DE3407476C2 - - Google Patents

Description

The present invention relates to a method of conversion from petroleum or petroleum residues to lower viscosity products and / or more hydrogenated products, in particular for Converting according to ASTM-D-1180 and D-1160 high-boiling or non-boiling petroleum residues.

Heavy oil fractions are known to be common only difficult to recycle to high quality products. Hydrogenations of Petroleum residues require very high temperatures and possibly the presence of hydrogen under high pressure.

GB-PS 20 75 542 describes a method for hydrogenation of petroleum residues, in which these in the presence of a Alkalisulfidkatalysators brought into contact with water vapor become.

The object underlying the invention was to to improve this known method and petroleum or petroleum residues as light and complete as possible in products of lower Viscosity and / or starch to convert hydrogenated products.

This is also intended for particularly heat-resistant petroleum residues with a boiling point above 454.44 ° C, especially asphaltenes or mixtures of oil, resin and asphaltenes, which are refined of petroleum as a residue and one after the Determined according to ASTM D-1180 or D-1160 Have boiling point of 573.78 ° C and higher, apply.

To solve this problem, the inventive method for converting petroleum or petroleum residues into products of lower viscosity and / or more hydrogenated products in the presence of an alkali sulfide catalyst without the addition of gaseous hydrogen, characterized in that in a first reaction zone in the presence of water vapor and an additional catalyst the petroleum or petroleum residues at a temperature of 160 to 650 ° C and a pressure of up to 10 bar, wherein the additional catalyst from at least 5 wt .-%, based on the total amount of catalyst, a saturated with hydrogen sulfide solution of 0.5 to 2 moles of alkali metal hydroxide in one mole of methanol, ethanol, 1-propanone, 1-butanol or mixtures thereof, optionally containing water,
separating the reaction product withdrawn from the first reaction zone into a top fraction containing vapor and gaseous products and a bottoms fraction and returning the bottom fraction either to the first reaction zone or in the presence of water vapor and the above two catalysts at the above or lower temperature and pressure implements.

The inventive method differs from the state the technology by the special catalysts, which the extremely heat-resistant constituents of petroleum residues, z. B. petroleum residues with a boiling point above 537.78 ° C, in very effectively convert into products of lower viscosity. In addition, the method according to the invention differs from the prior art in that in the implementation the most heat-resistant components of the residue, for example Asphaltenes, are specifically attacked and this Coke formation is avoided. In addition, arise procedurally further advantages from the use of a Siedebettreaktors.

The inventive method leads to increased overall yield and is able to the Ramsbottom carbon or the Conradson carbon into useful hydrogenated products too convict. The crude oil components are almost complete converted into useful products.  

The method according to the invention is also applicable to others Products like Gilsonite (also known as Uintait, an in Utah occurring black shiny asphalt), pitch (on Natural product and / or refinery product, tars, such as from tar pits in California, asphalts (natural) like the one found in asphalt lakes on Trinidad and in Utah Asphalt, and similar natural products or refinery processes, which oils, resins and asphaltenes in the residues concentrate, available products, in refineries accumulating sludge oil, which often no Owns distillation point, sump oil from coal gasification plants, for example, products of solvent extraction coal, coal extracts, soil fractions of shale oils, heavy soil products from the Lurgi gasification process, Sulfur products as produced by distillation, and both hard and soft resins and asphaltenes are applicable. It It is recommended to use some of the starting materials together However, this may be due to the prevailing reaction conditions be determined accordingly by experiments. Under Heavy crudes are crude oils with an API Zagl of less than 22, including the crudes with a negative API number, Roger that.

In the method according to the invention, the more volatile Petroleum fractions and similar hydrocarbons directly driven off, while the heavier soil products, so the heavier and in particular the more heat-resistant constituents, be recycled or treated in parallel. For example become the starting materials that require the greatest amount of treatment best treated after the parallel procedure. The parallel process in the second reaction zone treated bottom fraction of the first reaction zone can subsequently be further implemented.

For example, in such a parallel process the mixture of particularly heat-resistant products of a heavy soil fraction end products with a K value of about 11,17, so high molecular weight paraffins, what in terms of  is achieved that the top products from the first reactor easier to further process, for example reduce the bromine number and achieve similar.

The alkali metal sulfide catalysts are suitably alkali hydrogen sulfides, sulfides, polysulfides, sulfide hydrates and / or polysulfide hydrates. Preferably, they are supported catalysts.

Using such catalysts, the top fraction becomes the reaction product of the first reaction zone conveniently directly or after cooling in at least one further reaction zone in the presence of water vapor while reducing the Bromine number hydrogenated.

The reaction in the first reaction zone as well as the reaction the bottom fraction is conveniently adiabatic at up to 560 ° C. The reaction takes place in the second reaction zone appropriate at 320 to 450 ° C. The bottom fraction is expediently reacted at 360 to 440 ° C.

In the drawing means

Fig. 1 shows schematically the process carried out in an air-bed reactor according to the invention with recycling of the bottom fraction and

Fig. 2 shows schematically a continuous process according to the invention, in which the bottom fraction is treated separately.

As shown in Fig. 1, the reaction with the first reaction zone is a bottom-boiling reactor 10 having a reaction vessel 11 equipped at the top with a catch funnel 15 and a pump 14 for circulating the volatile or liquid products 8 participating in the reaction. The catalyst 9 is a dispersed supported catalyst suspended by the medium 8 . A participating in the implementation and the catalyst 9 contained medium 8 can be circulated by pumping the medium into the reactor and distributing within the reactor. Typically, the recycle medium is supplied in the region of the bottom of the reactor, but it may also be supplied to any other location or locations of the reactor. On a smaller scale, a stirrer may be provided as a close approximation to an aerated-bed reactor, as long as the supported catalyst is in one or more stainless steel cages or baskets and these cages or baskets are suspended from a suitable frame driven by an externally located motor is. By varying the speed of, for example, 20 to 180 min ^-1 , in particular 5 to 150 min ^-1 , the reactions can be largely aligned with the reactions occurring in a boiling-bed reactor reactions. It is also important to have sufficient contact between the steam, catalyst and medium.

A preheated starting material, such as asphaltenes or an OHA cut (oil, resin, asphaltenes) is introduced via a line 12 . Steam 13 may be introduced into the reactor 10 along with the feedstock, but may also be distributed from the bottom of the reactor 10 in the medium.

