NO840773L - DIVOLUTION AND HYDROGENERATION OF RESISTANT PETROLEUM REMOVAL PRODUCTS, LIKE ASPHALTENES, RESINTS OL - Google Patents

Description

The present invention relates to the treatment of petroleum and petroleum residues, in particular the invention relates to the treatment of pure petroleum and petroleum residues in general, but more particularly high-boiling residues according to the definition given in ASTM D-1180 and D-1160, or residual fractions without any boiling point according to the same definitions , or as destructively distilled with the formation of carbon residues. Furthermore, the invention relates to an independent treatment of the component parts in the above-mentioned petroleum fractions and petroleum residues to improve the total yield, and the space-time velocity, which includes, among other things, the use of somewhat higher pressure, i.e. up to NOK 10.7 per cm2. to further facilitate the recovery of the desirable products.

A number of attempts have generally been made in the past to upgrade heavy petroleum fractions, but with limited success. Many petroleum components such as residues are particularly resistant to further processing. However, by distilling all the oil components from the bottom, e.g. by means of steam distillation or the like, or by further processing these by means of hydrogen transfer reactions, partial coking, delayed coking and the like, it has been possible to improve these bottom fractions and thereby obtain further components from them. By e.g.

to use a hydrogen-yielding solvent, then some of the hydrogen from the solvent will be transferred to the high-boiling petroleum residue fractions, thereby improving the viscosity of part of the high-boiling fractions so that lighter products are obtained. Typically, these reactions will be converted at very high temperature, and with or without the presence of hydrogen under high pressure.

Other methods have also been proposed whereby hydrogen produced by the distillation of hydrogen, which is part of very resistant residues, has been used in a parallel flow, and where the hydrogen has been used as an ireactant gas, after which the coked product has either been used for other purposes or burned. The excess hydrogen can also be used as such in said hydrogen-releasing solvent system, or elsewhere in the refining cycle to improve the quality of the residues.

In an earlier application US-PS no. 140 604, published 18 November 1981, and in Great Britain as UK PS no. 2 075 542, a method for the hydrotreatment of petroleum residual materials is described, where the residual material is contacted with steam and with alkali metal hydroluphides or empirical monosulfides or polysulfides and/or their hydrates, or mixtures of such compounds, to hydrocrack, hydrogenate, hydrotreat, denitrogenate and/or demetallize and/or desulfurize the carbonaceous material. According to the procedure described in the above-mentioned application, a significant improvement has been achieved with regard to the treatment of petroleum and the yield of petroleum components, based on yield, conversion per review, space-time speed, API numbers and the like, so that the method produces upgraded products of greater value.

In US PS 4 366 044 and 4 366 045, various methods relating to the treatment of coal are described, and a number of references to previously known methods are indicated there. The two aforementioned patents and the references indicated therein provide background information regarding alkali sulphide chemistry in general, and also partially include references relating to the treatment of crude oil and the refining of petroleum residues. These references describe different methods to improve the treatment of petroleum residues or to treat different types of carbonaceous material.

Other previously known methods based on the US Patent literature are summarized by M.J. Satrian, Editor "Hydroprocessing Catalyst for Heavy Oil and Coal" Noyes Data Corporation, Park Ridge, New Jersey, U.S.A. 1982, especially pages 62 to 78. In the US Patent literature the following patents have been cited (although some of these are mentioned, among other things, in the above references): US PS 3,520,826; 3,775,346; 3,850,840; 4,117,099; 4 119 528 and 4 203 868. However, none of these previously known methods include the present invention which represents a new discovery, and where the present method is different from that described in the above-mentioned UK PS 2 075 542.

Other less relevant previously known methods are i.a. described, e.g. in Fuel, Vol.62, Feb.1983, which is an issue exclusively concerned with fundamental studies with respect to catalytic coal and carbon gasification (cf. Kikuchi et al., "Supported Alkali Catalysts for Steam Gasi-fication of Carbonaceous Residues from Petroleum, ibid., pp.226 and following, which indicate carbon deposits on bored catalysts, general attempts at a substantial gasification of carbon, but which have no references to the role the sulphide series of the alkali metal can play in these reactions).

Other prior art methods are described by Sikonia et al., "Flexibility of Commercially Available UOP Technology for Conversion of Resid Distillates", National Petroleum Refiners Association (NPRA) Publication AM-81-46, 1981, Washington, D.C. USA, Ritter et al., "Recent Developments in Heavy Oil Craking Catalysts", Publication AM-81-44, NPRA, 1981, Bartholic et al., "Utilizing Laboratory Equipment in New Residual Oil Development", Publication AM-81 -45, NPRA, 1981.

The present invention relates to a method using a boiling bed reactor and supported catalysts, as well as a method for treating particularly resistant petroleum residues, e.g. those that boil above 454°C, and especially the asphaltenes or ORA fractions, which are a mixture of oil, resin and asphaltenes of the type left over from petroleum refining and which are characterized by boiling at 537.7°C or higher according to ASTM D-1180 and D-1160 method. This aspect of the invention represents a further advantageous transformation of particularly resistant bottom products or residual products.

Furthermore, it is a purpose of the present invention to further provide advantages in the method described above. This aspect includes an easier and more advantageous treatment of petroleum or crude oil, such as a starting material in addition to less completely refined petroleum and crude oil residues, so that when you process crude oil, you will essentially not get any residues, and the starting material will essentially be converted to lower boiling products. This result is achieved with the help of the present method because crude oil is not subjected to the vigorous treatment that results in commonly known petroleum residues.

A specified boiling point, as described here, is understood to mean a boiling point such that the product will essentially be completely distilled at the boiling point (according to the defined methods), up to the specified temperature, and you will not recover significant amounts of residues if nothing else is specified.

Brief description of the invention.

The present method differs from previously known methods by the specific types of catalyst used, preferably in a supported form, to attack extremely resistant petroleum residue components, e.g. of the type that boils above 537 .7°C, and the process can be used for an efficient conversion of these residual components into valuable products of lighter viscosity.

The present method further differs from previously known methods in that the reaction particularly attacks the most resistant components in the residues,

such as the asphalts, while avoiding, due to the discovery of a catalytically facilitated thermal decomposition using the catalysts described here, the unfavorable effects arising from coke formation, which can occur if the method is not carried out properly. Other advantages are achieved by using a boiling bed reactor, and by carrying out the process continuously and at higher temperatures, i.e. up to 650°C, while simultaneously cooling the conversion product whereby a desired end product is obtained in one or more other reactors in combination with the first reactor.

In accordance with the present method, the ORA fraction obtained from residues, e.g. of the type arising from petroleum refining or refining other carbonaceous materials, which give these fractions, are advantageously treated according to the present method. However, the best results are obtained when the present method is used for the treatment of asphaltenes, i.e. solvent-extracted asphaltenes when these have been treated in one or two stages for further cracking and/or hydrogenation, thereby improving the total yield that can be obtained from a barrel of oil.

The present improved process is also characterized by its ability to specifically catalyze a conversion of Ramsbottom or Conradsen carbon into usable hydrogenated products without affecting the present process to a particularly significant extent.

Furthermore, the present method provides particularly advantageous substrate combinations for catalysts and where these can be used in a fluidized bed reactor and provide conversion products in a particularly advantageous manner, without the substrate combinations for the catalysts being affected by unwanted metal constituents which are often found in petroleum and which essentially accumulate in ORA - faction.

Furthermore, it has been found that after a first-stage conversion with the active catalysts, it is possible to provide another cooling stage where the separated product can be processed further with the help of a specially selected catalyst in another reaction stage, whereby a predetermined and/or or highly beneficial product. The second stage reaction is, however, dependent and based on the specific catalyst in the first stage, and is based on a suitably completed reaction in the first stage reactor.

Furthermore, the present method enables a recycling of the product which has not been suitably reacted in the first reactor, and which can then be converted into a desired end product. Furthermore, it has surprisingly been found that the present method can be further improved by parallel treatment of different products which are first obtained from a boiling or a fluidized bed reactor, and split into at least two streams, as will be described in more detail in the following. It has surprisingly been found that this method of splitting the product streams (which is a straight through treatment for the first gaseous phase or vapor phase, i.e. the top products) for the bottom fractions by a separation device such as a cyclone, further contributes to the total conversion of the starting material , which happens by separately treating the heavy fractions that are carried over, e.g. from the boiling layer. This is partly due to the fact that different catalyst combinations are used to avoid certain disadvantages in the catalyst compositions when these are used to treat highly varying components of the original feed, or recycled parts thereof. Furthermore, advantages will be achieved with the bottom fractions that have been pre-treated to a certain extent (e.g. something that can be characterized by a lower API value) by this parallel treatment.

This separate treatment of at least one parallel flow of a smaller, partially processed fraction, such as e.g. comes over from a boiling bed reactor, offers a number of advantages. The catalyst movement is from approx. 5 to approx. 100 cm per second with a residence time of 10 to 30 minutes, with an average of approx. 43 cm per second and a residence time of 20 minutes for the particles. (The catalyst movement is defined according to the technology used in boiling bed reactors). Further boiling is produced by means of steam and feed which is fed into the reactor. At the same time, this treatment will make it easier to achieve improved throughput, yield, conversion, type of products etc. for the top products from the first step, so that the applicability of the process for treating a number of different starting materials will be improved by this parallel treatment, as described here , which in particular makes it possible to split out certain fractions and process these, so that they can be adapted to the different starting materials, which are processed with different component compositions, besides making it easier to process mixed product starting materials that are particularly difficult to process.

The aspect of the present process where the heavy bottom fractions can be recycled, but without parallel treatment, gives very good results, but a further improvement has now been found. The former process is often called a "recirculating process", while the improved process is called a " parallel method". A separate treatment without recycling is thus more advantageous for a total improvement of the product and product yields that can be obtained from a number of different heavy crude oils, petroleum refining fractions, and especially the more resistant component parts. The parallel method described in the present application in reality provides almost a total recovery of all the crude oil components as usable products.

The present method is characterized by obtaining an improved product from a very poor starting material with a high hydrogen content, and where the cleaved products have low viscosity, and smaller molecular size, and can be more easily processed in commonly known processing steps.

The present invention includes a significant hydrogenation which improves the yield, product composition and gives smaller molecules and a less viscous product, and eliminates or reduces to a minimum the presence of free Conradsen or Ramsbottom carbon. With the help of the present method, it is now possible to use all fractions in a refinery product which previously could not be advantageously processed to the extent described here. The method described here, i.e. which consists of the "recycling method" and the "parallel method", is thus particularly advantageous for the petroleum fractions that were previously converted using previously known methods, under very strong conditions, can be converted using the ones described here catalysts in a milder way, than in a recirculating or in a parallel method. The very persistent bottom products which make up a smaller part of the total mixture and which are difficult to treat, were previously either pre-treated and recycled, instead of being drastically attacked during the first treatment, i.e. reboiled as a bottom fraction and/or recycled, can now be treated in a parallel treatment with the required degree of load. This processing load for the recycling method or the parallel method is markedly and significantly less than that required in previously known methods, while at the same time achieving far better results, which can be characterized by more valuable product components, better recovery of the product and simpler recovery steps for the components, as well as a number of other advantageous features which will be described in more detail in what follows.

