BR0003412B1 - Heavy oil oil hydro-cracking process. - Google Patents

Description

"HEAVY OIL HYDROCRACKING PROCESS" History of the Invention Field of the Invention

The present invention relates to a petroleum heavy oil hydro-cracking process and more specifically to heavy metal oil containing heavy oil cracking. In particular, the present invention relates to a heavy petroleum oil hydrocracking process containing heavy metals (such as atmospheric residue and vacuum residue), thereby producing light oil.

Increasing attention is being given to cracking technology to produce light oil, which is byproduced, from heavy oil, which is overproduced. This technology is becoming more and more important as oil reserves are inevitably running out.

Many processes have been proposed for thermal cracking or heavy oil hydro cracking, however, they present problems related to the conversion of heavy oil (such as vacuum residue) to light oil. This will be fully explained below.

A heavy oil, such as vacuum residue, tends to contain a large amount of nitrogen and sulfur compounds. In addition, it also contains a large amount of organometallic impurities that are extremely harmful to cracking in the presence of decatalizers. Such metallic impurities include mainly those containing nickel (Ni) and vanadium (V) and other metals less frequently. They chemically combine with comparatively polymeric organic compounds such as asphaltene in olives and considerably inhibit catalytic activity for decomposition and removal of nitrogen, sulfur and oxygen compounds and hydrocracking of heavy organic matter. There is a known process for thermal cracking of heavy oil such as waste vacuum, without using catalyst. This is the so-called decoking process. This process has the disadvantage of producing a large amount of coke as a by-product, since disposal of it is a problem, and also producing a large amount of gas split to over cracking, therefore decreases the distilled production. Another disadvantage is that the resulting distillers are of poor quality due to a high content of olefinic and aromatic components. There is a fixed bed type hydro-cracking process which is performed in a granular catalyst-filled reactor. The disadvantage of this process is that the catalyst is vulnerable to asphaltene and heavy metals, such as V and Ni, in heavy oil when excess decay occurs. Thus the catalyst decreases in activity and the catalyst layer becomes blocked due to the gradual deposition of coke as a byproduct and heavy metals. This prevents prolonged continuous operations.

There is another hydro-cracking process which is performed in an ebullated bed reactor using an extruded granular catalyst. This process causes such stirring in the reactor so that there is no problem with pressure increase due to coke accumulation. Additionally, it allows the catalyst to be loaded and unloaded during operation and thus allows for continued continuous operations while maintaining catalyst activity at a certain level. Its advantage over fixed bed hydro-cracking is reduced by the fact that it requires the catalyst to be recirculated during operation, thereby causing mechanical damage to the pumps, etc., making it difficult to operate. This process still has other disadvantages, such as requiring cost-effective catalyst, requiring high reaction pressure (usually 150-200 kgf / cm2), and resulting in not sufficiently desulphurized and denitrified products.

The present invention has been developed to solve the above mentioned problems. It is an object of the present invention to provide a process for producing light oil by hydrocracking from heavy oil from oil, such as vacuum residue, containing heavy metals. The hydrocracking process of the present invention is characterized by its condition of operating continuously for a long time without coke accumulation and economically obtaining a high production of high decayed yeastsulfurized and denitrified oil.

The aforementioned object of the present invention is achieved by means of a heavy petroleum oil cracking process which is defined in the appended claims. The process is shown below. The present invention is directed to a heavy metal hydrocarbon cracking process containing heavy metals, said process comprising steps (a) to (h).

(a) a first reaction step for feeding a suspended bed reactor with a feedstock paste composed of heavy oil containing heavy metals, limonite iron ore with catalyst and sulfur as cocatalyst, together with hydrogen gas, hence performing documental hydrocracking heavy oil under a reaction pressure of 30-160kgf 'cm2 (2.9-15.7 MPa), a reaction temperature of 430-455 ° C, and a reaction time of 30-180 minutes.

(b) a first gas-liquid separation step to separate the reactor effluent, obtained in the first reaction step (a), for a fasegasous fraction and for a liquid phase fraction.

(c) a second reaction step for feeding a fixed-bed reactor filled with Ni-Mo or Co-Mo catalyst with gas phase fraction obtained by separation from said first gas-liquid separation step, thereby performing hydrotreatment under a pressure 30-160kgf / cm2, at a reaction temperature of 310-380 ° C (at the beginning of the operation) and a liquid spatial hour velocity of 0.3-2 hour "1;

(d) a second gas-liquid separation step to separate the reactor effluent obtained in said second reaction step (c), for a fasegasous fraction and for a liquid phase fraction, (e) a distillation step to separate by distillation of the liquid phase fraction obtained by separation from said second gas-liquid separation step (d) for said prescribed fractions;

(f) a flash separation step for performing flash separation at a pressure lower than that of the first gas-liquid separation step (b) in the liquid phase fraction, which was obtained by separation from said first gas-liquid separation step ( b) so that the light reaction product in the liquid phase fraction vaporizes into the gas phase;