The resulting liquids of lower viscosity are a large-diameter discharged via a conduit 15 and, if necessary cooled in a quench zone 16 in a suitable manner and then introduced into the reactor 17 of the second stage. Although only one reactor is shown in the figure, multiple reactors connected in series or in parallel can also be used. These reactors may contain the catalyst as a fixed bed catalyst or be gas phase or steam reactors. A column in which the catalyst in the manner shown is supported on trays 17 a, is typical. Other known devices such as trickle bed reactors may also be used for this purpose. The bottoms product from the second reaction stage is collected in catchment zone 18 .

This bottoms product can be recycled in whole or in part together with the collected in the collecting funnel 15 medium in the boiling bed reactor 10 . If necessary, a pump 14 a can be used for this purpose. It can also be diverted a portion of the product and obtained for further processing. The top fraction from the second stage reactor 17 may be heated to reflux in a reflux kettle 19 , and the gaseous products thereof may be further worked up in a second reactor which is similar but not shown to the reactor 17 . All of these reactors are referred to as second stage reactors, unlike the first stage boiling-bed reactor 10 , in which the feedstock is split. The products from the second stage reactor and other gaseous products may be treated in another reactor, for example by passing through a suitable bath any undesired constituents, such as hydrogen sulphide, are separated. This can be done in a vessel 20 in which potassium hydroxide is dissolved.

The gaseous products are then in common Won way.

As shown in FIG. 2, the first stage reactor 110 is a fluidized bed reactor (boiling bed reactor). This reactor may also be a reactor with hinged basket or stirred cage. The stirring cages or baskets are not shown. However, since sifting bed reactors (fluid bed reactors) and stirred cage reactors are largely similar in performance and operation, these reactors are considered equivalent.

The reactor 110 is constituted by a reaction vessel 111 having a trap, such as a hopper 115 , and a pump 114 for circulating the feed 108 to be reacted. The feed 108 may be petroleum or a petroleum residue boiling below 454 ° C, but may also be an OHA fraction (oil, resin, asphaltenes) boiling above this temperature, for example above 565 ° C. A dispersed supported catalyst 109 is circulated with feed 108 . A sieve 109 a prevents the catalyst from flowing upwards together with the boiling charge 108 . The feed containing the catalyst, if the catalyst remains suspended therein, may be circulated by means of the pump 114 , but only a liquid phase trapped in the catch funnel 115 and freed from the catalyst may be circulated. As shown in the drawing, the boiling feed pumped by the pump 114 can be introduced anywhere in the reactor, although it is usually fed to the bottom of the reactor or near that bottom.

The feed 108 of the type described above is combined with steam and / or water 113 in a single inlet, but separate inlets may also be provided on the periphery of the vessel 111 for these substances. The heat-resistant petroleum components, which require a longer residence time, undergo the catalytic reaction and are discharged via the outlet conduit 115 a , which leads diretk into a cyclone 116 , in which the vaporous products of the liquid products and / or entrained by the vapor from the boiling bed Products are separated. Several cyclones 116 can also be used. The crude product 112 a from the cyclone is processed in a separate parallel flow.

The overhead product from the cyclones 116 can be quenched in the example illustrated in the drawing, in a quench zone 116, and then reacted immediately in a reaction vessel 117th It may be sufficient not to deter the products prior to introduction into the reaction vessel 117 if it can be quenched in the latter. The reason for the quenching will be explained later.

After insertion into the vessel 117 a later-described suitable carrier catalyst or unsupported catalyst in the manner shown schematically on a tray 117 a or by other means is suspended. The catalyst causes the reaction in the gas stream or liquid stream to a lighter fraction and a heavier fraction or causes hydrogenation of the gas stream. The bottoms product can be collected, for example in a vessel 118 , and worked up separately in a further parallel flow in the usual way or in a manner to be described later.

Further, not shown in the drawing reaction vessels on the type of reaction vessel 117 may be provided in the same flow path. The final removal of the gases from the reflux condenser 119 or other coolers, not shown, can be made in the usual way by cooling, if necessary under pressure.

It has proven advantageous to treat the gaseous components of the distillate obtained after the initial treatment in the vessel 117 at a temperature of about 130 ° C to 140 ° C in a reaction vessel of the type indicated in the figure 117 on. If more saturated products are desired, a properly selected and adjusted reflux column 119 can be used to capture these products.

The bottom product 112 of a separate, for example, in the cyclones 116, the reaction product is usually from more heat-resistant components of the crude oil or petroleum residues. This bottom product 112 a, to accept from which is that it has been implemented in the reactor 110 is partially (which results, for example from the improved API number), will best be further treated in a first parallel branch. Since this bottoms product has been reacted to some extent because of the catalyst 109 suspended in the digester bed, further reaction is best done with another specific catalyst, modified if necessary to suit the particular feed.

This second feed stream thus consists of the bottom product 112 a , which is supplied to the reaction vessel 130 together with steam. This feed 112 a is more amenable to hydrogenation, thus further improvements are achieved by ingredients of lower viscosity are obtained.

As a non-illustrated further reaction vessel may follow the reaction vessel 130, which gained in the reflux column 133 overhead products from the reaction vessel can be further processed 130th Usually exiting from the top of the reflux column 133 products by washing in a vessel 135 of undesirable components, such as H₂S, freed and the gaseous products are then recovered and reused. The reaction vessel 130 may be operated in a manner similar to the reaction vessel 117 .

If no further reduction in the molecular size of the top distillate from the first reactor is desired, this headstock vapor should be hydrotreated in the vapor phase of the reactors immediately upon exiting the cyclone 116 which separates the bottoms fraction from the top distillate. Quenching, for example in a quenching device 116a , prevents re-condensation of the cleaved components. The hydrogenation treatment is facilitated by the absence of the heavier fractions separated in the cyclone 116 . Hydrogenating treatment is by definition carried out without thermal or catalytic cracking by introducing hydrogen into a molecule. Since the hydrogenation reaction is exothermic, the aforementioned quench 116 a helps to control the temperature and to avoid degradation of the products by excessively high temperatures.

For this hydrogenating treatment can be fixed bed reactors or especially reactors for liquid catalysts be used.

The feed material, typically one of oil, resin and Asphaltene existing OHA fraction, is a material which at atmospheric pressure above 583 ° C boiling. Asphaltenes are for example, a catalytic processing particularly inaccessible and thus represent a faction which according to known methods only by burning or by known coking could be utilized.

If the bottom product 31 is difficult to split, it may be subjected to a further operation of the type shown for the feed 112 a and in particular in FIG. 2.

