (d) the reaction with the catalyst supported on a porous metal is carried out in a fixed bed reactor, a fixed bed reactor, a liquid bed reactor or a fluidized bed reactor. 21. The method according to claim 1, characterized in that the reaction in step (a) is carried out in a boiling-bed reactor with a catalyst supported on zeolites and the catalyst contains 25% of catalyst A and 75% of catalyst B. 22. The method according to claim 6, characterized in that the reaction in stage (d) is carried out in a boiling-bed reactor and the catalyst is supported on a zeolite which is a carrier and the catalyst consists entirely of catalyst A. 23. The method according to claim 6, characterized in that the reaction in step (a) is carried out in a stirred basket reactor or in a boiling-bed reactor and the catalyst is a mixture of catalyst A and catalysts B, C and D or mixtures of the latter, where the amount of catalyst A is 3 to 95% by weight and the catalyst is supported on a zeolite, a porous metal, a chromite spinel or alumina. 24. The method according to claim 6, characterized in that the reaction in step (d) is carried out in a stirred tank reactor or boiling-bed reactor and the catalyst is a mixture of catalyst A and catalysts B, C or D or mixtures of the latter, the amount of catalyst A is 50 to 100% by weight. 25. The method according to claim 24, characterized in that the catalyst A is present in an amount of 50 to 95 wt .-%. 26. The method according to claim 6, characterized in that the top products of the reactions in steps (a) and (d) are reacted further before the hydrogenation. 27. The method according to claim 26, characterized in that the top products from stages (a) and (d) by means of a catalyst A mixed with KOH, a catalyst C mixed with KHS, a catalyst B or one with 50% or more, based on mol K, catalyst B mixed catalyst B and A are hydrogenated.