(g) a solid-liquid separation step for feeding a solid-liquid precipitation type separator shares the liquid phase fraction containing the medium products, medium to heavy products, heavy products, waste, heavy metals and a catalyst used in said first stage. reaction (a), which were obtained by separating from light product in said flash separation step (f), and a solvent for solid-liquid separation, hence separating them into the solid eliquid components (both at normal temperature), said solid components. containing said residue, catalyst used in said first reaction step (a), and heavy metals, said liquid component containing said medium products, medium to heavy products, and heavy products; (h) a recycling step to return to said first reaction step (a) the liquid component obtained by separation in said solid-liquid separation step (g) and the phase fraction liquid containing medium products, medium to heavy products, heavy products, residue, heavy metals and a catalyst used in said first reaction step (a) which were obtained by separation from the light product in said flash separation step (f), as the remainder left after feeding into said precipitated solid-liquid separator.

Brief Description of the Drawings

Figure 1 is a schematic diagram showing an exemplary oil heavy oil hydrocracking process according to the present invention.

Figure 2 is a schematic diagram showing another example of the heavy oil oil hydrocracking process according to the present invention.

Detailed Description of the Invention

The present invention is practiced as follows as an example.

The process of the present invention uses, as feed material, petroleum heavy oil, such as vacuum residue, containing heavy metals. This feedstock is incorporated with catalyst and sulfur ferrous limonite ore as cocatalyst, and is used in the form of paste.

The feedstock slurry is fed together with hydrogen gas to a suspended bed reactor and is subjected to a first hydrocarbon cracking step (a) of said heavy oil under a reaction pressure of 30-160kgf / cm2 at a reaction temperature. 43 0-45 5 ° C and in a reaction time of 30 to 180 minutes. The reactor effluent [1] containing hydro-cracking products obtained in said first stage is subjected to a first gas-liquid separation step (b) for gas phase fraction separation [2] containing gaseous products and unreacted hydrogen and a fraction of faseliquid [3] containing liquid reaction products and the catalyst used in said first reaction step (a) and heavy metals. The gas phase fraction [2] obtained by separation in said first gas-liquid separation step (b) is fed to the Ni-Mo or Co-Mo catalyst fixed bed reactor. It undergoes a second hydrotreation reaction step (c) which is performed under a reaction pressure of 30-160 kgf / cm2, a reaction temperature of 310-380 ° C (at the beginning of operation) and a spatial velocity. 0.3-2 hour "1. The reactor effluent [4] containing the hydrotreating products obtained in the second reaction step is subjected to a second gas-liquid separation step (d) for separation to a fraction phase. [5] containing unreacted hydrogen gas and for a liquid phase fraction [6] The liquid phase fraction [6] obtained by separation in said second gas-liquid separation step (d) passes through a distillation step (e) for separation for the prescribed fractions.The liquid phase fraction [3] obtained by separation in said first gas-liquid separation step (b) undergoes a flash separation step (f) for flash separation under a pressure lower than that given above. gas-liquid separation step so that the light product in the liquid phase fraction vaporizes into the gas phase. The flash separation step (f) produces a liquid phase fraction [7] containing a medium products, a medium to heavy and a heavy products, residue, heavy metals and a catalyst used in said first reaction step (a) which have been separated. Light product. Part [7A] of the liquid phase fraction [7] is fed to a solid-liquid precipitating type separator together with a solid-liquid separating solvent. This solid-liquid separator performs a solid-liquid separation step (g) for separation for a solid component [8] and for a liquid component [9] (both at normal temperature), said solid component containing said residue, the catalyst used in the cited the first reaction step (a) and heavy metals. The said liquid component containing said medium products, medium heavy products and heavy products.Finally a recycling step (h) is performed to return to said first reaction step (a) the liquid component [9] which was obtained by unsaturated separation. solid-liquid separation step (g), and the remainder of said liquid phase fraction [7] (liquid phase fraction [7B] other than the faseliquid fraction [7A] supplied for said solid-liquid precipitation type separator) , ([7B] = all of [7] - [7A]).

The above is the manner in which the hydro-cracking process of the present invention is carried out in heavy petroleum oil. The distillation step (e) results in a prescribed fraction (light oil) continuously intermittently.

The present invention produces its effects as explained below.

In accordance with the present invention, the petroleum heavy oil hydrocracking process begins with a first reaction step (a), in which the heavy oil from oil incorporated with the catalyst iron limonite ore undergoes hydrocracking in a delighted reactor. suspended. Limonite iron ore is quite poor in catalytic activity for desulphurization and nitrification but has a high hydro-cracking catalyst activity. In addition, it is a low cost catalyst available. The entire process is operated in such a way that the intermediate products in each step are returned to the first reaction step.