Since most of the top products from the reaction vessel 11 are treated in a single pass, the bottom fractions 12 a collected in, for example, the cyclone 16 and the collecting vessel 18 form desired products which, however, account for a smaller proportion depending on the composition of the petroleum. These soil products can thus be treated in smaller vessels taking into account the proportions of the constituents in the processed petroleum product. Accordingly, a variety of mixtures of various petroleum oils depending on the components of the bottom product 12 a can be beneficially treated. The size of the reactor 30 may be sized according to the proportion of bottoms components in the petroleum, for example, those having a boiling temperature of 583 ° C and above.

Water vapor is supplied to the reactor 110 or to the reactor 130 at such a rate that the desired product dictates to some extent the amount of steam used. The steam can be used in an amount of about 50 to 100 and even 130 wt .-% of the recovered product. The lower limit of the amount of steam results from the fact that in the reactors 110 and 130 a coking would occur if steam would be present in an insufficient amount. The lower limit of the amount of steam supplied must therefore be such as to avoid any appreciable coking in the reactors 110 , 117 , 130 , etc. Additional steam can, for example, for the reactors 117 , 130 , etc. are supplied. As steam, low-tension steam, typically steam, is used.

The temperature in the reactor may generally be up to 650 ° C, although the most advantageous working temperature is below 450 ° C, for example below 425 ° C. In the following reaction vessels, for example in the reaction vessel 117 , however, lower temperatures can be maintained at which a more hydrogenated product is formed. However, the treatment of the more heat-resistant component in the reactor 130 may require the same temperature as in the reactor 110 .

Quenching may be carried out at a temperature of 390 ° C or below. The third stage reactor, not shown, may have a temperature in the range of 170 ° C and below down to 125 ° C (without condensation of steam). Lower temperatures are desirable because the hydrogenation is exothermic and most of the hydrogenation is carried out in the downstream reactors whereas most of the cracking is in the reactor 110 and the minor part of the cracking is in the reactor 130 .

The pressure in the vessel 111 and downstream thereof can be maintained at atmospheric pressure. Although pressures under atmospheric pressure are possible, they are not so convenient. Pressures between about 6.7 and 10 bar, in particular between 4 and 10 bar, are preferred. When working at higher pressures, condensation of steam must be avoided. Furthermore, instead of separating the liquid and / or entrained products in cyclones, such as the cyclone 116 shown in the drawing, another separating means may be provided, such as centrifugal separators or separating drums.

The treatment can be a variety of crude oils of any origin and any viscosity, if they are sufficiently fluid at the treatment temperature. Even with the usual oil refining under atmospheric pressure accumulating residues and soil products can be used. These starting materials are thus products which have a boiling point of less than about 454 ° C. Such with higher boiling point ingredients will be the best processed in parallel.

The inventive method provides products unusually improved viscosity, improved hydrogen content and better chemical structure, such as products with a higher content of high K paraffins or aromatic polycyclic compounds, which subsequently still to desired, olefinically less unsaturated cyclic Further processed compounds and linear or branched paraffins can be.

The hydrogenating treatment for reducing the bromine number and to reduce nitrogen and sulfur content the relatively low temperature is not appreciable to a crack, which reduces the boiling range would become.

Hydrogenating treatment should therefore form the final stage of the process described here and should only be carried out if the desired molecular size has been achieved by prior cracking. Thus feedstocks formed by residues from the vacuum distillation are initially cracked under less harsh conditions to smaller molecular size, and then the heat-resistant residues, namely, the feed stream 112 a, under the required harsh conditions which do not adversely affect the products of the main stream, cracked.

The smaller molecules forming the "cracking products" contain less nitrogen than the starting material. The metals and the nitrogen remain in the unreacted porphyrin backbone of the asphaltenes and are therefore not present in the top product of the asphaltene distillation or only in very small amounts. The metals and the nitrogen are present in the bottoms product, for example 112 a , in increased concentration. The sulfur content is not necessarily increased in the asphalt portion of the crude oil or in the residue. The resin content of the residue may contain up to 5% by weight of sulfur.

A suitable high efficiency catalyst for the reactor 110 is referred to as catalyst A and prepared in the following manner. 1 mole of potassium hydroxide is dissolved either in ethanol, methanol, most preferably in ethanol, or in a mixture of ethanol and methanol or, which is less advantageous because of the lower solubility, in 1-propanol or 1-butanol. The potassium hydroxide dissolved in this solution is then reacted with hydrogen sulfide passed through this solution. After complete saturation, the catalyst-alcohol mixture is separated and the alcohol separated by suction in vacuo.

The catalyst may be in the first reaction zone without carrier However, is used here and in the following reaction steps preferably applied to carriers and calcined.

Suitable catalyst supports are spinels, such as chromium spinel, Toner earth and most advantageously porous metals, such as stainless Steel, with a pore size up to 350 nm and more.

The catalyst preparation and its deposition on the support done in the absence of oxygen. After the abortion the volatiles become the catalyst together with the carrier to a suitable temperature, for example a temperature  from 320 to 450 ° C, heated. The catalyst then adheres firmly on the carrier.

The reaction in the first reaction zone may be at a higher Temperature, for example, a temperature in the range of about 320 to about 450 ° C, although even temperatures up to 560 ° C and even up to 650 ° C can be applied. For asphaltenes, the preferred temperature range is sufficient from about 360 to about 430 ° C, where it seems that a very good Working temperature range the range between 390 and 425 ° C is.

The steam is preferably in an amount of at least about 27% by weight of the feed, for example the OHA fraction, of crude oil or petroleum residues. If a stronger hydrogenation is desired, more water vapor however, the amount of steam typically exceeds not about 85% by weight of the feed, although based on the resulting hydrogenation product fed itself 130 wt .-% steam can be. In other words, the amount to be used of water vapor determined by the fact that the hydrogen content the charge of the hydrogen content of the desired Product subtracted and the difference obtained is multiplied by a factor of 9 (since the molecular weight of hydrogen ¹ / ₉ of the molecular weight of water. Typically, water is added to the first reactor introduced in an excess of 30%.

If in the reactor, water, for example in the form from steam, not introduced or the water supply off For some reason, it may become coking come, with carbon being deposited. By inadequate Steam supply or batch steam supply will also causes the formation of particularly heavy products in the reactor. It is important that enough steam is always introduced into the reactor and is thoroughly distributed in it.

On the other hand, excessive steam supply interferes with a flawless Reaction process obviously because partially implemented  Products are carried away.