The method according to the present invention can also be applied to other products such as Gilsonite (also known as uintait, a black, glossy asphalt occurring in Utah), pitch (natural and/or as a refinery product), various types of tars, e.g. of the type found in tar lakes in California, asphalts (natural) e.g. of the type found in asphalt lakes in Trinidad and of the type found in Utah, as well as a number of other natural products or products obtained after refinery treatments where a concentration of oils, resins and asphaltenes occurs in residual components, suspended oils of the type which is often without a distillation point• bottom oils from coal gasification plants, e.g. solvent extraction products from coal, or coal extracts of the type shown by Satrien, reference as shown above, page 206, or bottom fractions after treatment of shale oil; heavy bottom fractions from a Lurgi gasification process, and where these fractions are produced during operation of the SASOL process in South Africa; so-called delayed coking products which are obtained during bottom distillation of these products, resins of both the hard and soft type, besides asphaltenes. It is not advisable to mix some of these starting materials, but this can be determined in each case by an experiment carried out under the real reaction conditions. Heavy crude oils are crude oils with an API number of less than 22, and this includes a number of API crude oils with a minus number.

The combination process described here is also advantageous in that the more liquid fractions from petroleum and similar hydrocarbons can be treated by direct distillation, while the heavier bottom fractions containing the heavier components, especially the more resistant components, can be treated with advantage by a recirculating treatment or a parallel treatment depending on the load that is required. Thus, e.g. the starting materials that require the strongest load are best processed by the parallel method. One thus chooses the method which is most suitable for the starting material and which will not impair the yield from the first reactor system, while at the same time improving the recovery of bottom products by a subsequent recirculating treatment or parallel treatment, whereby an improved total yield is obtained.

Although the "parallel method" has been described with reference to two successive reactors or a parallel treatment of two streams, the method can also be used for the treatment of stable products obtained from another reactor^

and which can be further processed without recycling.

As a result of the combination described above

of stages, the present invention provides advantageous space-time rates, yields on an absolute basis, yields per parallel pass, or combined yields.

As an example of parallel processing one can mention

that products obtained from a very heavy bottom fraction of particularly resistant products gave end products with a K-factor of approx. 11.71, i.e. highly paraffinic products, which was achieved more easily than a subsequent treatment of the top fractions from the first reactor, e.g. to reduce the bromine number or the like.

Description of the present method by means of

a process diagram, and a description of the attached drawings.

On the attached drawings are:

Figure 1 a schematic illustration of the recycling method in a boiling bed reactor, and that

figure 2 schematically shows a continuous process, where different starting materials are converted into usable products using the parallel method.

As shown in figure 1, the first stage reactor is a boiling bed reactor 10. It consists of a reaction vessel 11 in a collecting funnel 15 on top, and a pump 14 for recycling the liquid 8 which is subjected to a reaction. The catalyst 9, which is a dispersed-borne form, boils together with the liquid 8. A liquid 8 which undergoes a reaction together with the catalyst 9 is therefore circulated by pumping the liquid into the reactor and distributing it there. Typically, the circulating liquid will be introduced at the bottom, but the supply can also take place at other places in the reactor. On a smaller scale and very close to a boiling bed reactor,

then a stirred tank would be appropriate where a supported catalyst material has been placed, either in small cages or in baskets, e.g. 4 stainless steel baskets made of a fine mesh, and these are attached to a suitable frame driven by an external motor. By varying the rotation speed, e.g. from 20 to 180 revolutions per minute, more typically from 50 to 150 revolutions per minute. minute, then reactions are achieved that are very close to those achieved in a boiling bed reactor. It is again of course important that a suitable contact between steam catalyst and liquid is achieved in order to achieve the desired result.

You can also use circulating fluidized bed reactors where a supported catalyst circulates together with the liquid, or you can also use fluidized bed reactors. In a similar way, a fixed bed reactor can be used where the liquid flows either upwards or downwards.

A continuously supplied preheated starting material such as asphaltenes or an ORA fraction is supplied via pipe 12. Steam is supplied via pipe 13 together with the supply, or can also be distributed inside the reactor 10 in the liquid at the bottom of the reactor.

The produced liquids of lighter viscosity are fed by means of a large diameter pipe 15a and suitably cooled if desired in a cooler zone 16, and then fed into a second stage reactor 17. Although only one reactor is shown in the figure , then one can of course use a number of reactors in series or in parallel. These reactors can have a fixed catalyst bed, or they can be gas or vapor phase reactors. A typical column 17a where the catalyst is carried in a trough is shown in the drawing. You can also use other known devices for this purpose, such as reactors with a seeping layer and the like. The bottom fractions from the second reaction step are collected in zone 18. These are classified as No. 1 bottom fractions. These bottom fractions can be fully or partially recycled to the boiling reactor together with the liquid collected in pipe 15. If necessary, a pump 14a can be used for this purpose. Part of the product can also be diverted and recovered for further processing. The top fraction from the second stage reactor 17 can be refluxed using column 19, and the products can be diverted from this column and the gaseous products can be further worked up in another reactor, e.g. of the type indicated for 17, but not shown here, and all these second stage reactors differ from first stage reactors in that the starting material is split. Products and other gaseous products from the second-stage reactor can be processed in a further reactor, by e.g. bubbling through a suitable bath to remove unwanted constituents such as hydrogen sulphide. This can be carried out in a vessel 20 where potassium hydroxide has been dissolved.

The gaseous products can then be recovered in a conventional manner.

As shown in figure 2, the first stage reactor 110 is a boiling bed reactor. It can also be a reactor with stirred baskets or cages, or a fluidized bed reactor. The rotating device itself is not shown in the drawing. Since a fluidized bed reactor or a stirred cage or basket type reactor are very similar in terms of reactions, they will be discussed here as if they were the same.

The reactor 110 consists of a reaction vessel 111 with collection devices in the form of a funnel 115 on top, and a pump 114 to recycle the liquid 108 which is subjected to a reaction. The starting material or liquid 108 can be petroleum or light petroleum residues that boil below approx. 454.4°C, it can also be ORA (oil, resin, asphaltene) that boils above that temperature, e.g. above +565.5°C, i.e. which has no boiling point. A catalyst 109 that is dispersed and carried in form boils together with the liquid 108. A sieve 109 prevents the catalyst from flowing up together with the boiling liquid 108. The liquid that is subjected to the reaction and thereby forms part of the catalyst, however, cannot or does not necessarily need to circulate by it is pumped into the reactor via a pump 114, because the catalyst must be suspended in the liquid (for circulating fluidized bed reactors), and one can possibly only collect a liquid phase free of catalyst in the funnel 115. As shown in the drawing, the boiling liquid is pumped via pump 114 and it can be supplied at any place in the reactor, typically, however, this will happen near the bottom of the reactor.

The supply 108 of the type described above is fed in through a single supply pipe together with steam and/or water 113, or one can have several supply openings along the periphery of the vessel 111. The permanent components in the petroleum that have a longer residence time are subjected to a catalyzed reaction and is led out of the reactor via line 115a. This line leads directly to a cyclone 116 which separates the vapor phase products from the liquid phase and/or entrained products are carried over with the steam from the boiling layer. Often more than one cyclone 116 can be used. The cyclone bottoms 112a are processed in a separate parallel stream.

The top products from the cyclone 116 can be cooled in a cooling zone 116a as shown in the drawing, after which they are immediately reacted in a reaction vessel 117. It is not necessary for the products to be cooled before they are fed into the vessel 117, as the cooling can optionally take place there. The reason for carrying out a cooling will be described in detail in what follows.

After the products have been introduced into vessel 117, suspend

one a suitable catalyst which will be described hereinafter, either in supported or unsupported form, as shown schematically in trough 117a or otherwise. The catalyst will produce a reaction in the gas or liquid stream, so that a lighter fraction and a heavier fraction are again produced, or the gaseous stream can be hydrogenated. The bottom products can be recovered in a vessel 118, and further processed in a known manner, or as described here, processed separately in another parallel flow.

Although it is not shown in the drawing, additional reaction vessels can be used in the same stream of the type shown as 117. The final recovery of gases from the reflux cooler 119 or other coolers (not shown) can be carried out in a commonly known manner, e.g. .ex. by cooling and/or low-temperature cooling and/or pressure condensation as required.

It has been found advantageous that the gaseous products from the first direct distillation, after treatment in vessel 117, can be further treated at temperatures such as from 130 to 240°C in a reaction vessel of the type shown in the figure at 117. If if more saturated products are desired, a reflux column 119 can be used which can be adjusted to collect these products.

The bottom fractions 112a from the reaction product which is separated in the cyclones 116 will usually contain the more stable parts of the petroleum, or petroleum products. These products 112a, although it is assumed that they are partially converted in the reactor 110 (which can be characterized, among other things, by their improved API numbers), are best dealt with in a first parallel review. After these bottoms have reacted to some extent due to the catalyst 109 suspended in the boiling bed, further reaction will best be carried out with another specific catalyst suitably modified if necessary to a particular starting material .

The bottom fraction 112a is therefore fed to a reaction vessel 130 together with steam, and this mixture is then subjected to a catalytic treatment while it is in a hot state. The material 112a is easier to hydrogenate, and further improvements are therefore achieved by obtaining components with a lighter viscosity than those found in the heavier parts of the petroleum. The bottom fraction 112a can advantageously be treated with a different catalyst composition in the presence of steam in the same ratio with regard to the mixing ratio between steam and the starting material as in the first reactor 111. Different catalyst compositions have been found which are more advantageous in the subsequent reactions in relation to the catalyst which is used in the boiling bed reactor and which is designated 109. The use of another catalyst is due to a partial reaction of the more resistant components in vessel 111. The composition of the different catalysts for the treatment of different bottom fractions will be described in more detail below the following.

The reaction vessel 130 can be followed by another reaction vessel (not shown), so that the top products recovered in the reflux cooler 133 from the reaction vessel 130 can therefore be further processed. Usually, the products from the reflux column 133 will be washed in vessel 135 to remove unwanted constituents, e.g. H^S, after which the gaseous products are recovered and used. The reaction vessel 130 can be used in the same way as vessel 117, i.e. either as an upstream or a downstream reactor, with a supported or unsupported catalyst suspended in a bed or on trough 130a, or with fixed catalyst beds. A liquid catalyst, such as a hydrated melt of various alkali sulphides, and which is liquid at the operating temperatures used in said reactors, is described in US PS

4 366 044, and is very well suited in these reactors.

If no further reduction is desired

of the molecular size of the distillate product of the first reactor 110, then this should be hydrotreated in vapor phase reactors immediately after the product emerges from the cyclone 116 which separates the bottom fractions (the heavy partially reacted or unreacted components) from the top distillate which is the product. Cooling in cooler 116a prevents a re-condensation of the split components. Hydrotreatment is facilitated by an absence of the heavier fractions removed in cyclone 116. By definition, the hydrotreatment is carried out in the absence of either thermal cracking or catalytic cracking, and includes a hydrogen partition to a molecule. As the hydrogenation reaction is exothermic, the previously described cooler 116a will control the temperature and prevent the product from being damaged if temperatures are too high.

Fixed bed reactors can be used for this hydrotreatment, or fluidized catalyst reactors can be used. At present, however, it appears that liquid catalyst reactors are most efficient at temperatures below 280°C, which is due to the fact that the catalyst solidifies above 280°C due to a transformation of the hydrates.

The fixed bed reactors are less efficient with regard to the degree of hydrotreatment achieved in relation to

liquid catalyst reactors.