(a) by the recycling step (h). Intermediate products include the liquid component [9] containing heavy products which are obtained in the first production step (a) and its subsequent steps from (b) to (f), and the liquid phase fraction [7] containing said weighted products as well as catalyst used. The recycled liquid component [9] and the liquid phase fraction [7] pass through the first reaction step (a) so that the heavy products contained therein are subjected to hydro cracking. This hydro cracking is facilitated by the remaining catalyst used in the liquid phase fraction [7]. (catalyst still used for hydro cracking and thus works effectively in the first reaction step (a)). With respect to the first reaction step (a) and the recycle step (h), the process of the present invention allows a large production of light oil without coke accumulation. In addition, the process of the present invention is economical because the first reaction step employs a catalyst of The first gas-liquid separation step is designed to separate the reactor effluent [1] containing hydrotreating products, which was obtained in the first reaction step (a) for the gas phase fraction. [2] containing gaseous reaction products and unreacted hydrogen and for the liquid phase fraction [3] containing liquid reaction products and said catalyst (equation used in said first reaction step (a)) and heavy metals. In the aforementioned second reaction step (c) below, the gas phase fraction [2] thus separated goes through hydro-treatment in the Ni-Mo or Co-Mo filled fixed bed reactor. It should be noted that the gas phase fraction [2] fed to the second stripping step (c) does not contain asphaltene in the first stripping step (a) and almost all asphaltene, together with heavy metals are present in the separated liquid phase fraction [3]. in the first gas-liquid separation step (b). Therefore, the gas phase fraction [2] undergoes hydro-treatment in the fixed bed reactor in the absence of heavy metals and asphaltene, which disables the desulfurization and denitrification catalyst and increases the coke that makes deactivation of catalyst and catalyst layer blocking. The result is the prevention of coke accumulation and decatalyst contamination and the extension of catalyst life and prolonged continuous operation. A high degree of sulfurization and denitrification and hydro cracking occurs in the second reaction step. The second stripping step (c) is followed by the second gas-liquid separating step (d), the distillation step (e), the flash separating step (f), the solid-liquid separation step (g) and the (c) as mentioned above. The first gas-liquid separation step (b) and the second reaction step (c) are primarily responsible for preventing coke accumulation and prolonged continuous operation. Therefore, the entire process allows for continuous production of highly desulphurized and deenitrified light oil.

In summary, according to the present invention, the petroleum heavy oil hydrocracking process performs the petroleum heavy oil hydrocracking (such as vacuum residue) containing economically heavy metals in high production volumes in a continuous continuous operation without accumulation. Coke Thus the process allows the production of highly cracked, highly desulfurized light oil with high cracking. A detailed description is given below of the effect produced by the present invention.

The first reaction step (a) employs a hydro cracking catalyst. Basically, this catalyst must be highly active (for efficient water cracking) and readily available at low cost. Low-cost catalysts include iron-based catalysts such as iron sulfide, iron oxide and red mud. Unfortunately, they have the disadvantage of having little catalytic activity. The present invention studies the effect of iron based catalysts on the hydrocracking of heavy petroleum oil. Therefore, it has been found that iron limonite ore shows a high catalytic activity when used in heavy petroleum oil. This discovery led the inventors of the present invention to employ iron oreolimonite as a catalyst for hydro cracking in the first reaction step (a). In the first reaction step (a), limonite iron ore (as a heavy oil hydro-cracking catalyst containing heavy metals containing oil) is used in combination with sulfur as a cocatalyst. The heavy oil incorporated with catalyst and cocatalyst takes the form of slurry. This slurry feed material together with hydrogen gas is fed into a suspended bed reactor. Heavy oil hydro-cracking is performed under a reaction pressure of 30-160 kgf / cm2, the reaction temperature of 430-455 ° C, and a reaction time of 30 to 180 minutes.

The first reaction step (a) results in the reactor effluent [Ι] which contains hydro-treatment products. This effluent [1] is fed to the first gas-liquid separation step (b) for separation to the gas phase fraction [2] containing unreacted hydrogen gas reaction products and to the liquid phase fraction [3] containing liquid products and the catalyst (used in the first reaction step (a)) and heavy metals.The first gas-liquid separation step (b) is followed by the second reaction step (c) in which the gas phase fraction [2] is separated from the liquid phase fraction [3] is fed to the Ni-Mo or Co-Mo filled fixed bed reactor. Hydro-treatment is performed under a reaction pressure of 30-160 kgf / cm2, a reaction temperature of 310-380 ° C (at the beginning of the operation) and a liquid hourly spatial velocity of 0.3-2 hour "1. The gas phase fraction [2] is composed of hydrogen sulfide gas (H2S), hydrocarbon gas (C1-8), light product (C5-171 ° C), medium products (171-343 ° C), medium to (343-450 ° C) but containing very little heavy metals and asphaltene adversely affect cataitic activity in the second reaction step (c) This can be explained as follows: The first reaction step (a) includes decomposition of those compounds formed by heavy metals with heavy organic compounds The heavy metals released by decomposition are adsorbed by the catalyst and thus separated from the organic compounds or heavy oil hydro-cracking products (obtained in the first reaction step (a)). The heavy weight in the heavy petroleum oil (as feed material) is contained together with the catalyst in the liquid phase fraction [3] in the first gas-liquid separation step (b). The phaseliquid fraction [3] is eventually discharged to the downstream step. Additionally, the first reaction step (a) involves the decomposition of asphaltene contained abundantly in the heavy oil. Thus, the effluent from the first reaction step (a) decreases the deasfalten content. The first gas-liquid separation step (b), asphaltene (which has a high boiling point) does not vaporize the gas phase but for the most part remains with heavy metals in the liquid phase. It follows, therefore, that the reaction gas product in the gas phase fraction, which has entered the second reaction step (c), undergoes hydro-cracking in the heavy metal fixed bed reactor (which deactivates the sulfurization and denitrification catalyst) and asphaltene (which causes coke accumulation, catalyst deactivation and catalyst layer blocking).