If carbon separates for any reason, it usually happens at hot spots, such as the heated reactor walls or catalyst support. The reactor 10 is therefore preferably operated adiabatically. Carbon deposits on the catalyst can be eliminated and converted into useful products by exposing the catalyst to the action of the vapor for a certain period of time without adding further starting material, so that the catalyst is then again usable for the production of the desired product desired fraction can be used. Bulk supply of hydrogen sulfide or sulfur may be useful in general, and especially for low sulfur content feedstocks, for example, sweet crudes.

It has already been found that when the working temperature is increased to about 320 to 420 ° C, depending on the composition of the feed, for example, due to their content of asphaltenes, an exothermic reaction can proceed. The exothermic reaction can cause a temperature increase up to 600 ° C, however, this temperature increase is dependent on the amount of water vapor supplied. A larger amount of steam gives rather lighter carbon compounds. Excessive temperatures are not desirable, so temperatures below 440 ° C are preferred. In the reactor 17 shown in Fig. 1 or in the second reactor 117 shown in Fig. 2, where the further reactions take place, the products coming from the first reactor rapidly in the presence of a catalyst to about 250 ° C, usually on Temperatures between 250 and 390 ° C cooled. The quality of the product is increased with increasing working temperatures to 430 ° C, preferably 425 ° C, but there is no increase in the degree of conversion above 390 ° C. The cooling in the cooler 16 a takes place to such an extent that steam does not condense and disturbs the reaction. The catalyst in the second reactor 17 and 117 is preferably also a supported catalyst, however, the catalyst can also be used without a carrier. The catalyst supports may be the same as in the first reaction stage, such as alumina or zeolites.

A typical catalyst for the second stage reaction, ie, in Reactor 17 or 117 , is a less effective catalyst B. This catalyst is prepared by mixing technically pure or reagent grade potassium hydroxide containing about 86% potassium hydroxide in absolute or 95% strength. ied ethanol or methanol (preferably ethanol) and the resulting solution is saturated without hydrogenating the alkanol with hydrogen sulfide.

A designated Catalyst C catalyst is dissolved by dissolving 6 moles of KOH in 4.5 to 7.5 moles of H₂O without heating from the outside, where per mole of KOH a small amount of an alkanol, for example 2 to 2.5 ml of methanol or ethanol, is added. The resulting solution is then added 4 moles of elemental sulfur, which reacts exothermically. Subsequently, elemental sulfur is added to obtain a sulfide of the desired empirical formula, which may range from K₂S to K₂S₅ but can also range from K₂S _1,1 to K₂S _2,5 , depending on which end product is desired. For a higher gas content in the product, catalysts are used which are less saturated in sulfur. For a higher liquid content in the product, more saturated catalysts are used with sulfur.

A further catalyst, called catalyst D, is prepared as follows. 1 mol of KOH is dissolved with vigorous stirring in 1.0 mol of water. Immediately after the KOH has dissolved, 2 ml of methanol or ethanol are added. Immediately thereafter, 0.66 moles of elemental sulfur are added, whereupon the sulfur is left to react in the course of a more vigorous reaction. The catalyst is brought under vigorous stirring by adding a suitable amount of sulfur to the desired empirical sulfur content, for example, ¼ of ² / ₃ moles of dissolved sulfur K₂S _1.5 and ½ of ² / ₃ moles of dissolved sulfur K₂S _2.0 . The composition of the catalyst may be in the range of K₂S _1,1 to K₂S _2,5 or even to K₂S₅. The lower sulfur content catalysts in admixture with Catalyst A are good crevice catalysts and are generally good catalysts for less severe treatment of crude oils. This catalyst is also a good hydrogenation catalyst. The catalyst A is added to the catalyst D usually in an amount of 3 to 25 wt .-%, usually in an amount of less than 10 wt .-%, added.

Once the catalyst is prepared in this way it becomes a flowable slurry in vacuo then evaporated over with glycerol treated and cooled as described above extrudate (if the carrier is up to 300 ° C or higher has been heated), being stirred in a high vacuum and finally sucked to dryness. The catalyst is further sieved in the dry state and immediately then introduced into the second stage reactor, which, if the catalyst as a hydrogenation catalyst used, was rinsed to atmospheric oxygen to displace.

A further catalyst, referred to as Catalyst E, is a non-supported catalyst which is capable of reducing the molecular size of the product in a first stage reactor (or which may be used in a further second reaction stage). Catalyst E is obtained by adding powdered dry KHS or a slurry thereof in suitable amounts of any of the above-identified reagent mixtures (catalysts) A, B, C or D. There are ¹ / ₅ to ¹ / ₃ mol, based on K mole of KHS, for example, added to K₂S or K₂S _1.5 , which is reduced by these additions of KHS, the molecular size of the product.

Favorable catalysts for the first reaction zone are supported catalysts from catalyst D with 5 to 80, preferably at least 75% by weight of catalyst A. Another favorable catalyst for the first reaction zone consists of a zeolite as support with only catalyst A or with 25 wt .-% of catalyst A and 75 wt .-% of catalyst B or from From 3 to 95% by weight of catalyst A mixed with catalyst B, C or D on zeolite, porous metal, chrome spinel or alumina.

The catalyst for the further treatment of the soil fraction is suitably a mixture of catalyst A with catalyst B, C or D, preferably with 50 to 95 wt .-% of catalyst A. The top fraction is appropriate before hydrogenation with a KOH-added catalyst A, a KHS-added catalyst C, Catalyst B or a mixture of catalyst A and B further implemented.

From the feed vanadium can be completely removed. The separation of nickel is promoted when part of the catalyst is present as hydrogen sulfide. Nickel and iron form with the alkali metal sulfides and hydrogen sulfides complexes similar to ferrites. These complexes become immediate in water hydrolyzed to insoluble iron hydroxide or nickel hydroxide, which precipitate.

In general, the catalysts for the reaction in the second and further stages the hydrogen sulfides and the sulfides, that is, monosulfides and polysulfides of the metals of the group IA of the Periodic Table of the Elements, and are based on Alkanol solutions produced. Sodium and potassium sulfide catalysts are by far the most advantageous of which two  Potassium is particularly preferred. In the reactor of the first However, it is also possible to use rubidium in a mixture of 40% by weight. Rubidiumsulfiden, 26 wt .-% potassium sulfides and 60 wt .-% sodium sulfides, so in the form of the different types of sulphides, available. The range of relationships for the aforementioned Mixtures is 1: 1.5 to 2.5: 3.5 to 4.5, however, these mixtures must be in the for the catalyst A be prepared manner described. The typically used catalysts are in the form of hydrates used, but is a small portion of the catalyst, u. between about 15% by weight and typically less than 10% by weight or even less than 5% by weight, apparently in Form of an alkanolate ago. Alkalimetallethionate are as transition intermediates also, hydrates and alkoxides of these compounds are very complex and subject at the beginning a number of conversions depending on the prevailing reaction conditions. It seems that the presence of mixed Hydrates and alcoholates for the outstanding results is required.