The starting material will usually be a so-called ORA fraction, i.e. which contains oil, resin and asphaltenes, and at atmospheric pressure it will boil at a temperature above 537.7°C. Although this is a relatively rough description because the amount

of the oils, resin and asphalt present is not precisely known, it is nevertheless an appropriate term for this most persistent component of an oil. It is known that asphaltenes can be solvent extracted from petroleum residues or from an ORA fraction, whereby the oil and resin can be further processed. Asphaltenes are known to be particularly amenable to further processing, e.g.

by catalytic conversion, and they thus form a fraction that can either simply be burned or coked using previously known methods to extract all available hydrogen.

An appropriate characterization of the asphalt is that it is the part of the asphalt that is soluble in carbon disulphide,

but insoluble in paraffins, e.g. ethane, paraffin oil or ether. Resins from ORA fractions can be extracted with propane. Asphalts are also soluble in carbon disulfide. Carbenes which are constituents of asphalt are insoluble in carbon tetrachloride, but soluble in carbon disulphide. Furthermore, the oily or soft components in an asphalt will also be referred to as asphaltenes. These can be dissolved in white spirit. Maltenes are pentane-soluble compounds, while asphaltenes are pentane-soluble compounds.

The above definition also includes natural asphalts such as petrol, mineral tar, earth tar, Trinidad tar, petroleum tar and original mixtures of hydrocarbons such as amorphous solid or semi-solid fractions produced by oxidation of residual oils.

As there is no fixed definition of these compounds, such as maltenes or asphaltenes, mixtures of these compounds are often designated as one or the other. Furthermore, the solvents and extraction and precipitation techniques used will to a greater or lesser extent affect the properties of the final product. Because of this, the solvent will extract asphaltenes, e.g. carbon disulfide dextra-cured asphaltenes that precipitate from heptane are not considered pure compounds, as they have no specific melting points but only softening points. The softening points for asphalt can be up to 204.4°C and higher. A convenient way to define the ORA fraction is that it is one that boils at 537.7°C

and higher, although this temperature limit is somewhat arbitrary,

and lower temperatures can also be used, e.g. 482.2°C, because it may be desirable that not all the starting material is distilled over. A lower temperature of 426.6°C therefore only characterizes a less intractable composition.

Thus, e.g. the asphalt extracted using a solvent and with a high softening point, have a softening temperature of approx. 132.2 °C, a specific density at 15.5 °C of

about 1.1149 and a viscosity of 4060 poise at 135°C. The specific gravity at 135° is 1.026, and the viscosity is 3.957 stokes or 395.700 cST (centistokes) (stokes is obtained by dividing the poise value by the specific density at the indicated temperature). The viscosity at 148.8°C for the same high softening point asphalt is 877 pois with a specific gravity of 1.016 at the same temperature, and 86,000 cST. At 162.7°C, this asphaltene has a viscosity of 261.5 poise and a specific density of 1,006, which gives a cST value of 26,000.

An analysis of the above-mentioned solvent-extracted asphaltene is set forth in the following table:

ASPHALT ANALYSIS

Calorie content in asphalt: 47,592.9 kcal/l.

Although the above-mentioned asphaltenes can be considered to be representative, other asphaltenes depending on their starting point may have other properties.

Based on the above product analysis, it appears that these products contain significant amounts of metals. The quantities will vary based on the starting material, and can go up to 6,000 parts per million (ppm) for vanadium, but is usually only up to approx. 600 ppm. Nickel and other constituents can too

be present in amounts of up to 600 ppm. These metals will consequently also affect the ability of the residues to be treated using normal methods for treating petroleum residues.

Based on various analyses, the hydrogen content in the ORA fraction will usually lie in the range from 13.5 to approx. 7% or lower per weight, but again this is not a precise characterization. In a petroleum residue that boiled above 537.7°C, the hydrogen content will be approx. 12.5% or lower. A significant percentage of "free" carbon can also be found (as Conradsen carbon), i.e. up to 45% by weight. The free carbon is defined as Conradsen carbon, or Ramsbottom carbon, but these analyses

are not identical because different methods are used to define the term "free carbon", which in reality does not necessarily have to be "free". In the ORA fraction, the amount of Conradsen carbon can be up to 40% by weight or more. In each case, such carbon residues can be converted using the present method.

Any cracking, whether it occurs by heat or by catalysis, will give smaller molecules than those that were present in the starting material. A smaller molecule will have a lower boiling point than a larger molecule. Hydrotreatment does not seem to cause cracking (and thereby the formation of smaller molecules). The boiling point temperature range appears to be about the same as before and after treatment with hydrogen.

Hydrofining is e.g. effective in the hydrogenation of olefins to paraffins, and in the elimination of diolefins and triolefins in the product by these being hydrogenated either to mono-olefins or paraffins. Aromatic compounds are also converted to naphthenes by hydrofining. Real paraffins do not seem to react during hydrofining.

Hydrofining can also reduce the sulfur and nitrogen content of the product. If, however, the asphaltenes, which contain most of the nitrogen, are not previously cracked, then the nitrogen in said asphaltenes will not be available for removal by hydrorefining. The degree of sulfur and nitrogen removal corresponds to the hydrofining of an unsaturated bond (where no hydrogen is present at its position, where the hydrogen is introduced into the molecule). The sulfur-containing and nitrogenous compounds that can be removed by hydrorefining are those where the sulfur and nitrogen atoms can be replaced with hydrogen under a thermodynamically favorable yield, during the formation of ammonia or hydrogen sulphide and are more saturated molecules.

The bromine number is a measure of the product's degree of unsaturation. The slightly unsaturated product fractions will have the highest bromine numbers. A hydrofining of the product will bring the bromine number down to an acceptable level, i.e. to a value between 40 and 50 for the fractions that first begin to boil, and up to those that boil at 204.4°C (obtained by distillation using the methods described above).

As mentioned earlier, the bottom products 31 can form a separate parallel stream, which can be used as a starting material, but they can also be recycled because a further reaction can lead to further complications. If the bottom products are nevertheless difficult to process, then they can be subjected to another treatment of the type shown for feed 12a, which is specifically indicated in Figure 2.

As most of the top products from the reaction vessel 11 are treated by direct distillation, the bottom fractions 12a which are collected in the cyclone 16 or in vessel 18 will be products, but to a lesser extent depending on the component composition of the petroleum. These bottom products can thus be processed in smaller vessels based on the quantity of the components in the petroleum products that are processed. Consequently, different mixtures of petroleum can now be treated depending on the component in the bottom product 12a. The size of the reactor 30 can vary in relation to the bottom component of the petroleum, i.e. those that boil at a temperature of 537.7°C and higher.

Steam is introduced into the reactors 110 and 130 in quantities which depend to a certain extent on the products one wishes to produce. The amount of steam can be from 50 and up to 100 to 130% of the recovered product per weight. Steam can also be supplied in smaller quantities, i.e. so that you do not get significant coking due to an absence of steam, but which would occur in the reactors 110 and 130 if there was not sufficient steam present. In smaller quantities, steam can thus be supplied so that substantial coking is avoided in the reactors 110, 117 and 130 etc. Additional steam can e.g. is supplied to the reactor 117, 130 etc. as required. The steam that has a low pressure will usually be excess steam.

The temperature in the reactor will usually be up to approx. 650°C, although the most advantageous operating temperature is below 450°C, e.g. below 425°C. The temperature in the subsequent reaction vessels, e.g. 117 etc, can, however, be lower, and this seems to give a more hydrogenated product, which is characterized by a lower bromine number. The more resistant components in reactor 130 may, however, require the same temperature as in reactor 110.

The cooling can be carried out at a temperature of 390°C or lower. A third stage reactor not shown in the drawing can have a temperature as low as 170°C, or lower, i.e. down to 125°C (but without steam condensation). Lower temperatures are desirable because the hydrogenation is exothermic and most of the hydrogenation is carried out in downstream reactors, while most of the cracking is carried out in reactor 110 and to a lesser extent in reactor 130.

The pressure in vessel 111 and in the subsequent reactors can hold

at atmospheric pressure. Although superatmospheric pressure is possible,

then this is not appropriate. If the pressure is between 7.1 and up to 10.7 kg per cm2 (although you can use up to 25 atmospheres), it seems that you get an improved reaction. The most preferred ratio is between 4.28 and 10.7

kg per cm2 when using elevated pressure. Steam condensation must be avoided at the higher pressures. Although the reason for the conversion is not fully understood, it seems that better hydrogenation is achieved in the aforementioned pressure range, while a pressure beyond this range leads to increased costs in terms of equipment and capital investment.

Instead of a cyclone separation as shown in the drawing, i.e. that you use one or more cyclones 116, then you can use other separation devices to recover the liquid and/or entrained products, and thereby carry out a parallel treatment of the petroleum components, and such other devices can be centrifuge separators, blowout drums and the like.

With regard to the materials that can be subjected to a treatment, these can be different crude oils of different origin and viscosity. Such crude oils, as long as they have sufficient fluidity at the processing temperature, can be processed and will give the results described here. Residues and bottom fractions obtained at atmospheric pressure conditions can be produced by ordinary petroleum refining, can also be used. All of these starting materials are characterized by having a fraction that boils at a temperature lower than 454.4°C. However, the parts and reaction products (reactor 110) which have a boiling point above 454.4°C can best be treated in a parallel flow as described here.

A further treatment of said secondary stream will

give products with improved viscosity, improved hydrogen content and improved chemical structure (i.e. products with a higher paraffinic content and with higher K values, or aromatic polycyclic compounds have been produced which can subsequently be processed to give desirable less olefinically unsaturated cyclic compounds, as well as linear or pure paraffinic products). The gaseous products which are produced by means of the present method, besides the conversion of unsaturated products into more saturated products, further contribute to an improved result by means of the present method.

With hydrofining of the product according to the present invention method to thereby reduce the bromine number and further reduce the nitrogen and sulfur content in said products, it is necessary to remember that hydrofining at lower temperatures does not seem to "crack" the product to a particularly significant degree, and thereby form less molecules, which in turn would reduce the boiling point range. Hydrofining of the product will require further processing, e.g. by cracking (as determined by the lower temperature range of the boiling point, which is due to the formation of smaller molecules), more difficult due to the fact that the process is apparently unable to crack paraffins.

Hydrofining should therefore be the final part of the present method and should not be used until the desired molecular size has been achieved by means of the "cracking" according to the present method. Thus, vacuum residues can first be cracked into smaller molecules under less stressful conditions, after which the more resistant esters, i.e. stream 112a, can be cracked under suitable stressful conditions that do not affect the main products. The resulting product has a reduced temperature range with respect to boiling point. A molecular size reduction in this way also leads to a reduction of sulfur and nitrogen content in the top product compared to what is found in the starting material.

The resulting smaller molecules that make up the "cracked" end products do not contain the high amount of nitrogen found in the starting material. The nitrogen will normally be concentrated in the heavy asphalt parts of the petroleum residues. Metals are bound through the nitrogen bond to the methylene groups in the porphyrin structure in the asphaltenes, and they will thus be able to be separated by said parallel treatment. Metals and nitrogen bound to the central porphyrin structure in the asphaltenes are almost absent

only present in small amounts of the top distillate product,

which is derived from vacuum residues containing asphaltenes,

and they have become relatively concentrated to the asphaltene part of the crude oil which formed said vacuum residues. The metals and nitrogens are more concentrated in the No. 1 bottom fractions, i.e. 112a, which do not form part of the top distillate product. The sulfur content will not necessarily be concentrated in the asphalt part of the petroleum residues. The resin portion of the residues may contain as much as 5% by weight of sulphur.