In addition, the first reaction step (a) involves alkylation which reduces the steric hindrance to alkaline substituent-active catalyst sites. Thus the second reaction step (c) allows Ni-Mo or Co-Mo catalysts to show complete catalytic activity for sulfurization and denitrification.

The aforementioned practice produces the effect of preventing coke accumulation and catalytic contamination by extending catalyst life and permitting stable prolonged continuous operation. The result is the production of pure hydro-treated oil by high degree of desulphurization and denitrification. The hydrotreated oil is accompanied by hydrocarbon gas, hydrogen sulfide gas, and ammonium gases resulting from the hydrotreatment.

The second reaction step (c) results in a reactor effluent [4] containing hydrotreated products. This effluent enters the second gas-liquid separation step (d) for separation for the gas phase fraction [5] containing unreacted hydrogen and for the liquid phase fraction [6].

The liquid phase fraction thus obtained [6] enters the distillation step (e) for separation into the fractions prescribed by distillation. Fractions include naphtha (light product, Cs-171 ° C), kerosene and gaseous oil (medium products, 171-343 ° C) and vacuum gas oil (medium to heavy duty products, 343 ° C).

The liquid phase fraction [3] obtained by separation in the first gas-liquid separation step (b) enters the flash separation step (f), which causes the product to evaporate from the liquid phase fraction [3] to gas phase.

As mentioned above, the first reaction step (a) employs limonite iron ore or a low cost available catalyst. After the separated tersed light product to the gas phase [7] in the flash separation step (f) connected to the liquid phase line from the first gas-liquid step (b), the catalyst is separated from the liquid phase and discharged, along with waste and heavy metals from the system in the solid-liquid separation step (g).

Separation of the catalyst (solid) from the liquid is usually accomplished by distillation, centrifugation or filtration. The disadvantage of distillation is that there is a large oil loss resulting from the fact that the solid-liquid mixture is discharged from the bottom of the distillation column and thus a large amount of liquid accompanied by solids is discharged from the system. The disadvantage of centrifugation or filtration is the high cost of the equipment and low efficiency for fine particle-containing fluids (tenths of a micrometer maximum) as in the present invention.

Accordingly, the present invention employs the solid-liquid separation (g) step of the settling type in which a solvent is used for solid-liquid separation. A detailed description is shown below.

As mentioned above, the flash separation step (f) provides the liquid phase fraction [71] separated from the light product. The liquid phase fraction [7] contains medium products, medium products for heavy and heavy products, residue, catalyst (used in the first step (a)), and heavy metals (contained in oil heavy oil as raw material). If this liquid fraction [7] is to be separated into liquid and solid by a solid-liquid precipitation type separator, precise separation of a large size equipment because the liquid phase fraction [7] contains fine solid particles having a precipitation slow. This problem is solved by supplying the solid-liquid precipitation type separator with a solvent that has a lower boiling point than the faseliquid fraction [7] and partially dissolves the heavy and waste products. The remaining undissolved substance consists of pure solids and pseudo-solids with intermediate solid and liquid properties. The latter acts as a binder for inorganic solids (such as catalyst, heavy metals adsorbed to catalyst and coke), hence promoting flocculation. As a result, solid particles grow and settle in a brief period of time. The resulting precipitators (in the form of sludge) are discharged from the bottom part of the precipitating solid-liquid separator. The solvent is recovered and recycled. The liquid is discharged from the top of the solid-liquid precipitation type separator. The discharged liquid contains very little inorganic solids, and its solvents are recovered by flash separation. The recovered solvent is recycled to the liquid-solid separation step (g) and the remainder is recycled to the first reaction step (a).

In other words, what happens in the solid-liquid separation step (g) is as follows. The precipitation type separator is fed with the liquid phase fraction [7A] part [7] from the flash (f) separation step which has been separated from light products. This liquid phase fraction [7] contains medium products, medium to heavy products, heavy products, catalyst waste (used in the first reaction step (a)) and heavy metals. Part [7A] of the liquid phase fraction is fed together with a solid-liquid separation solvent. The separator separates it into two components, the first being a solid component [8] (at normal temperature) containing the residues, the catalyst (used in the first reaction step (a)) and heavy metals, and the second being a liquid component [9]. ] containing medium products, medium to heavy products, and heavy products. The liquid component [9] obtained by separation in the solid-liquid separation step (g) is recycled to the first reaction step (a). The remainder * of the phaseliquid fraction [7] is also recycled to the first spinning step (a) in the recycling step (h).