If the feed of the first stages of heat resistant components, u. between asphaltenes of various molecular size, resins and oils such as sludge oils, in appreciable quantities of more than 50% by weight and, including a fraction having a boiling point higher than 583 ° C, up to 95% by weight Implementation in the first stage at least a portion of the above-defined specific catalyst, recordable example of the catalyst A required. The lower the proportion of feed to more heat resistant compounds, the less catalyst A is required. Also for the feed 12 a catalyst A is recommended. Depending on the desired molecular resolution, this catalyst can also be combined with catalysts of a different composition, since this is the best way of effecting molecular splitting without the formation of excessive amounts of gas. Another catalyst for the feed 12 a is the catalyst B, the KHS in an amount of 5 wt .-% was added (on the basis of elemental potassium), or which has been replaced, taking into account the same proportions by up to 95% of catalyst A , The suitability of the catalysts for the molecular cleavage of the heat-resistant compounds is generally as follows: Catalyst A (catalysts prepared using ethanol), Catalyst A (catalysts prepared using methanol), Catalyst A (using methanol and ethanol Catalysts), mixture of catalyst A and catalyst B (contains at least 50% of catalyst A based on the moles of potassium), catalyst D, catalyst E (with an addition of about ¹ / ₃ mol KHS), catalyst B and finally catalyst C. Of course, by increasing the amount of KHS added to the catalysts, their cracking capacity is increased.

The order of catalysts suitable for hydrogenation is the following: 0.5 part of catalyst A + KOH with added another 0.5 parts Catalyst A (very exothermic hydrogenation reaction), then Catalyst C with mixed KHS (considering the required Degree of cleavage), Catalyst B + A (50% or more on catalyst B based on the moles of potassium) and catalyst B.

Example 1

An asphaltene of the properties given below, which had a high softening point of 132.22 ° C and by Solvent extraction was obtained in the treated as follows to obtain the product A. The catalyst was the previously described catalyst A. If that when treating the asphaltenes in the first stage product obtained was reacted in a second stage, was obtained in B labeled product. The two Catalysts were used without support.  

The above data clearly show that in terms of Viscosity and specific gravity of products achievable Improvement, the significant increase in hydrogen content and the noteworthy reduction in metal content recognize from the later fractions.

The following examples show the results obtained. As a feedstock were by solvent extraction asphaltenes obtained or above and in Example 1 specified type used.

All experiments were batchwise in a stirred tank made, whose inner diameter is 15.85 cm and whose height was 25.4 cm. The stirred tank was with a Stirrer and a steam distributor equipped. In one the municipal water supply directly connected Steam boiler became steam at a pressure of 2.67 bar generated. The steam passed over an inner diameter of 9.53 mm lines with atmospheric pressure to the steam distributor. The steam distributor had a diameter of about 8.89 cm and several exit holes, which the steam after directed upwards. The distributor was located at the bottom of the reactor.

If a catalyst without carrier was used provided for the reactor a stirrer. The engine was immediate arranged above the reactor. The area, in which the agitator shaft was fed into the reactor was sealed by a gasket. The stirrer consisted Double circles through angled, curved blades were connected. The stirrer could be replaced by baskets be, which in later to be described manner Charge catalyst were charged. On the stirrer shaft were arranged four baskets loaded with supported catalysts. The baskets together with the stirrer shaft gave a total diameter of nearly 15,88 cm. The baskets were about 15.24 cm high and about 12.7 mm thick.

The top of the reactor had a gasket in which the agitator shaft turned, the riser, which distillate  discharged from the head of the reactor, and a pressure relief device with a valve, which at a pressure of more than 2.00 bar opened. This pressure relief device was also used to feed the reactor with solid feed to feed.

As a rule, two thermocouples were in the lid of the reactor used. A thermocouple measured the temperature in the lower half of the reactor and the other thermocouple measured the temperature in the upper half of the reactor.

The riser was about 22,86 cm high and possessed an inner diameter of 1.91 cm.

The second stage reactor was one 30.48 cm longer Tubular reactor with an inner diameter of about 3.81 cm. The capacity of this reactor was 347.5 cc. This Reactor was with three radiators wound around the reactor fitted. The temperature of each radiator was controlled by a thermocouple via regulators, which were arranged on a portable control panel.

The second stage reactor, from bottom to top passing gases were then passed through a wash bottle with a diameter of 40.64 cm, which is not was cooled with water.

The first wash bottle was arranged vertically and equipped at the bottom with a collecting flask of 500 ml. The collecting flask was passed through a jacket radiator normally kept at 240 ° C. The bottom of the Receiving flask was equipped with a tap to the product deducted.  

At the top of the collecting flask was a second cooler (wash bottle), the parallel to the first cooler (wash bottle) was. The second Cooler was not cooled with water. The second cooler was also a glass cooler with spherical cooling areas.

One from the second cooler in the downward direction inclined pipe resulted in a water-cooled Cooler. This water-cooled cooler was arranged vertically, about 45.72 cm long and from the top in an unheated 500 ml piston inserted at the bottom with a drain cock was equipped. From the second fitting This piston performs a water-cooled parallel and vertical Cooler way. Another water-cooled radiator was placed directly on this piston.

The top water cooled cooler was with a obliquely upwardly extending pipe of 30.38 cm equipped. This line had a diameter between 12.7 and 19.05 mm. This line leads to one with ice cooled radiator.

This iced cooler was a double-walled one Vessel as it is usually used to catch fumes their entry into a vacuum pump is used. The inner container contained a mixture of water and ice, whereas gases and vapors pass through the outer shell of the vessel streamed. The gases and vapors occurred at the bottom of the On and at the top of the vessel. The Bottom of the container had a collecting compartment with a Content of 50 ml. The collection compartment was with a drain cock equipped for draining product.

The remaining gases and vapors became another Cooler passed, similar to the iced cooler was. This cooler was made with a mixture of carbon dioxide snow and 2-propanol cooled. Again, the product became collected in a collecting compartment, which is below the  Radiator provided and equipped with a drain cock was.