With regard to the catalysts which are sometimes called reagents, these are used in such a form that they give the desired results, i.e. sufficient cleavage of the bonds in the first and further parallel flow reactors, and in the final hydrogenation. The results are achieved by loading the starting material appropriately, but not too strongly, with the catalysts described here,

and these will be described with reference to the results obtained. Furthermore, the pre-treated more resistant components will be suitably loaded in a parallel current. Mixtures of catalysts may also be highly desirable for crude oils and their residues, or for mixtures of crude oil based on different crude oil components.

A suitable suitable loading catalyst for reactor 110

is designated catalyst A, and can be prepared in the following way. A mole of potassium hydroxide is dissolved either in ethanol or methanol, preferably in ethanol, or in an ethanol and methanol mixture, less advantageous due to poorer solubility, in 1-propanol or 1-butanol. The solubility of the catalyst product is lower in the two latter alcohols, and one must therefore use larger amounts of the alkanol if these two types of alcohol are used. The alkanol can be an absolute alkanol, although 95% ethanol can also be used.

The potassium hydroxide dissolves in this solution and then reacts with the hydrogen sulphide which is bubbled through the solution. After saturation, the catalyst - alcohol mixture will be recovered and the alcohol separated by vacuum distillation. The residue will be the catalyst. When using a mixture of the solvent, typically 1 mol of potassium hydroxide will be dissolved in 200 cm3 of ethanol and 130 cm3 of methanol. Typically, one will use pellets of potassium hydroxide of analysis quality (approx. 86% KOH), absolute ethanol or 95% ethanol and absolute methanol. As mentioned above, the quantity ratio in the ethanol and methanol mixture can be varied. The solution is evaporated in vacuo until no remaining alkanol can be removed. These catalysts seem to be best suited for decomposition, especially decomposition of the more resistant components, and they can therefore be used in appropriate amounts, possibly together with other catalysts, to achieve the desired degree of decomposition.

The catalyst can be unsupported in the first stage reaction, but it is most advantageous both for the first and subsequent stages that the catalyst is deposited on a substrate, a so-called net.

With regard to catalyst substrates, these will preferably be used to increase the surface area. Spinels, e.g. chromite spinel (CrO), and most advantageously porous metal, e.g. stainless steel of available AISI grades and the like. The latter is achieved by sintering very finely sieved uniform, pulverized, metallurgical particles, or they can be produced as thin plates obtained by leaching the leachable components into thin (i.e. 3 mm thick) plates, whereby small intercommunicating passages are provided. A metal substrate can e.g. produced by weaving very fine nets, i.e. 0.2 to 5 mm thick, and cut these into lengths of 2 to 5 mm. Other substrates that can be used are aluminum oxide with pore sizes varying from 50 to 350 Angstroms or even up to 1000 Angstroms, but these must be protected as explained below. Although the treatment in the subsequent reactors may be less demanding with regard to the properties of the substrate, the treatment in the first stage reactor in accordance with the present invention will preferably be carried out with a catalyst placed on a strong inert substrate, such as a porous metal substrate which can have a pore size of up to 3500 Angstroms or greater, i.e. the volume of the substrate can be made up of from 10% metal and 90% pores, although the metal can also make up up to 25% by volume of the substrate.

In all reactions, the catalyst will react in the absence of oxygen. In a similar way, the catalyst must be deposited on the substrate in the absence of oxygen, which is also the case when the volatile compounds are to be driven out of the substrate.

After the volatile compounds have been driven off, the catalyst and substrate will be heated to a suitable temperature, between 320 and 450°C or even up to 560°C. The catalyst will be strongly attached to the substrate and can thus be used in a so-called spinning arc reactor, a boiling bed reactor or a fluidized bed reactor.

If the substrate is too strongly attacked by the catalyst, which can happen with aluminum oxide in the first stage reactor, then the following procedure is used. The above-mentioned catalyst is evaporated to considerable dryness, dissolved in glycerol, after which the glycerol - catalyst mixture is deposited on the aluminum oxide substrate. Other less resistant substrates that can be attacked by the catalyst, e.g. such as molecular sieves, are treated in a similar manner. Such molecular sieves will usually be of Y and X, e.g. YL-82 type, low sodium (available from Union Carbide, Danbury Ct or a comparable substrate from Mobil Oil, New York, NY). However, the molecular sieves function in this case as a substrate for the catalyst, i.e. to increase the catalyst's contact area.

After deposition on the substrate, the glycerol - catalyst mixture will be progressively heated to approx. 560 °C, whereby the volatile compounds are driven off.

Glycerol can therefore first be deposited, heated to approx. 200°C, after which the catalyst is deposited on the substrate with glycerol (after the substrate has cooled), after which the product is heated to the desired temperature.

The reaction in the first reactor can take place at a high temperature, i.e. from approx. 320 to approx. 450°C, although temperatures of up to 560°C, and even up to 650°C have been used. For the asphaltenes, the preferred temperature range will be on

about. 360 to 430°C, and it seems that the temperature range from 390 to 425°C is a very good operating range.

For these catalysts, the reaction must always be carried out in the presence of steam to facilitate hydrogenation, hydrocracking etc. The steam is used in a minimum ratio of approx. 27% by weight based on the weight of the starting material, whether this is an ORA fraction, a crude oil, petroleum residues etc. Conversely the amount of water added in the form of steam at the operating temperature must be increased or decreased based on the desired degree of hydrogenation (which also takes place in the first reactor to some extent). If more hydrogenation is desired, more steam must be added, but usually the amount of steam should not exceed 85% by weight in relation to the starting material, although 130% by weight can be used based on the hydrogenated product that is removed (if the gaseous fraction produced is used for a liquid equivalent). Put another way, the amount of water used must be determined by subtracting the hydrogen content of the starting material from the hydrogen content of the desired product on a weight basis and multiplying the amount by 9 (as water makes up 1/9 of the hydrogen weight). Typically, up to 30% excess will be injected into the first reactor.

If water is not supplied to the reactor in the form of steam, or is interrupted for one or more reasons, coke formation will easily occur. Carbon will thus be deposited or developed by a method corresponding to catalytic thermal cracking, and in this case the catalyst will act as a cracking catalyst, albeit with a certain advantage in a different way (because this catalytic thermal cracking takes place at a relatively low temperature, i.e. approx. 320°C), but far less efficiently than when it functions in the presence of steam as a hydrogenation catalyst and/or as a cleavage catalyst. As mentioned earlier, some coke formation can occur, but this occurs to a very small extent.

Interruption or insufficient steam supply will also give

a production of particularly heavy products in the reactor 110. It is therefore important that steam is supplied continuously in a suitable manner to the reactor and that the steam is evenly distributed (i.e. that there is no space that is free of steam and/or reagent) .

Nevertheless, it should be mentioned that too large amounts of steam will also prevent the reaction from proceeding in the desired direction, which is presumably due to too strong entrainment of partially converted products.

If carbon is deposited for one reason or another, this will typically be on the hot spots, such as the heated reactor walls or on the catalyst substrate. The reactor 10 is therefore preferably operated adiabatically. Carbon deposits on the catalyst can be operated by, i.e. can be converted into usable products by expanding against steam for a sufficient period of time without adding additional starting material, and then the catalyst can again be used for the production of desired products. An occasional addition of hydrogen sulphide or sulfur will often be beneficial, especially for starting materials with a low sulphide content, i.e. for non-corroding crude oils.

Furthermore, it has been found that if the temperature, e.g. in connection with the above-mentioned catalyst, is increased to between 320 and 420°C, depending on the composition of the starting material (when, for example, this contains a significant part of asphaltenes) then an exothermic reaction occurs at approx. 440°C. This reaction can raise the temperature to up to 600°C, but this will also depend on the amount of steam supplied. More steam will in these cases tend to give lighter carbon products. Too high temperatures are not desirable, and it is preferred that the temperature stays below 440°C.

In the second stage reactor, which is shown as no. 17 in Figure 1 or as 117 in Figure 2, and where further reactions take place, the products from the first reactor should advantageously be able to quickly cool down to approx. 250°C in the presence of a catalyst, usually to temperatures of about 250 to 390°C. The quality of the product will increase when the process is operated at a temperature of up to 430°C, preferably 425°C,

but the conversion will apparently not increase after the temperature has passed 390°C. The cooling in cooler 16a is carried out in such a way that the steam does not condense and affect the reaction. The lighter fractions, i.e. hydrogenated products will not condense.

The catalyst used in another reactor 17 or 117

The ig is preferably a supported catalyst, but can also be unsupported, and is used for the treatment of vaporous or gaseous products. A typical catalyst for second stage reaction, i.e. reactor 17 or 117, is a less burdensome catalyst B. This catalyst can be prepared by dissolving potassium hydroxide either of technical or analytical quality, i.e. up to 86% potassium hydroxide, in absolute or 95% ethanol or methanol (preferably ethanol) and saturated with hydrogen sulphide, but without boiling off the alkanol, after which one collects and condenses any alkanol that evaporates in a subsequent vessel. Other vessels where a reaction occurs may be downstream to collect the hydrogen sulphide. When the last vessel containing KOH shows a reaction, the reaction in all upstream vessels is stopped.

If the reaction is continued in other reactors, e.g. 17 or 117, then the advantages of the present method will lie in the combination achieved by an immediate cooling of the reaction products from the first stage reactors to approx. 300°C, preferably 250°C, in the presence of a catalyst, after which these reactions are carried out on the products of the first stage in this second stage. For this purpose, it has been found particularly advantageous to leave the catalyst placed on a suitable substrate. These substrates can be the same as for the catalyst in the first stage,

but in each case the substrates must be inert under special reaction conditions prevailing in the reactor, e.g. reactor 17 or 117, or 130. These second stage reactors 17 or 117 can be fluidized bed reactors (circulating, fluidizing bed reactors, or partially circulating or statically fluidized bed reactors), fixed bed reactors or fluidized bed reactors. Reactor 30 can also be a boiling bed reactor or a fluidized bed reactor or a circulating fluidized bed reactor.

It has been found acceptable that in second-stage reactors, i.e. 17 or 117 use substrates of the type that are usually available, e.g. aluminum oxide - aluminum oxide silicates of a solid zeolite type, e.g. a molecular sieve, where the sodium or potassium in the zeolite is replaced with ammonia. E.g. zeolites of type X or Y (10 and 13) can be suitably used. It is preferred to use molecular sieves of zeolite type Y, and of these it is particularly advantageous to use molecular sieves with a low sodium ratio (i.e. approx. less than 1.0% Na 2 O). The molar ratio between silicon dioxide and aluminum oxide in these zeolites is greater than 3 to 1, i.e. about. 5 to 1, and the Na20 content is approx. 0.2% by weight. Such zeolites are available from several commercial sources, either in the form of powdery spheres, cylindrical extrudates or of other suitable sizes, e.g. extrudates or spheres with a diameter of approx. 3 mm. Although it has been suggested that such substrates can be poisoned or destroyed by alkali metals, as they are processed by the subsequent process,

then it has been shown that these substrates are none the less advantageous for a deposit of the alkali-sulphide reagents described here. Such substrates can also be used in the first stage reactor 10.