* (The remainder is defined as the liquid phase fraction [7B] other than the liquid phase fraction [7A] supplied for the solid-liquid precipitation type separator).

([7B] = all of [7]) - [7A])

What is recycled to the first reaction step (a) consists of the liquid component [9] and the remainder of the liquid phase fraction [7]. They are mostly heavy products (or heavy distillates) having a boiling point of 450-525 ° C under normal pressure. They readily undergo hydro cracking in the first reaction step (a). The remainder of the phaseliquid fraction [7] contains the catalyst used in addition to the heavy products and this used catalyst still has catalyst activity for hydro cracking. Therefore, this catalyst used effectively works in the first reaction step (a) enabling the reduction of the amount of catalyst to be replaced.

The above is the reason for light oil production in large volumes. The low cost catalyst and its recycling for the first reaction step (a) contributes to production economy.

In conclusion, according to the present invention, the petroleum heavy oil hydrocracking process performs the petroleum heavy oil hydrocracking, such as vacuum residue, containing economically heavy metals in large production volumes for an extended continuous operation without coke accumulation. Thus the process allows the production of highly cracked, highly desulfurized light oil with high cracking. According to the present invention, the heavy petroleum oil hydrocracking process involves the first reaction step (a) employing a suspended delight reactor in which The reaction is carried out under a reaction pressure of 30-160 kgf'cm2, at a temperature of 430-455 ° C, and a reaction time of 30-180 minutes. The reason for this will be explained below.The reaction pressure should be 30-160 kgf / cm2, because of the reaction under a pressure of less than 30 kgf / cm2 [9,80665x10 ^ 4 Pa / (kgf / cm2) x 60 Kg / cm2 = 2.9x106 Pa- 2.9 MPa] will allow coke formation in large quantities due to low partial hydrogen pressure and, on the other hand, the reaction under a pressure greater than 160 kgf / cm2 makes the production cost high. without any additional effect. The reaction temperature should be 430-455 ° C because the reaction at a temperature below 430 ° C does not promote hydro cracking but presents difficulties to the production of light oil in large volumes and the reaction at a temperature above 455 ° C causes pyrolysis. intense and increase the rate of polycondensation, hence greatly increasing coke formation. The reaction time should be 30-180 minutes because the reaction in less than 30 minutes presents difficulties in producing large volumes of oil and the reaction in times longer than 180 minutes causes excessive cracking, resulting in a large amount of hydrocarbon gas, resulting in that the amount of light oil decreases and the amount of hydrogen consumption increases more than necessary, consequently increasing the production cost. The reaction in the suspended bed reactor should preferably be performed under a reaction pressure of 50 to 100 kgf / cm2 at a reaction temperature of 440 to 450 ° C, and a reaction time of 60 to 120 minutes. The reaction under these conditions maintains a low coke formation and economically yields a large production of light oil. The second reaction step (c) employs a fixed-flow reactor in which the reaction is performed under a tapping pressure of 30 to 160 kgf / cm2, at a reaction temperature of 310 ° to 380 ° C (at the beginning of the operation) and a liquid hourly space velocity of 0.3-2.0 hours. "1. The reason for this will be given below.

The reaction pressure should be 30-160 kgf / cm2, which is identical to that of the first reaction step (a). In other words, the liquid phase fraction from the first gas-liquid separation step (b) is not pressurized for economic reasons. Reaction at this pressure prevents coke formation and allows prolonged continuous operation. The reaction temperature should be 310 ° to 380 ° C (at the beginning of operation) because sandblasting at a temperature below 310 ° C is slow and incomplete for desulphurization and denitrification and hydro-cracking, and a reaction at temperatures greater than 380 ° C. C produces more hydrocarbon gas (with less oil) due to excess hydrocracking. The hourly liquid space velocity should be 0.3-2 hour "1 because a speed less than 0.3 hour" 1 does not increase desulphurization and denitrification reactivity (resulting in a high production cost) and a speed greater than 2 hour "J results in inefficient desulphurization and denitrification.

The reaction in the fixed bed reactor should preferably be performed under a reaction pressure of 50 to 100 kgf / cm2, a reaction temperature of 330 ° - 360 ° C and a net hourly liquid velocity of 0.5-1hour-1. . The reaction under these conditions is more efficient and economical for prolonged continuous operation and allows the production of highly desulphurized and edenified light oil.