The remaining gases and vapors were then treated with potassium hydroxide washed, the 6 moles of KOH dissolved in 360 ml of water contained. The amount of gas was then by passing the Gases determined by a gas meter. After this Measurement were taken at intervals samples and the other gases passed into the fume cupboard.

For the following experiments, 1300 g of solid asphaltene were used weighed in and on for loading the reactor crushed suitable particle size. Under protection against oxygen was in the previously purged example by helium Reactor is a liquid or solid catalyst brought in. About 40 g of the theoretically anhydrous catalyst A were introduced into the reactor.

The second stage reactor was under the same Precautions with about 300 ml of a Charged catalyst fed. The second stage reactor was initially heated to the water from both the zeolitic Abort carrier as well as from the catalyst.

After heating the second stage reactor at over 300 ° C, the water from the reactor room largely was removed, the reactor of the first stage was heated.

It was obtained by solvent extraction Asphaltenes with a melting point of 93.33 ° C and 204.44 ° C, respectively used.

After the temperature required to melt the column was reached in the reactor of the first stage, the stirrer was switched on. In general, the stirrer was initially operated at a speed of about 30 to 60 min ^-1 .

Steam was then fed when in the Reactor of the first stage reaches a temperature of 220 ° C.  had been. At this time, the reactor should be the second stage reaches a temperature of 424 ° C or up have settled this temperature.

Helium was in an amount of about 200 ml / min prior to introduction supplied by steam into the system via the steam distributor, to keep the distribution openings free and the system oxygen rich to keep.

example 2

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

The catalyst used was one in the following manner prepared modification of the catalyst A. To the previously described Solution of KOH became a solution of 1 mol KOH in 30 ml H₂O, whereupon the resulting solution with Hydrogen sulfide was saturated. The solution separated in two layers, of which the upper one third and the bottom two-thirds of the total fluid level accounted for. The two layers were separated and dried, whereupon the two residues obtained were mixed again. The weight of the The catalyst used was theoretically anhydrous Condition 40 g.

The second stage reactor had been charged with a zeolite-deposited catalyst during assembly of the apparatus. The second stage reactor contained about 300 g of carrier and catalyst. The catalyst was the catalyst D, ie K₂S _1.5 . To start the plant, it was necessary to heat the second stage reactor to at least 175 ° C prior to heating the first stage reactor.

The first stage reactor was then heated. In this case, only the lower half of the reactor was heated, however, the upper half of the reactor is not heated. At 220 ° C was placed in the reactor above the bottom Steam distributor introduced a small amount of water vapor.

When reaching a temperature of about 220 ° C on The bottom of the first reactor slowly and steadily set the Production of a hydrocarbon product, which in the flask below the water-cooled twin cooler was condensed. However, this product was essential heavier than that at working temperatures in the range of Product obtained from 390 to 424 ° C.

Once in the bottom of the reactor of the first stage a Temperature of 360 ° C had been achieved, one resulted considerable improvement in production speed for the product. The reaction became exothermic, so that the temperature increased rapidly and leveled off to about 415 ° C. This temperature was from this point in the ground maintained the reactor. The temperature at the head of the Reactor had reached 360 ° C.

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

If the contents of the first stage reactor with the Stirrer was in contact, the process ran in the reactor the first stage at about 415 ° C evenly, the Temperatures in the second stage reactor between 440 and 46 ° C fluctuated.

The amount of steam was about 20 ml in steam transferred water per minute. Towards the end of the experiment was left at the top and bottom of the first stage reactor temperatures rise to 440 ° C.

The above-described modification of the catalyst A provided almost no gas. The amount of gas was less than 6 liters.  The main part of the product was from a bottom product No 2 formed below the water-cooled radiator was caught. This product together with the trapped in the collection compartment cooled by water and ice Product a total of 458 g of product with one API number of 23 (density = 0.9158 at 15.56 ° C).

The bottom product # 1, which is below the air-cooled radiator weighed a total of 33 g and possessed at 15.56 ° C. a density of 0.96587 and an API of 15.

Under the cooled with dry ice and 2-propanol calibrated trap, 44 ml of bottoms were collected, however, only 28 ml could be withdrawn, as the slight Tail vaporized. The API number of this light tail was 81 at -10 ° C (density of 0.6553 at -10 ° C). Since the evaporation process is quick, which is for the Density indicated value very inaccurate.

There was an unusually light coke, the Leave in the reactor in the form of 50 mm thick layers. The volume of this coke was 1800 ml, however The weight of this coke was only 513 g.

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

The total conversion to products obtained was (1045.7 / 1300) x 100 = 80.42%. The yield was (soil products + gases) (532.7 / 1300) x 100 = 40.97%.

example 3

In this example, in a similar way as in Example 2, but with one catalyst in another Form used and the reactor of the first stage was heated up faster.

The catalyst A used for this example resulted in its production only a single catalyst layer, so that required in Example 2 above Separation of the two layers formed during drying was not required. In this experiment  a temperature of 420 ° C was maintained.

During the 42-minute trial period converted about 40 g theoretically anhydrous Catalyst after reaching the working temperature from about 420 ° C, a feed of 1300 g by solvent extraction obtained asphaltenes in the following soil products around:

• a) Bottom product no. 1, which is below the air-cooled Cooler was collected and a total of 29 g Hydrocarbons with an API number of 11.5 (d = 0.9895 at 15.56 ° C).
• b) Bottom Product # 2, which is below the water-cooled Cooler and the cooled with water and ice Cooler was collected and what 398 g of hydrocarbons with an API number of 29 (d = 0.8816 at 15.56 ° C) provided.
• c) Bottom Product No. 5, which is under dry ice and 2-propanol cooled trap in an amount of 63 ml was collected and at room temperature an API number of 83 possessed.
• d) A total of 135 liters of gas were produced. The amount of gas was behind the Alkalihydroxidwäsche the gas-steam mixture after being cooled with dry ice and 2-propanol Trap measured. These gases were not analyzed, however resulted from the average gas analysis of similar experiments, that about 5 mole% of gases that are not hydrocarbons such as hydrogen, carbon monoxide and carbon dioxide. The remaining gaseous hydrocarbons have an average molecular weight of 49, so that about 280 g of gaseous hydrocarbons were obtained.
• e) This experiment was also the case with the Experiment according to Example 2 resulting light coke. The weight of the coke was 489 g.

The products were obtained in the following quantities:

Bottom Product No. 1 | 29.0 g Bottom product no. 2 398.0 g Bottom product no. 5 41.5 g (cooled trap) gases 280.0 g coke 489.0 g a total of 1238.0 g corresponding to 95.23%.