Other zeolites are the potassium-type ELZ-L zeolite described in US PS 3,216,789, and the silicate compounds described in US PS 4,061,724. The latter has a pore dimension of approx. 6 Angstroms. Other substrates can e.g. be those described in GB PS 1 178 186, i.e. with a very low sodium content, i.e. usually less than 7 vol%, e.g. ELZ- -6, or ELZ-E-6, E-8 or E-10. Other substrates that can be used are mordenite and erionite with a very low sodium content, which are produced by ammonia extraction and of the calcined type. Of the above-mentioned molecular sieves, type Y with a very low sodium content, i.e. 0.15% by weight, and the yield with ammonia, be available under the trademark LY-Y82 from Linde Division, Union Carbide Corporation, Danbury, CT, Mobil Oil Corporation, New York, NY, and other commercial firms. In each case, the stability and durability of these molecular sieves as a substrate must be tested under the reaction conditions, and it should be investigated how the catalyst functions in the second stage reactor.

The procedure for producing the substrate to be used in the second stage reactor is as follows. Zeolite extrudates with low sodium content and ammonium yield, e.g.

in the form of powders or cylinders, stars, rings, spheres etc., or extrudates with a size of approx. 3 mm to 6 mm or the like, is treated with glycerol or similar polyhydroxyalkane compounds, e.g. partially reacted polyhydroxy compounds such as hexahydroalkanes, by first impregnating the said products in a closed reactor. Then, e.g. if you use glycerol, you heat the product and remove decomposition products from the powders or the product, at temperatures from room temperature up to 265 or 280°C, or even up to 560°C,

where a reaction occurs. The reacted substrate is then sieved, sieved and cooled in a closed container if the temperature has been brought up to 5 6 0°C.

When the above-described substrate has cooled, it is impregnated with a catalyst of the general form K^S^

(empirical), and this catalyst is acceptable but not particularly good for cracking. Although the catalyst is designated by an empirical formula, the particular method used for its manufacture will determine the catalyst's function. Consequently, one can prepare different catalysts with the same empirical formula, but with greatly different properties and different starting materials treated will give different products. This catalyst, designated as catalyst C, is prepared by dissolving 6 mol of KOH in 4.5 to 7.5 mol of water without

external heating is used, after which a small amount of an alkanol is added, e.g. 2 to 2.5 cm 3 of methanol or ethanol per mol of KOH. 4 mol of elemental sulfur is then added to the aforementioned solution, which will then react exothermically. An appropriate amount of sulfur is then added to adjust the catalyst to the desired sulfur level, which occurs by adding additional sulfur so that an empirical sulfide is formed, i.e. from<I>^<S>,K^ S ^ ^ to K2S2 5 '°9nerved includes K2S5' depending on the desired product composition. In order to

to get more gas in the product, less sulfur-saturated compounds are used. If you want more liquid components in the product, you use more sulfur-saturated compounds.

Another catalyst, i.e. catalyst D, is prepared

in the following manner. One mol of KOH is dissolved in 1.0 mol of water under vigorous stirring. 2 ml of methanol or ethanol are then added immediately after the calyx has dissolved. Then 2/3 mol of elemental sulfur is added and a forceful reaction then occurs. The catalyst is adjusted to the desired empirical sulfur content by adding appropriate amounts of sulfur with further stirring, i.e. 1/4 to 2/3 mol of sulfur relative to the empirical sulfur content of K2S, i.e. that from 1/4 to 2/3 mol of dissolved sulfur will give K^S1 5, while 1/2 to 2/3 mol will give K2S2 Q etc, and here other suitable fractions are included. The catalyst can thus vary from KnS„ . to K_S_ c or even up to K»SC.

z 1 , 12.2. , D ^5 Catalysts with a low sulfur content, e.g. K2S^ are good cracking catalysts when mixed with catalyst A, and are a good catalyst in general for less burdensome treatments of crude oil. This catalyst is also a hydrogenation catalyst.. Catalyst A will usually be mixed with from 3 to 25%, but typically less than 10% by weight of catalyst D.

Once the catalyst is prepared, it is vacuum evaporated into a flowing suspension. It is then poured over the glycerol-treated and cooled extrudate as described above (i.e. if the substrate has been heated to 300°C or higher) under very low vacuum, then stirred and dried. The catalyst is then sifted when dry,

and then fed into the second-stage reactor, which is cleaned of atmospheric oxygen, when it is to be used as a hydrogenation catalyst.

Another method of protecting the substrate, if a glycerol-treated substrate is heated to between 260°C

and a decomposition point (which indicates a significant reduction in the liquid condensate that is collected)

is of the catalyst suspension described above, is added and the vessel is covered and heated to at least 440°C, i.e. up to 560°C. In each case, the catalyst will be calcined above the temperature at which it is used.

Another method is to mix glycerol, i.e. up to 88 ml of glycerol to approx. 1 mol of a catalyst in solution, or by mixing the above-mentioned catalyst mixtures with glycerol. The mixture of glycerol and catalysts is heated to drive off the water and/or alcohol, whereby a glycerol solution of the catalyst is obtained. The temperature is then driven up to 190°C. The mixture is then poured over the substrate and, while stirring, it is heated to at least 450°C, or even up to 560°C. The supported catalyst will then have a very unpleasant smell. It must be manufactured under closed, well-insulated conditions.

For use in a first stage reactor of approx. 4 liters together with a second-stage reactor, you can add approx. 2/3 mol of the supported catalyst (K base) to the second reactor. Thus, e.g. an alumina-borne (empirical)

catalyst is added to the second stage reactor.

Another catalyst, i.e. catalyst E, is an unsupported or supported catalyst capable of reducing the molecular size of the product from the first stage reactor (or when used in the second stage reaction. Catalyst E can be prepared by adding a dried KHS powder

or suspension in appropriate molar portions (based on the desired molecular size of the product) to each of the above-described reagent mixtures, i.e. A, B, C or D. You can either use supported or unsupported forms of the catalyst. I.e. that one adds from 1.5 to 1.3 mol on a molar basis of K of said KHS to e.g. K2S (empirical) sulfide or K„S1 (empirical), and the molecular size of the product will decrease under these additions of KHS.

When the procedure is carried out with these supported catalysts in the second stage reactor, appropriate adjustments should be made, e.g. with K^ ^ller ^S^^ 5 to get more hydrogenation, while K2S2 gives larger molecules (also more distillate on less gas ). These reactions are carried out in temperature ranges from 113 to 440°C. Similar catalyst adjustments can be made in other reactors, i.e. when using more than one second stage reactor 17 or 117.

In each case, the first step reaction will be carried out with a more closely specified reagent in order to thereby achieve the desired degree of cleavage of the more resistant and irreducible components, i.e. residues of a crude oil, an ORA fraction and especially asphaltenes. In a parallel treatment, the unreacted irreducible compounds will then be treated catalytically to achieve a further degree of cleavage. The total combination in the process in further stages, i.e. possibly a second, third, fourth etc. step, will depend on the reaction of the first step and how large the quantities are

unreacted components are removed.

The amount of catalyst that is deposited on the substrate is

from 4M of the catalyst (K basis) to approx. 0.5M or can be as low as 0.1M (K basis) per 500 cm 3 of the substrate.

Another basis that is not completely identical is that the amount of catalyst is approx. 20 g per 300 cm3 substrate, but can vary from 3 or 5 g reagent per 100 cm3 substrate to approx. 25 to 30 g per 100 cm3 substrate.

The vanadium metal can be completely removed from the petroleum starting material, and the heaviest product can contain substantially all of the vanadium. If the reaction is carried out with the available catalysts, then vanadium will only be detectable in very small amounts or not at all in the bottom fraction from reactor 2, i.e. product 18. The removal of nickel is facilitated if some of the reagent is present in the hydrosulphide form. No reaction occurs between alkali metal sulfides and nickel sulfides, but there is a solubility reaction in the presence of alkali metal hydrosulfides and nickel sulfides. Nickel (and iron) forms complexes such as ferrites with alkali metal sulphides - hydrosulphides. These complexes are immediately hydrolysed in liquid water and will form a precipitate of iron or nickel hydroxides. In liquid water, the vanadium complex with the catalyst will be easily soluble and stable. Iron will normally be present in the residues after distillation determinations, in amounts of between 3 and 5 ppm, but the amount will of course depend on the amounts in the original starting materials.

Typically, the catalysts for the second and further stage reaction will be hydrosulfides and sulfides, i.e. monosulfides and polysulfides of the elements of Group 1A of the periodic table other than hydrogen, and they can be prepared from an alkanol solution as described above.

Although for the purpose stated above you can use

sodium, potassium, rubidium and lysium, then it will usually be most beneficial to use sodium and potassium. Of these two, potassium is preferable. Although rubidium compounds are acceptable, rubidium with lysium would not be advantageous to use due to cost. In the first stage reactor, however, rubidium can be very advantageous in a mixture of rubidium, potassium and sodium in the following proportions: 14% rubidium, 26% potassium and 60% sodium sulphides, based on the elemental metal per weight. The ratio range for the aforementioned mixture is 1:1.5-2.5:3.5-4.5, respectively, but these compositions must be prepared in the same way as indicated for catalyst A. The catalysts used are usually in the form of hydrates, but a smaller amount of the catalyst is apparently in the form of an alkanolate (hydrate analogue) i.e. up to 15%, but usually less than 10% or even less than 5%. Alkali metal thionates are also present as intermediates. The hydrate (and alkanolates) of these compounds are very complex and undergo a number of transitions during the initiation of the reaction. It appears that a presence of mixed hydrates and alkanolates is necessary to obtain the best results. No attempt has been made to investigate how these transitions occur, or how the real structure is under the stated reaction conditions for sulphides, hydrates, alkanolates, thionates or mixtures of these compounds. However, it is sufficient to be able to state that the added catalyst can be a mixture of a number of hydrates and/or alkanolates, or an euthetic mixture of different hydrates and/or alkanolates, but in each case the specified cleavage catalyst must be used on the more resistant connections.

In a similar way, during the reaction there will be an internal transformation of the sulphur-containing compounds of the sulphides, and no attempt has been made to describe this internal transformation. components, i.e. asphaltenes of different molecular size, resins and oils such as suspended oils, i.e. more than 50% by weight, and this may include up to 95% of a fraction that boils above 537.7°C,

then it is necessary that in first stage reactors you have at least a part of the specific catalyst A

as defined above. As the percentage of the more stable compounds in the starting material decreases, the less catalyst A needs to be used. Also for the feed material 12a, it is recommended to

use catalyst A. This catalyst can also be added to other catalyst compositions depending on the desired bonds that you want to split (as this can be characterized by the product's API number, viscosity, etc.), without too much gas production. For No. 1 bottom fraction, i.e. stream 12a, then the preferred reaction vessel is a boiling bed reactor, or a stirred cage or curve reactor, although a fluidizing bed reactor or a fixed bed reactor can be used. Another catalyst for feed 12a is catalyst B with 5% by weight (on an elemental K basis) of KHS or replaced in the same amounts with up to 95% of catalyst A. The order of the catalysis for bond cleavage of more resistant compounds is as follows : Catalyst A (ethanol-made compounds); catalyst

A (methanol-made compounds); Catalyst A (ethanol-methanol prepared compounds); Catalyst A + B (mixture of 50% or more of Catalyst A on a K basis mole); Catalyst D; catalyst E (about 1/3 mol added KHS); catalyst B, and finally catalyst C. Of course, increasing amounts of KHS to the catalysts will increase their cracking ability.