In the first reaction step (a), the limonite iron ore as catalyst should be added in an amount of 0.3-2% by weight (in terms of iron) based on the amount of heavy petroleum oil. With a quantity of less than 0.3 mass%, the catalyst tends to produce more asphaltene and coke. With an amount greater than 2% by mass, the catalyst does not work effectively to produce oil having a boiling point below 450 ° C. Sulfur as cocatalyst must be added in order to convert iron limonite ore into pyrrhotite which shows catalyst activity. The amount of sulfur should be greater than that of iron in limonite iron ore (atomic ratio), as the upper limit being three times for economic reasons. The limonite iron ore used as catalyst in the first reaction step (a) should preferably be a fine powder form having an average particle size of less than 2 µm. A catalyst having a large medium particle size has a small effective surface necessary for efficient contact with heavy oil and thus has little catalyst activity. A catalyst having an average particle size of less than 2 μm has a large effective surface area and therefore high catalyst activity. For this reason the average catalyst particle size should preferably be less than 1 μπι.

Limonite iron ore in the form of fine powder having an average particle size of less than 2 µm can be obtained by mechanical spraying of limonite iron ore.

This spraying should preferably be performed by wet process (in petroleum solvent) rather than dry process. The dry process powder (using pneumatic blender) coagulates remarkably in the heavy oil-containing feedstock paste and thus has poor dispersibility while the wet process powder (in petroleum solvent) does not coagulate in the feedstock paste, having boadispersibility, which allows a high catalytic activity.

Iron limonite ore is known to contain iron oxyhydroxide and iron oxide by X-ray powder diffraction. It varies in composition according to its origin (location and mine lot). The inventors of the present invention investigated the effect of composition on catalytic activity. It has been found that ferrous limonite ore substantially free of iron oxide has the highest catalytic activity. It is therefore desirable to use a substantially iron oxide free limonite ore in the first reaction step (a).

The following explains why ferrous limonite ore substantially free of iron oxide has the highest catalyst activity.

In general, any iron compound is sulfurized with sulfur or sulfur compounds to form iron sulfides called pyrrohotites (Fej.xS), which are species presenting catalyst activity. If conversion to pyrrohotite occurs at low temperatures, this means that active species exist before the hydro-cracking feed material begins to decompose. As a result, pyrolytic radicals are sufficiently supplied with hydrogen. This prevents the polymerization of pyrolytic radicals and promotes hydro-cracking (which means high catalytic activity).

The conversion of α-iron oxyhydroxide to pyrrhotite occurs at a lower temperature than Cl-iron oxide. Therefore, iron oxide-substantially limonite ore free of iron has the highest catalytic activity. Incidentally, the conversion temperature for pyrrhotite is 200 ° C for α-iron oxyhydroxide and 350 ° C for α-iron oxide. In the case of piritanatural (FeS2) it is 350 ° C as another example of iron based catalyst. The above indicates that iron oxyhydroxide converts to pyrrohotite at a very low temperature. The term "substantially iron oxide-free limonite iron ore" denotes an iron limonite ore containing less than 10% iron oxide mass that can be detected by X-ray powder diffraction.

The recycling step (h) should preferably be performed so that the total amount of liquid component and heavy products and waste in the liquid phase fraction to be returned to the first stage (a) is 10-130% by weight of the heavy oil from oil supplied for the first reaction step (a) (this total amount is called the BTM recycling amount). If the amount of BTM recycling is less than 10% by mass, there is only a small amount of heavy products and waste to be returned for the first reaction step (a) for hydro cracking providing a small production. If the amount of recycling is greater than 130% by mass, the slurry mixed with liquid component and liquid phase fraction decreases flowability, and this leads to a poor handling property. In particular, when the liquid component and the liquid phase fraction are supplied to the feedstock pulp preparation tank, the pulp obtained in the pulp preparation tank decreases flowability.

In the solid-liquid separation step (g), it is desirable to operate the precipitating-type solid-liquid separator at a temperature of 200-300 ° C and under a pressure of 20-40 kgf / cm2. In this temperature range the substance, consisting of heavy product and waste, is partially dissolved and undissolved components (organic and inorganic components) precipitate rapidly. This means efficient solid-liquid separation. In this pressure range the solid-liquid separating solvent remains unevaporated and thus allows stable pulp handling.

There are no specific restrictions for heavy metal oil containing heavy oil used as a feedstock in the present invention. This includes any heavy petroleum oil containing heavy metals, such as atmospheric and vacuum residue. It also includes super heavy oils containing heavy metals, such as bitumen, tar and oil sand. Heavy metals are those which cause catalytic contamination by Ni-Mo or Co-Mo catalysts, and are not specifically restricted and include for example Ni and V. According to the present invention, heavy petroleum oil hydrocracking is performed by means of the apparatus and The process flow rates shown in Figure 1. The following is a detailed description. As shown in Figure 1, the waste preparation tank (1) is fed with heavy petroleum oil, limonite iron ore (as catalyst) and sulfur. (as cocatalyst). They consist of the paste of feed material. The feedstock paste, together with hydrogen gas, is fed to the preheater (2). The preheated slurry enters the suspended bed reactor (3) in the first reaction step (a). Heavy oil hydro cracking is performed under a reaction pressure of 30-160 kgf / cm2, a reaction temperature of 430-455 ° C and a reaction time of 30 to 18021.

minutes The suspended bed reactor (3) may be a bubble column type reactor [first step (a)].