Yield = (Soil Nos. 1, 2 and 5 + Gases / Feed) · 100 = 57.57%.

Both Example 2 and Example 3 was in the second stage reactor the supported catalyst completely clean and free of pitch and carbon.

The basic difference between the examples 2 and 3 is the significantly higher gas production at Example 3.

In both examples, the soil products separated from the water and condensed the soil products from the Steam, with no emulsion formed.

example 4

In this example, a catalyst mixture was used, which to ² / ₃ from the non-separated catalyst according to Example 2 and about ¹ / ₃ from the catalyst according to Example 3 (potassium) existed. The catalyst was without carrier in an amount of about 40 g, as a theoretical anhydrous, used.

The reactor reached one during this experiment Temperature of 420 ° C.

The products obtained in this experiment were as follows:

Bottom product no. 1 195.8 g (d = 0.9793 at 15.56 ° C or API number 13) Bottom product no. 2 309.5 g (d = 0.86 at 15.56 ° C or API number 33) Bottom product no. 5 28.0 g gas 276.4 g (133 l x 0.95 / 22.4 x 49 = 276.4 g) coke 473.0 g a total of 282.7 gg

Conversion = (1282.7 / 1300) x 100 = 98.67%.
Yield (on bottoms Nos. 1, 2 and 5 and gases) = (809.7 g / 1300 g) x 100 = 62.28%.

The coke was created in the room below the stirrer and a Part of this room was also under the steam distributor.

example 5

The essential difference between this example and the previous examples was the use of a Supported catalyst in place of the added before the start of the experiment strapless catalyst. The catalyst was in four baskets on a sieve made of sintered stainless steel Steel supported. The stainless steel was 3.175 mm thick and was cut into strips of 3.175 mm, the in turn divided transversely to 4,763 mm pieces were. The dimensions of the carrier were therefore 3.175 mm · 3.175 mm · 4.763 mm. The catalyst used was the same as in Example 2.

The supported catalyst was in baskets of dimensions 6.35 mm x 63.5 mm x 152.4 mm, with four baskets were connected to the stirring shaft, so that the baskets at Rotation of the agitator shaft formed the stirrer. The supported catalyst was held in place by a sieve, and the wire mesh baskets were supported by a frame.

When using the same catalyst as in Example 2 has been reduced by using much less Catalyst on the support the amount of gas produced is reduced from 135 liters to 90 liters. The largest part of this gas quantity originated towards the end of the Try as the temperatures rose to 440 ° C.

The speed of the agitator shaft carrying the baskets was initially 60 min ^-1 and later increased to 120 min ^-1 during the experiment.

gas 187.0 g (90 l of gas 0.95 / 22.4 x 49 = 187.0 g) Bottom product no. 1 407.0 g Bottom product no. 2 368.0 g Bottom product no. 5 43.0 g coke 199.5 g a total of 1204.5 g

The obtained from by solvent extraction Asphaltene existing feed weighed 1290 g.

Conversion = (1204.5 g / 1290 g) x 100 = 93.%
Yield = [(bottom products Nos. 1, 2 and 5 + gas) / 1290] × 100 = 77.90%.

The API number of the combined soil products No. 2 and No. 5 was 32.0 at 15.56 ° C. The bottom product No. 1 was in a liquid portion at 200 ° C and a non-liquid portion at 200 ° C separated. The liquid fraction had a calculated API number of 5 and weighed 228 g, accounting for 56.03% of the total 407 g of bottoms no. The rest was obvious from some upgraded forms of solvent extraction formed asphaltenes. One up Reason for the liquid content of the soil products at 200 ° C No. 1, gave a new calculation Yield or a degree of conversion of 64%.

With all these calculations of yield and eligibility was based on the assumption that 95% of the gas from a hydrocarbon with one average molecular weight of 49 was formed.

example 6

In the above described procedure according to example 2 became a load from a mixture of different composition Residues used.

Analysis of the feed:

Fraction collected at 537.78 ° C | 8.6% API number 6.6 Density at 15.56 ° C 1.0243 Sulfur,% by weight 3.91 Nitrogen, wt% 0,478 Hydrogen, wt.% 10.35 Carbon, wt% 84.72 Oxygen,% by weight 0.471 Oils,% by weight 29.6 Resins,% by weight 57.8 Asphaltenes, wt% 12.6 Ramsbottom carbon, wt% 21%

metals:

Vanadium | 228 ppm nickel 52 ppm iron 14 ppm sodium 4 ppm viscosity 1430 cSt at 27.78 ° C 250 cSt at 121.11 ° C floating point 54.44 ° C

The feed was gradually heated to as long as possible cause no coking than the mixture not yet pumpable or not yet the reaction temperature reached about 425 ° C. At this Temperature was the reaction in the presence of the catalyst and carried out by steam. The loading away a total of 15 650 g.

The boiling point defined by ASTM D-1160 and D-1180 the residue from existing feed was at a Sample of 200 ml of the feed according to the mentioned Standard determined. The data obtained were as follows:  

This product was converted by means of a supported catalyst in a stirred reactor in which the catalyst was placed in baskets. The support was a zeolite and was in the form of extrusions of 3.175 mm. The catalyst had the empirical formula K₂S _1.5 and was the catalyst D prepared above, to which 5 wt .-% of the catalyst A had been added. The amount of catalyst was 20.16 g per 220 ml of the described carrier.

The initial conversion of the feed to the first product (top product from the reactor 110) 71% (the combined products without the very volatile and in the cooled with dry ice and isopropanol Trap trapped materials, gases and  those ingredients of the feed that are at a temperature could be distilled off under 565.56 ° C). The top products were with the separated soil products with 98.1% of the original Charge calculated, with the remaining amount obviously gases and volatile distillates were.