The order of catalysis for hydrogenation is as follows: % catalyst A + KOH on a molar basis, then added the other half of catalyst A (highly exothermic hydrogenation reactions); catalyst C with KHS mixture (based on the required degree of cleavage); catalyst

B + A (50% or more of catalyst B on a K mol basis), and catalyst B.

It should be emphasized that in the above discussions and descriptions the various carbonaceous starting materials and their components will be termed permanent on the basis of what has been common in the past. As a comparison, it can be stated that the present method is carried out at lower temperatures and lower pressures, and is capable of converting all previously known stable compounds with great advantage and with great ease. However, since carbonaceous starting materials and their components have different "permanent" natures according to the present process, i.e. that paraffins are the most resistant, while asphaltenes and aromatic compounds are not particularly resistant, the present invention is based on the fact that the present method can treat all components of these carbonaceous starting materials in a suitable order with the help of a suitable catalyst, and that the method is characterized by obtaining yields which were previously unattainable when treating these carbonaceous starting materials.

In the following examples, different reactions are described. The invention is not limited to the specified examples which are only given to illustrate the applicability of the invention.

Example 1

An asphaltene with a high softening point of 132.2°C as described below, and of a solvent-extracted type, was treated with the following reagent to produce product A. The catalyst was catalyst A as described earlier. When the product from the first stage treated asphaltenes was reacted in the second stage, the product was identified as B. The second stage catalyst was the same as in the first stage. None of the catalysts were placed on an inert substrate.

132.2°C - softening point

§ Aromatic compounds and olefins were not clearly separated in the column, which is due to a heavy fraction that was distilled above 315.5°C.

The above data clearly indicate a significant improvement with regard to the density and viscosity of the products, a dramatic increase in the hydrogen content and a significant removal of the metals present from the later fractions.

The following examples show the results obtained. The starting material was solvent-extracted asphaltenes of the type defined above and in Example 1.

All experiments were carried out on a production basis in a stirred tank reactor. The reactor had a diameter of 15.8 cm and a height of 25 cm. It was equipped with a stirrer and a steam shower.

A steam generator, which was connected to ordinary drinking water, produced steam at 2.58 kg/cm2. The steam was passed through a pipe with an inner diameter of 9.5 mm to said shower and was supplied at atmospheric pressure. However, the reactor can use pressures from 2 atmospheres to 5 atmospheres or even higher, as needed.

Even the shower head had a diameter of 8.9 cm and had a series of holes, which directed the steam upwards. The shower was located at the bottom of the reactor.

A stirrer was used in the reactor when catalysts without an inert substrate were used. The engine was located directly above the reactor. Above the reactor there was a gasket through which the stirring axis passed. The stirrer was of a double type, and fitted with bent blades. However, the stirrer can be replaced with baskets where a supported catalyst has been placed as described earlier. In this case, 4 baskets containing the supported catalyst were placed on the stirring axis itself. The baskets and stirring axis had a total diameter of approx. 15.8 cm. The baskets were 15 cm high and had a thickness of 1.25 cm. The unassembled basket is a rectangle with a deep deposit 1.25 cm.

Through the above packing on top of the reactor goes

there also a riser that leads out distillates from the reactor,

a pressure equalization tube which includes a valve that opens at a pressure of 2.15 kg per cm2, and a withdrawal tube for withdrawal of the reactor's contents. The above-mentioned pressure equalization tubes can also be used to fill the reactor with solid starting material.

Usually two thermocouples will be placed on top of the reactor. One thermodome measures the temperature in the lower part of the reactor, while the upper thermodome measures the temperature in the upper half of the reactor.

The riser consists of a wire that is approx. 22.5 cm high and

has an inner diameter of 18 mm.

The second stage reactor is a tube with an inner diameter of 3.25 cm and a length of 30 cm. The reactor's capacity is 347.5 cm3. This reactor is equipped with three heating elements. A thermodome controls each of the heating elements through control devices placed on a portable device.

Gases and vapors go up through the second stage reactor and are then led down through a 40 cm long glass bubble cooler. This cooler is not water cooled.

The first cooler is placed vertically and the bottom of the cooler is connected to a 500 cm3 collecting flask. The flask is normally kept at 240°C using a heating jacket. The bottom of the flask has a glass tap for collecting product.

Another cooler is also attached to the above-mentioned flask and is parallel to the first cooler. The second cooler is not water-cooled. The second cooler is also a glass cooler with bubble-type cooling areas.

The downwardly directed pipe from the second cooler is connected to a water-cooled cooler. The water-cooled cooler is placed vertically and approx. 45 cm long, and is placed on

top of an unheated -liter flask fitted with a glass stopcock at the bottom. A parallel vertical water-cooled cooler is also located on this flask. Another water-cooled cooler is located directly above this cooler.

The water-cooled cooler on the top is equipped with a 30 cm long cable that runs diagonally upwards. This wire has an inner diameter between 12 and 18 mm. This line is connected to the ice cooler.

The ice cooler is a double-walled vessel of the type usually used to collect vapors so that these are not fed into a vacuum pump. The central part of the vessel is filled with a water - ice mixture, while gases and vapors are passed through the outer part. The gases and vapors are fed into the bottom of the vessel and exit at the top. The bottom of the vessel contains a collection vessel of 50 cm3. This collection vessel is also equipped with a glass tap for removing product.

The remaining gases and vapors are sent to another cooler of the same type as the ice cooler. This cooler is cooled using a mixture of solid carbon dioxide and 2-propanol. The product is again collected in a 50 cm3<1>s vessel under the cooler, and the collection vessel is again equipped with a glass tap.

The remaining gases and vapors are then washed in a solution of potassium hydroxide. This solution contains 6 mol of KOH dissolved in 360 ml of water. The gases are then measured by passing them through a gas meter. After this measurement, the samples are periodically collected and collected gases are released into the atmosphere.

In the following experiments, 1300 g of solid asphaltenes were weighed out

and crushed to a size so that they could be added to the reactor. Liquid or solid catalyst protected from oxygen,

was then added to the reactor which had previously been purged with helium. About. 40 g of theoretical anhydrous catalyst A was then added to the reactor.

The second reactor was filled under the same precautionary rules, usually with approx. 300 cm3 of a supported catalyst.

The second reactor was heated to drive out the water in it

both the zeolite and the catalyst.

After the water content in the second reactor had been reduced at the temperature raised to approx. 300°C, then the primary reactor was heated.

Solvent-extracted asphaltenes have a melting point between 93.3 and 204.4°C. The melting point determines the particular form of asphaltenes that can be used.

After sufficient temperature had been reached for the asphaltenes to melt in the primary reactor, the stirrer was turned on. Normally, the stirrer was initially driven at 30 to 60 revolutions per minute.

Steam was normally supplied when the temperature in the first reactor reached approx. 220°C. By this time, the temperature in the second-stage reactor had risen to approx. 424°C.

Helium was normally passed through a shower or spreader head before adding steam to the system to keep the spreader holes clear and the system free of oxygen. The helium was supplied in a quantity of 220 cm3 per minute.

Example 2

1,300 gr of solvent-extracted asphaltenes were reduced to a size so that the material could be fed. into through the opening at the top of the reactor, which had a diameter of

about. 2.5 cm. The asphaltenes were not heated, but added to the reactor in solid form. After the addition, the reactor was purged with helium.

The catalyst was another version of catalyst A which was prepared as follows. The previously described solution of KOH was added to a solution of 1 mol of KHO dissolved in 30 cm3 of water, after which the solution was saturated with hydrogen sulphide. The solution separated into two layers, approx. 1/3's top layer and a 2/3's bottom layer. The layers were separated and dried after which the two parts were again mixed. The mixture can be carried out in the same quantity ratio as when the layers were prepared (which was the case in this example) or the quantity of the two catalysts can be varied. The catalyst can also be dissolved or dispersed by mixing again for deposition on a substrate.

In a theoretical anhydrous state, the catalyst weighed approx. 40 gr.

A zeolite-supported catalyst was added to the second-stage reactor during the assembly of the apparatus. This reactor contained approx. 300 gr substrate and the catalyst. The zeolite was an L2-Y82 zeolite, while the catalyst was catalyst D, i.e.<K>2<S>15(emPirisk) to improve the hydrogenation of the cleaved product. At start-up, it is necessary for the second reactor to hold at least 175°C before starting the primary reactor.

The first stage reactor was then heated. Only the bottom part of the reactor was heated, while the upper top was not heated. At 220°C, a small amount of steam was added to the reactor through the bottom spreader head.

At approx. At 320°C, a slow production of hydrocarbon product was obtained at the bottom of the first stage reactor, which condensed in the flask below the two water-cooled coolers. However, this product was much heavier than the product obtained at the process temperatures, i.e. in the range from 390 to 424°C.

When the temperature in the first stage reactor had reached 360°C, there was a significant improvement in terms of production speed. The reaction became exothermic and rose rapidly to 415°C, and remained at this level. This temperature was then maintained at the bottom of the reactor. The top of the reactor had then reached 360°C.

The temperatures in the second stage reactors were then in the range from

220 to 460°C.

When the contents of the first stage reactor came into contact with the stirrer, the process occurred at approx. 415°C in the first stage reactor and between 440 and 460°C in the second stage reactor.

The amount of steam was calculated to be approx. 20 cm3 of water converted to steam per minute. At the end of the experiment, the temperature at the top and bottom of the primary reactor was allowed to rise to 440°C.

The Catalyst A variation above (as described) gave almost no gas through the gas gauge. The amount of gas was less than 6 litres.

The main quantity of the product was No. 2 bottom fractions which were collected under the water-cooled coolers. This product was combined with the product collected under the water - ice trap, and a total of 458 gr was obtained. This product had an API number of 23 (specific gravity at 15°C = 0.9158).

No. 1 the bottom fractions totaled 33 gr and had a specific heat of 0.96587 (AFI number 15) at 15.5°C. No. 1 the bottom fraction was collected under the air-cooled coolers.

The amount of bottom fractions collected under the dry ice-2-propanol trap measured 44 cm3 in the calibrated trap. However, when it was collected, only 28 cm3 was obtained due to evaporation of the light products. The API density of these light products was 81 at 10°C (specific gravity at 10°C = 0.6553). Due to the rapid evaporation, the density determination is not completely precise.

A very light coke was formed in the form of approx. 5 cm thick layer in the reactor. The coke had a volume of approx. 1800 cm3, but the weight was only 513 gr.

The dead space under the stirrer resulted in a delayed coking under the real reaction zone.

Example 3

This example was carried out in the same way as in example 2, except that a different form of catalyst was used and a more rapid heating of the first stage reactor.

The catalyst A in this example was a single layer of catalyst, and did not require the layer separation during the drying phase that was used for the catalyst as described in example 2. The sustained temperature in this experiment was 420°C.