The first reaction step (a) gives rise to the reactor effluent [1]. This reactor effluent [1] enters the first gas-liquid separator [4] to be separated for the gas phase fraction [2] and the liquid phase fraction [3]. The first one contains gaseous products and unreacted hydrogen in a state of high pressure and high temperature.The second containing liquid products, the catalyst used in the first reaction step (a) and heavy metals [first gas-liquid separation step (b)] .The gas phase fraction [2] obtained by separation in the first gas-liquid separation step (b) enters the fixed bed reactor [5] filled with Ni-Mo or Co-Mo catalyst in which hydro cracking occurs under a reaction pressure of 30-160 kgf / cm2 at a reaction temperature of 310 ° C. -380 ° C (at start of operation) and at a net hourly spatial velocity of 0.3-2 hour * 1. The fixed delight reactor [5] may be a tubular flow reactor, for example [second reaction step (c)].

The reactor effluent [4] (containing hydrotreating products) obtained in the second reaction step (c) entrancing second gas-liquid separator [6] for separation to the gas phase fraction [5] containing unreacted hydrogen and to the product liquid phase [6] [second gas-liquid separation step (d)].

The liquid phase product [6] obtained by separation in the second gas-liquid separation step (d) enters the distillation column [7] for fractional distillation to yield desired fractions such as naphtha (C5-171 ° C), kerosene, and gas oil (171-343 ° C) and vacuum degassing oil (343-525 ° C) [distillation step (e)] Incidentally, the gas phase fraction [5] obtained in the second reaction step (c) enters in the gas purification step (g) and the purified gas is partly used as fuel gas and partly used as a cooling gas in the first reaction step (a) and the second reaction step (c).

The liquid phase fraction [3] obtained by separation in the first gas-liquid separation step (b) enters the flash separator [9], as a gas-liquid separator, for light product separation from the phaseliquid fraction under a lower pressure. than the first gas-liquid separation step [flash separation step (f)] The flash separation step (f) results (after separation from light products) in the faseliquid fraction [7] containing products such as heavy products , waste, the catalyst (used in the first stripping step (a)), and heavy metals. Part [7A] of the liquid phase fraction [7] is fed, together with a solid-liquid separation solvent, to the solid-liquid separator of the precipitation type [10] so that part [7A] is separated to a solid component [ 8] and for a liquid component [9] (both at normal temperature), said solid component containing said residue, catalyst used in the first reaction step (a) in heavy metals, said liquid component containing said medium products, medium to heavy products , and heavy products [solid-liquid separation step (g)] The liquid component [9] obtained by separation in the solid-liquid separation step (g) and the remainder of the liquid phase fraction [7] (or liquid phase fraction [ 7B] other than the liquid phase fraction [7A] supplied for the solid-liquid precipitation type separator) ([7B] = all of [7] - [7A]) are returned to the pulp preparation tank [1] or to the leit reactor the suspended [3] in the first reaction step (a) [recycling cycle (h)].

The fact that the liquid phase fraction obtained in the second gas-liquid separation step is returned to the first reaction step offers the advantage of reducing coke deformation and increasing low boiling point oil production. The reason for this is that the liquid phase phase fraction obtained in the second gas-liquid separation step is a highly hydro-treated component having a high hydrotreating capacity.

In accordance with the present invention, the petroleum heavy oil hydrocracking process is carried out by the apparatus and the process flow shown in Figure 2. In the case where the heavy oil contains a large amount of quinolines, indoles, benzothiophenes, etc., hardly desulphurized and denitrified, it is desirable that the second reaction step (c) and the second gas-liquid separation step which follows in the second stripping step (c) should be followed by a third stripping step. The second gas-liquid separator separates the oil produced in the second reaction step into a liquid phase and a gas fraction, the first containing high boiling fractions (such as quinolines, indoles, and benzothiophenes) and the second having a relatively boiling point. low. The gas phase fraction is denitrified and desulfurized in the third reaction step to give a good quality oil.

Specifically, the third reaction step is performed as follows. The gas phase fraction obtained in the second separation stage is fed to the Ni-Mo or Co-Mo catalyst-filled leitofixed reactor. The heat treatment in the fixed bed reactor is performed under a reaction pressure of 30 to 160 kgf / cm2 (preferably 50 to 100 kgf / cm2) at a reaction temperature of 310 ° to 420 ° C (preferably 340 ° C 390 ° C) at the beginning of the operation and at a net hourly liquid velocity of 0.3-2 hour "1 (preferably 0.5-1 hour" 1). The effluent from the reactor is separated into a liquid phase fraction and a gas phase fraction in the third gas-liquid separation step. The liquid phase fraction is distilled to separate to the desired fractions. The fasegasosa fraction is purge or recycled gas for the first, second, and third reaction steps for temperature control.