A 7 g H₂O containing sample to 1392 g of the overheads from the reactors 110 and 117 (which, however, did not contain the above-mentioned condensates) had the following boiling analysis:

Initial boiling point to 182.22 ° C: 71.6 g or 5.24 wt%

182.22 to 342.33 ° C 203.4 g or 14.89 wt% 343.33 to 537.78 ° C 1084.1 g or 79.36% by weight

The product from the reactor 110 and 117 had a boiling point at 182.22 ° C and had the following properties:

API number 54.5 Density at 15.56 ° C .7608 viscosity 0.72 cSt at 37.78 ° C 0.568 cSt at 65.56 ° C ash content 0.009% by weight nitrogen content 0.05% by weight sulfur content 1.0% by weight

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

API number | 34.2 density 0.8540 at 15.56 ° C water 0.002% by weight floating point -28.89 ° C viscosity 2.57 cSt at 37.78 ° C 1.61 cSt at 71.11 ° C sulfur 2.18% by weight ash 0.01% by weight

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

API number | 17.8 density 0.9478 at 15.56 ° C floating point 46.11 ° C viscosity 37.34 cSt at 65.56 ° C 12.05 cSt at 98.89 ° C Carbon content 85.03% by weight Hydrogen content 11.55% by weight nitrogen content 0.26% by weight Conradson Carbon 2.56% by weight

The obtained at temperatures above 565.56 ° C. Residue contained:

Vanadium | 5.6 ppm iron 48.6 ppm sodium 8.2 ppm potassium 22.3 ppm nickel 0.48 ppm

The properties of the fraction with a boiling point of 343.33 ° C and above, were the following:

Initial boiling point at 344.44 ° C and the following yields:

10% to 388.89 ° C
20% to 416.67 ° C
30% to 453.33 ° C
40% to 470 ° C
50% to 478.89 ° C
60% to 497.78 ° C
70% to 508.89 ° C
80% to 524.44 ° C
90% to 547.78 ° C
Boiling point 565.56 ° C
at a yield of 94.4%.

The K factor of the overhead fraction above 343.33 ° C from reactor 110 (final stage hydrogenation) was 11.71.

The bottom products recovered from the first stage reactor 110 at 212 a were returned in batches when a sufficient amount had been weighed. The catalyst for the feed 212 a was the Stahlsiebteilchen on the dimensions of 3.175 mm x 6.35 mm x 4.76 mm deposited Catalyst A and contained in a total volume of 200 ml 15.7 g catalyst. A total of 1286 g of sample were used. The total yield was 750.7 g (including 8 g of dry ice product without gas). About 29% of the feed was recovered as bottoms. The analysis of the feed 112 a (bottoms) was as follows:

carbon 84.78% by weight hydrogen 8.74% by weight ash 0.98% by weight sulfur 4.64% by weight Conradson Carbon 40.57% by weight

About 48.4% of the feed was converted.

The united won products possessed following properties:

carbon 82.24% by weight hydrogen 11.95% by weight oxygen 0.64% by weight water 2.8% by weight

They have a boiling range of 182,22 to 343,33 ° C Fraction had the following properties:

API number | 27.2 floating point -26.11 ° C viscosity 8.533 cSt at 37.78 ° C 4.028 cSt at 65.56 ° C nitrogen 0.17% by weight water 2.8% by weight

The boiling range of 343.33 ° C until the end of boiling owning fraction possessed the following characteristics:  

API number | 17.4 density .9444 floating point 35 ° C viscosity 21.54 cSt at 65.56 ° C 7.989 cSt at 98.89 ° C carbon 84.7% by weight hydrogen 11.27% by weight oxygen 0.66% by weight Conradson Carbon 2.2% by weight nitrogen 0.27% by weight vanadium 3.5 ppm iron 7.6 ppm sodium 1.6 ppm potassium 10.3 ppm nickel 0.45 ppm

distillation residue

carbon 85.22% by weight hydrogen 11.06% by weight oxygen 0.49% by weight sodium 3.07 ppm vanadium not determined nickel 0.344 ppm iron 4.6 ppm potassium 27.4 ppm

The product passing above 343.33 ° C (from feed 112a ) had the following boiling analyzes for a sample of 191 g, of which 163 g was withdrawn overhead: it was a 175 ml sample at 15.56 ° C. The collected fractions were as follows:

10% to 343.3 ° C
20% to 412.78 ° C
30% to 443.33 ° C
40% to 465.56 ° C
50% to 486.11 ° C
-
-
80% to 554.44 ° C
85% to 565.56 ° C

This distillation product had at 15.56 ° C a API number of 20.0, and the residue product had one API number of 16.4. The low metal content and the pretty high API number makes this residue a good one Feed, for example for catalytic cracking.

The residue constituting about 15% formed the bottom product in the reaction vessel 130 . His analysis was as follows:

carbon 83.99% by weight hydrogen 6.88% by weight oxygen 2.9% by weight water 0.05% by weight

Based on the 58.37% conversion (product recovered divided by the weight of the 83% feed) (reflux of the bottom products from reactor 110 , feed 112 a) and the overheads of reactor 110 , the total conversion was Liquids in the last dry ice trap and the gases withdrawn with the top products of the reactor 110 as well as the presumably distillable products (which are presumed to still be in the feed and distill below 565.56 ° C) were 87.9%, but the actual yield was still higher if the gases and products trapped in the dry ice trap are charged with an average molecular weight of 49.

The yields were based on the total amount of starting material used based.

Claims (6)

A process for converting petroleum or petroleum residues into lower viscosity products and / or more hydrogenated products in the presence of an alkali sulfide catalyst without the addition of gaseous hydrogen, characterized in that in a first reaction zone in the presence of water vapor and an additional catalyst, the petroleum or oil Petroleum residues at a temperature of 160 to 650 ° C and a pressure of up to 10 bar, wherein the additional catalyst from at least 5 wt .-%, based on the total amount of catalyst, a saturated with hydrogen sulfide solution of 0.5 to 2 moles of alkali metal hydroxide in one mole of methanol, ethanol, 1-propanol, 1-butanol or mixtures thereof, optionally containing water, separating the reaction product withdrawn from the first reaction zone into a top fraction containing vapor and gaseous products and the bottoms fraction either back to the first reaction zone kführ or in the presence of water vapor and the above two catalysts at the above or lower temperature and the above pressure. 2. The method according to claim 1, characterized in that one the top fraction of the reaction product of the first reaction zone directly or after cooling in at least one another reaction zone in the presence of water vapor and a supported catalyst with alkali hydrogen sulfide, sulfide, polysulfide, sulfide hydrate and / or polysulfide hydrate Reduction of bromine number hydrogenates. 3. The method according to claim 1 or 2, characterized that the bottom fraction of the reaction product of the first Reaction zone at a temperature of 360 to 440 ° C converts. 4. The method according to claim 1, characterized in that one in the first reaction zone and in the reaction of the bottom fraction the reaction product of the first reaction zone works adiabatically at a temperature up to 560 ° C. 5. The method according to any one of claims 1 to 4, characterized that in the first reaction zone at a Temperature of 320 to 450 ° C implemented. 6. The method according to any one of claims 1 to 5, characterized in the presence of at least 27% by weight Steam, based on the feed rate, works.