In the 42 minutes of this experiment after reaching the operating temperature of 420°C, the catalyst converted,

which was approx. 40 gr on a theoretical anhydrous basis, the following bottom fractions from the original added 1300 gr of a solvent extracted asphaltene;

a) No. 1 bottoms fraction, collected under air-cooled coolers, a total of 29 gr hydrocarbons with an API density of 11.5 (specific gravity of 0.9895 at 15.5°C). b) No. 2 bottom fractions, collected under the water-cooled coolers and in the water-ice-cooled coolers, a total of 398 gr of hydrocarbons with an API density of 29 (specific density at 15.5°C = 0.8816). c) No. 5 bottom fractions, collected under the dry ice-2-propanol trap, a total of 63 cm3 with an API density of 83 at room temperature. A significant part of the No. 5 bottom fractions was lost during the determination of the density.

d) A total of 135 1 gas was produced.

The gases were measured after an alkali hydroxide washing of the gas vapours

(after the dry ice-2-propanol trap). These gases were not collected for analysis, but a milder analysis of similar trials yielded approx. 5% (volume%) non-hydrocarbon gases such as hydrogen, carbon monoxide and carbon dioxide. The remaining hydrocarbon gases had an average molecular weight of 49, and on the basis of this it can be calculated that approx. 280 gr of hydrocarbons in said gas.

e) The same light coke as in example 2 was formed

in this experiment. The weight of the coke was 489 gr.

The specified weights are: following;

In both examples 2 and 3, the catalyst in the second reactor was completely clean and free of tar, carbon etc.

The total settlement for products produced in example 2 was 1.045 7/1300 0 80.42%. The transformation, i.e. no. 1,2,5 bottom fractions + gases = 532.7/1300 = 40.97%.

The main difference between the two examples was that there was much higher gas production in example 3.

In these two examples, the bottom fractions were completely separated

from the water that condensed from the steam, and no emulsion was formed.

Example 4

In this example, a mixture of catalyst was used, i.e. about. 2/3 of a catalyst described in example 2, but not separated, and 1/3 of a catalyst from example 3 (K basis). The quantity ratio can be varied, and this also includes the quantity ratio of the catalyst layers in example 2. The catalyst was not placed on a substrate, and the weight was approx. 40 gr on a theoretical anhydrous basis.

The temperature in the reactor was kept at 420°C during this experiment.

The products obtained during the experiment were the following:

No. 1 bottom fractions 195.8 gr. (specific density 0.9793 at 15.5°C or

API 13) No. 2 bottom fractions 309.5 gr. (specific density = 0.86 at 15.5°C or

API 33) No. 5 bottom fractions 28.0 gr.

Gas 276.4 gr. (133 1. X 0.95/22.4 X 49 = 276.4 gr).

Coke 473.0 gr.

A total of 1,282.7 gr.

Total product = 1,282.7/1300 = 98.6% conversion

Conversion (no. 1, 2, 5 bottom fractions + gas = 809.7 gr /1300 gr) = 62.28%.

Coke formation occurred in 480 cm3 of the space below the stirrer, and some of this space is also below the spreading head.

Example 5

The main difference between this example and the previous examples was that a supported catalyst was used, which was added before starting the experiment. The catalyst was placed on sintered stainless steel meshes in 4 baskets. The steel was 3 mm thick and was cut into 3 mm wide strips which were again cut into 4.5 mm long pieces, and each piece for 3x3 was 4.5 mm. The catalyst was the same as in Example 2.

The catalyst on the steel pieces was placed in baskets measuring 0.6 x 6.25 x 15 cm, and 4 baskets were used and the baskets were placed on supports and stirred around using a stirrer, so that the baskets became the stirrer itself. A net held the catalyst on the steel in place, and the baskets themselves were placed on a frame.

With the same catalyst as in example 2 (and considerably less of the catalyst in the supported form) the amount of gas dropped from 135 to 90 litres. Most of the gas was produced at the end of the experiment when the temperature rose to 440°C.

The stirring speed was first 60 revolutions per minute, then it was increased to 120 revolutions per minute.

The starting material was 1290 gr. of solvent-extracted asphaltenes. Total product = 1204.5 grams/1290 grams = 93%. Conversion = No. 1,2 and 5 bottom fractions + gas/1290 = 77.90%.

o The API number for bottom fractions 2 and 5 was 32.0 at 15.5 C. Of the No. 1 bottom fractions, a distinction was made between the part which was liquid at 200°C and the part which was not liquid at this temperature. The liquid part had a bergnet API number of 5, and this part made up 228 grams of a total of 407 grams, or 56.03%. The remaining part appears to be a somewhat worked-up form of the solvent-extracted asphaltenes. A calculation of the conversion based on the liquid part at 200°C

of No. 1 bottom fractions gives a conversion of 64%.

In all these conversions and indication of total product, the produced gas was calculated as 95% hydrocarbons and where the hydrocarbon molecules had an average molecular weight of 49.

It is obvious that when the starting material is held under the spreading head, you get a competitive process. This includes a reduced temperature catalytic cracking with an improved limit with regard to thermal cracking, i.e. when steam is not present with the catalyst. When the catalyst is carried, there is no coke formation, if the steam, the asphaltene material and the catalyst are in intimate mixture. If there is insufficient steam or no steam, the asphaltene reaches the material, and you will then get a catalytically assisted thermal cracking. Although this is carried out at lower temperatures than a normal thermal cracking, and at atmospheric pressure, and at a speed of approx. 10 times faster than the normal one, this catalytic hydrocracking - the hydrogenation will be far more desirable because you get a high yield, high space velocities, you get a desirable and adjustable product composition and easily controllable reaction conditions.

Example 6

Using the method described in example 2, a mixed material containing different types of residues called FHC-353 was used as starting material.

An analysis for this material is as follows:

Metals

The starting material was gradually formed so that no coking was obtained until the mixed residues became pumpable and obtained a reaction temperature of 425°C, and at this temperature the reaction was carried out in the presence of catalyst and steam. The total weight of the starting material was 15,650 gr., and the recovery was 98.1% of the starting material based on the various condensates obtained.

A redetermination of the boiling point of this starting material according to the method ASTM-D-1160 and D-1180 was carried out on

a sample of 200 cm3 of this material, and they obtained data as follows (expressed as temperatures observed during the determination) :

This product was converted in a stirred basket reactor,

and where a supported catalyst was placed as described earlier. The substrate was a Y-82 zeolite from Union Carbide, Danbury, CT in the form of extrudates with a diameter of about 3 mm. The catalyst had an empirical composition of K2si 5'

i.e. it was prepared as catalyst D above and to which 5% by weight of catalyst A had been added. The amount of catalyst was 20.16 grams per 220 cm3 of the above-mentioned substrate.

A first conversion of the product (top product from reactor 110) was 71% (for total product, apart from highly volatile compounds from the dry ice trap and gas). This conversion does not include distilled material in the starting material which was recovered below 537.7°C. The top products + recovered bottom fractions made up 98.1% of the original starting material, while the rest

was gas and volatile distillates.

A sample of 1.392 gr contains 7 gr of water from the top products

from reactors 110 and 117, apart from the condensers mentioned above, had the following component distribution: Initial boiling point at 182.2°C, 71.6 gr or 5.24% by weight

182.2 to 343.3°C 203.4 gr or 14.89 -"-343.3 to 537.7°C 1084.1 gr or 79.36 -"-Total distilled: 1359.1 gr.

The product from reactors 110 and 117 had an initial boiling point of 182.2°C, and this product had the following properties:

The fraction boiling from 182 to 343.3°C had the following properties:

The fraction that boiled from 343.3°C to the end point had the following properties:

The recovered residium, i.e. what remained after a high temperature of 565.5°C had been reached, contained the following metals:

The vacuum gas oil fraction, i.e. the fraction boiling from 343.3°C and above had the following properties.

Initial boiling point was 344.4°C and the recovery was as follows:

at 9 4.4% recovery.

The K factor for the fully treated overheads from reactor 110 for 343.3°C plus the fraction (the hydrogenation as the last step) was 11.71.

The bottom fractions from the first stage reactor 110 recovered as product 112a were recycled on a batch basis when sufficient amounts were recovered. The catalyst for the starting material 112a was catalyst A placed on steel mesh particles of a size of 3 x 6 x 4.5 mm, and a total volume of 200 cm 3 containing 15.7 g of catalyst was used. Total starting material was 1286 gr. Total recovery was 750.7 gr. (this also includes 8 gr from the dry ice trap, but no gas). About. 29% of the starting material was recovered as bottom fractions. An analysis of the 112a bottom fraction was as follows:

About. 58.4% of the starting material was converted. On a continuous basis, higher conversion can be achieved, but on a batch basis, there is a dead space in the reactor under the rotating cages or baskets, which does not provide a complete reaction space, besides the reaction not being completely complete due to the constant presence of some material in the transmission lines.

Total recovered products had the following characteristics:

The fraction boiling from 182.2 to 343.3°C had the following properties:

The fraction that boiled from 343.3°C to the end point had the following properties:

Distillation residue:

343.3°C + product (from starting material 112a) had the following distillation characteristics for a sample of 191 gr, of which 163 was recovered on top, and the sample volume was 175 ml at 15.5°C.

The fractions won were as follows:

This distillation and residual product had the following properties: API numbers of 20.0 at 15.5°C and 16.4 respectively. Furthermore, the lower metal content and the relatively high API number will make this residue a good starting material, e.g. for a catalytic cracking.

Although the residue amounted to approx. 15%, then it will be the bottom fraction from the reactor vessel, i.e. vessel 130, and these fractions can again be subjected to further treatment, e.g. in a further parallel process in the same way as for the starting material 112a. An analysis of the residual product from the reaction vessel was as follows:

Based on the overall conversion of 58.37% (product recovered divided by the weight of the starting material with a total production of 83%) (retreatment of the bottom products 112a from reactor 110) and the top products from reactor 110, but where the liquids have not been included which was present in the dry ice trap, and not the recovered gases for the top products from reactor 110 and not assumed to be distillable products (assumed to be in the FHC-353 starting material which distills below 565°C), the total conversion is 87.9%. If, however, you include gases and products in the dry ice trap on an average molecular weight basis of 49 and convert the volume to weight, and here you also include the distillate content in FHC-353, then the real recovery is somewhat higher.

The dividends in the above examples are calculated on the basis of

that these dividends will also be consistent with normal practice. The conversion with respect to space-time speed can also be calculated, i.e. treated starting material per unit volume of the catalyst per given unit of time.

Dividends are usually calculated on the basis of the total

amount of the starting material and the conversion are calculated for the components that boil at a boiling point of less than 5 6 5.5°C. The boiling points are determined by previously stated methods which are those used in the petroleum industry. In the same industry, a temperature limit of 565.5°C is usually used

to describe very persistent components in crude oil or in distillation products.

The yields can also be based on the part of the components that have been converted into products with a lighter viscosity, more than on

an API product base. The residues which were collected by the cyclone 116 and which were used as starting material 112a were calculated into the yields as an addition to the original amount of starting material which was calculated and recovered. The residues from vessel 130 boiling at or above 565.5°C were then subtracted from the total recovered material provided that these residues are not further processed.

Sometimes the yields are based on the treatment of residues with certain boiling points, i.e. those which boil at temperatures between 343 and 454 .4°C, and only take into consideration the part which can be converted by ordinary known methods, whatever these may be. However, this method is not very well suited for evaluating the present method, as in it almost all previously unconvertible residues are converted, and it is therefore assumed that it is best to characterize the products obtained by a direct distillation, and where one takes the products that can be distilled at less than 565°C in a first or another parallel treatment as additions. The yields are therefore calculated on the basis of the products from the first direct distillation and the second and third parallel treatment, thereby also including products from reaction vessel 117, and all the product amounts are combined by the distillation described, after which the residues that boil above 565 are subtracted, 5°C.

In practice, gaseous hydrocarbon products will often be included in the starting material which can be distilled at temperatures below 565.5°C, as well as low-temperature condensable compounds. The gas can be included on the basis of a milder molecular weight calculated from the volume that has been collected, after which the volume is converted to weight% to include all products.

It appears from the above description and example 6 that

by a specific treatment of separated fractions from petroleum crude oils, low-boiling residues that boil up to 454.4°C or high-boiling residues up to and even above 565°C, direct distillation at low pressure and at low temperatures can convert almost all the hydrocarbon values in crude oil or in petroleum residues to usable products if one includes a different or additional treatment, as indicated here with specifically selected catalysts.

In the past, considerable efforts have been made to convert insoluble parts of the crude oil or petroleum residues into usable components, but the yields have been low and numerous recycling treatments have had to be carried out in order to achieve an acceptable yield. It should be emphasized that previously up to 50% of these unmanageable parts had to be subjected to a very strong treatment, so the present parallel method, which gives a significantly higher yield and significantly better end products, will at the same time achieve a higher space-time speed, and at the same time if poorer starting materials are used and the product is treated under less drastic conditions, it is obvious that the present method will be a significant contribution when it comes to utilizing all usable products from a given volume of crude oil or crude oil residues.

Claims (1)

1. Process for converting carbon-containing starting materials such as crude oil, residues of crude oil and the like with stable components without the addition of hydrogen gas as a reactant, in order to thereby produce products with a lower viscosity and/or end products that are more hydrogenated, and where said starting material is converted in the presence of an alkali metal sulphide catalyst, characterized in that one: a) reacts in a first reaction zone in the presence of added water in the form of steam and a catalyst, at least one carbonaceous starting material which includes a heavy crude oil, natural asphalt compounds, natural tar, pitch, Gilsonite, a suspended oil, a solvent extracted asphaltene, pitch or tar prepared from coal, petroleum residues, oil, resins and asphaltene mixtures or an oil - resin - asphaltene fraction from a distillate with a boiling point of up to 454.4°C and higher , a distillation residue without any boiling point below the point at which a destructive distillation of the same residue takes place, a coal oil extract, a bottom fraction of a processed shale oil; a heavy bottoms fraction produced by coal gasification of the SASOL type, a delayed coking distillation bottoms product or mixtures of the aforementioned, and where said catalyst consists of an alkali metal sulphide catalyst and a further cleavage catalyst in an amount of at least 5% in order to thereby pre-treat the more resistant components of said starting material in said first reaction zone, and where the cleavage catalyst is in a mixture with said alkali metal sulphide catalyst in order to thereby treat said resistant components in said starting material, and where said cleavage catalyst is a supported or unsupported catalyst composition containing: 1) a first solution of an alkali metal hydroxide dissolved in methanol, ethanol, 1-propanol or 1-butanol or mixtures of these alkanols, or 2) another solution of said alkali metal hydroxide - alkanol as defined in point 1 above, to which water-dissolved alkali metal hydroxide has been added and where said a alkali metal hydroxide on a mole basis in said solutions, is from 0.5:1 to 1:0.5, and where the first and second solutions are saturated with hydrogen sulphide, so that in each of the solutions 1) a single phase solution or 2) two phase solutions are formed. and where said catalyst composition is either said single phase solutions of (1), said two-phase solution of (2), each of the phases under section 2 taken individually, mixtures of the phases under section 2, or a mixture of each of the individual phases of (2) with single phase solutions of (1), or a mixture of the two phases from (2) together with the single phase from the single phase from (1); b) recovers a top reaction product from said first reaction zone and this includes gases as volatile and gaseous products, and separates bottom products; c) recycles the bottom products or separates the bottom products in a parallel stream from the products from said first reaction zone, which are liquid products, entrained liquid products and partially pretreated resistant components of said starting material; d) reacting in the presence of the first supplied steam or later added steam, the separated bottom products in the presence of said cracking catalyst as defined under paragraph (a) above and said alkali metal sulphide catalyst, and e) recovering the products prepared in step (d). 2. Method according to claim 1, characterized in that the catalysts used in steps a and d are placed on a metal. 3. Method according to claim 1, characterized in that the products from step c are recycled instead of being separately processed. 4. Method according to claim 1, characterized in that the vaporous and gaseous reaction products from the top in step b or step d are further reacted directly after cooling the products, in at least one further reaction zone in order to thereby hydrogenate the products in the presence of a supported catalyst and added steam , and where said supported catalyst includes an alkali metal hydrosulfide, sulfide, polysulfide, a hydrate of a sulfide, a hydrate of a polysulfide or mixtures thereof, so that the bromine number of the product is reduced compared to the bromine number of the products from step b or step d. 5. Process according to claims 1 and 3, characterized in that the vaporous and gaseous top reaction products from step (b) are further processed directly after cooling the products in at least one further reaction zone, in order to thereby hydrogenate the products in the presence of a supported catalyst with added steam , and where said catalyst includes an alkali metal hydrosulfide, sulfide, polysulfide, a hydrate of a sulfide, a hydrate of a polysulfide or mixtures thereof, so that the bromine number of the product is reduced together with the bromine number from the products in step (b). 6. Method according to claim 1 or 5, characterized in that the catalyst in step (a) or step (d) is a supported catalyst and includes a/ Catalyst A, or 2/ Catalyst A in an amount of up to 95% on a K mol basis of catalyst A, in mixtures with catalysts B, C, D or E, or mixtures thereof, and where each of the aforementioned catalysts is prepared as defined here: where catalyst A is prepared by dissolving a mole of potassium hydroxide in an ethanol, methanol, a ethanol-methanol mixture, 1-propanol or 1-butanol, after which one reacts the potassium hydroxide solution with hydrogen sulphide that bubbles through the solution, recovers the catalyst - alcohol mixture and separates said alcohol from said solution, and where catalyst B is prepared by dissolving potassium hydroxide of technical or analytical quality and containing approx. 86% potassium hydroxide in ethanol or methanol, and saturate the solution with hydrogen sulphide in a series of vessels by introducing said hydrogen sulphide into a first vessel without boiling off the alkanol, collecting and condensing each alkanol given off by an exothermic reaction between said potassium hydroxide and hydrogen sulphide in a subsequent vessel, and where the addition of the hydrogen sulphide is stopped when the last vessel containing KOH shows a reaction, and where catalyst C is prepared by dissolving approx. 6 mol KOH in approx. 4.5 to 7.5 mol of water without external heating, after which a small amount of an alkanol is added, from approx. 2 to 2.5 cm3 of methanol or ethanol per moles of KOH; adds approx. 4 moles of elemental sulfur, and adding an appropriate amount of sulfur to adjust the catalyst to the desired sulfur content by adding supplemental amounts of sulfur to produce an empirical sulfide from K2S^ ^ to K2S5' and wherein catalyst D is prepared by dissolving 1 mole KOH in 1.0 mol of water, after which 2 ml of methanol or ethanol are immediately added after said potassium oxide has been dissolved; adds 2/3 mole of elemental sulfur, adjusts the catalyst to the desired empirical sulfur content by adding sulfur with further stirring, and wherein said catalyst composition varies from ^ to K2S5' and where catalyst E is prepared by adding a dried KHF powder or a suspension of such a powder to each of the above described catalysts or mixtures of catalyst A, B, C or D, and where said addition takes place with from 1/5 to 1 /3 mol on a molar basis of K of KHS to K2S (empirical) sulfide to ~& 2^ 2 5 (emPir;>-sk) on a molar basis. 7. Method according to claim 1, characterized in that the catalyst in step (a) or step (d) is a supported catalyst and that the substrate is a porous metal, chromite spinel, zeolite or an aluminum oxide. 8. Method according to claim 1, characterized in that the catalyst in step (a) or (d) is a catalyst supported by a porous metal substrate and where the latter is a stainless steel with up to 35% metal by volume. 9. Process according to claim 6, characterized in that the carbonaceous starting material that is reacted in step (a) is a heavy crude oil, or a residue without any boiling point or an asphaltene, and that the supported catalyst contains at least 75% by weight of catalyst A in combination with catalyst D deposited on a catalyst substrate. 10. Method according to claim 9, characterized in that the catalyst contains from 5 to 80% by weight of catalyst A, while the rest is catalyst B. 11. Method according to claim 6, characterized in that the catalyst in step (d) is a catalyst A placed on a porous metal substrate and where the reaction is carried out at temperatures from 360 to 440°C. 12. Method according to claim 1, characterized in that the reaction is carried out at a pressure from subatmospheric pressure to less than NOK 10.7 per m2 at temperatures from approx. 160 to about 600°C, in the presence of a catalyst placed on a porous metal substrate. 13. Method according to claim 6, characterized in that the catalyst for the further reaction is a catalyst A composition with KHS as an added component in an amount of up to 15% by weight, and where the resulting catalyst product is placed on a porous stainless steel substrate. 14. Method according to each of the preceding claims, characterized by the substrate for a catalyst for reactions as defined in claim 1, is protected and is a porous metal, a chromite spinel, an aluminum oxide, a zeolite or a mixed substrate. 15. Method according to claim 14, characterized in that the catalyst is placed on a substrate, and in that this is aluminum oxide with a pore size of from 6 to 13,000 Angstroms, preferably 350 to 900 Angstroms, and wherein said aluminum oxide is protected against attack from said catalyst in that said catalyst is deposited in a mixture with glycerol, after which the catalyst and the substrate can be defined in the absence of oxygen, or deposited together with a polyfunctional alkanol with up to 6 carbon atoms, and where either glycerol or polyfunctional alkanol is deposited on said substrate and heats the substrate to 200°C, after which it cools and after which said catalyst is deposited on the substrate, which is then again heated to approx. 560°C. 22. Method according to claim 6, characterized in that the reaction in step (d) takes place in a boiling bed reactor, and in that said catalyst is placed on a zeolite and in that the catalyst consists exclusively of catalyst A. 23. Method according to claim 6, characterized in that the reaction in step (a) takes place in a boiling bed reactor, or a reactor with rotating curves, and in that the catalyst is a mixture of catalyst A and catalysts B, C and D or a mixture of the latter, and where the amount of A is from 3 to 95% by weight, and where the catalyst is placed on a zeolite, a porous metal, a chrome spinel or an aluminum oxide. 24. Method according to claim 6, characterized in that the reaction in step D is carried out in a boiling bed reactor or in a reactor with rotating or spinning baskets, and in that the catalyst is a mixture of catalyst A and catalysts B, C or D, or a mixture of the latter, where the amount of A varies from 50 to 100% by weight. 25. Method according to claim 24, characterized in that the catalyst A constitutes from 50 to 95% by weight. 26. Method according to claim 6, characterized in that the top product from steps (a) and (d) is reacted further before they are hydrogenated. 27. Process according to claim 26, characterized in that the top products from steps (a) and (d) are hydrogenated with catalyst A with added KOH, catalyst C with added KHS, catalyst B and A with more than 50% of catalyst B on a K mol basis, or by means of catalyst B.