In this case, too, the liquid phase fraction obtained in the second gas-liquid separation step can be returned to the first reaction step to reduce coke production to improve oil production under boiling point. The reason for this is that the liquid phase fraction obtained in the second gas-liquid separation step is a highly hydro-treated component having a high hydrotreating capacity.

EXAMPLES

The present invention will be described with the following examples, which are not intended to restrict the scope of the invention.

Example 1

The first reaction step (a) was performed as follows. In this example, the vacuum residue (here, VR) from the Middle East was used as heavy oil containing heavy metals. This VR was incorporated with pulverized iron limonite ore as a catalyst in an amount of 0.6 mass% (in terms of iron) based on the total amount of VR and with sulfur as a catalyst in an amount of 1.2 times the amount of iron (1, 2 times, atomic ratio, the amount of iron in the limonite iron ore). Thus, a folder of feed material was obtained. Incidentally, VR had the fraction composition, shown in table 1.

Table 1

<table> table see original document page 25 </column> </row> <table>

The feedstock paste was placed in a 5 liter stirring autoclave. The autoclave was supplied with hydrogen gas for documented hydro cracking VR (vacuum residue or heavy oil). Hydro cracking was performed under a reaction pressure of 10 MPa (102 kgf / cm2), which is equal to the hydrogen supply pressure, at a reaction temperature of 450 ° C, and a reaction time of 60 minutes.

After the reaction was completed, the volume of gas product and each liquid fraction was examined. Also, each liquid fraction was examined for the amount of carbon, hydrogen, nitrogen, and sulfur by elemental analysis. Results are shown in tables 2-4.

Table 2

<table> table see original document page 26 </column> </row> <table>

In Table 2, the production of gas, liquid fractions and coke (collectively called fractions, etc., hereinafter) are expressed as terms of ratio (mass%) of the amount of fractions, etc. in relation to the amount of RV supplied. This definition applies to the production of tables 5, 6, 8, 10, 13, 14, 16, 18 and 19 given thereafter.

It is apparent from Tables 1 and 2 that the amount of fractions (or vacuum distillates) having a boiling point below 450 ° C increased from 4.4% by mass to 55.6% by mass [+ 8.4% by mass (fraction 1) + 28.7% by mass (fraction 2) + 18.5% by mass (fraction 3)] by hydro cracking while maintaining the decoque volume at 1.5 mass%. Incidentally, hereafter mass% will be expressed simply by%. Table 3

<table> table see original document page 27 </column> </row> <table>

It is clear from table 3 that hydro cracking increased the H / C ratio from 1.4 in the VR feed material to 2.0 in fraction 1, (with a boiling point less than 171 ° C) and to 1.7 in fraction 2 (with a boiling point of 171-343 ° C). It should also be noted that the sulfur content in fractions 1, 2 and 3 was 0.2%, 1.95%, and 2.8%, respectively. This suggests that the hydro cracking performed a desulphurization to some extent. In contrast, hydro cracking performed little denitrification.

Table 4

<table> table see original document page 27 </column> </row> <table>

It should be noted in Table 4 that metallic elements in the VR feed material were not detected in all distillates but were detected in THFI. This suggests that the metal elements were removed by adsorption on the suspended catalyst. Incidentally, THFI indicates those components that are insoluble in tetrahydrofuran.

Example 2

The hydrocracking of the first reaction step (a) was carried out under the same conditions as in example 1 in the fractionsop (or bottom or distillation residue having a boiling point greater than 450 ° C) obtained in example 1, which was charged together with the VR feed material. (The amount of the heavy fraction was 80% of the VR feed material quantity). Incidentally, the amount of limonite pulverized iron ore with catalyst was 0.6% by weight (in terms of iron) of that of VR feedstock, and the amount of sulfur as cocatalyst was 1.2 times the amount of iron, as in the case of example 1.

After the reaction was completed, the production of gas and liquid fraction was examined. Results are shown in table 5.

Table 5

<table> table see original document page 28 </column> </row> <table>

It should be noted in Tables 2 and 5 that deep recycling greatly reduces the amount of fraction + 450 ° C.

In Example 1 (without background recycling) the amount of fraction remaining as a vacuum distillation residue (or fraction amount + 450 ° C or fraction amount having a boiling point greater than 450 ° C) was about 40% (= 27%). , 4% + 9.7% + 1.5%), while in example 2 (bottom recycling) the fraction amount + 450 ° C is about 5% (= 1.6% + 0.7% + 2, 4%).

In addition, bottom recycling increased fractions output having a boiling point of less than 450 ° C, from about 55% to about 90%. This suggests that a large portion of VR feed material can be converted into light distillable fractions.

Despite the conversion to light oil for such high yields, coke yield was only 2%. Incidentally, distilled oil had identical properties (amount of C, Η, N, S, metal elements and H / C ratio) to those of Example 1. .Example 3

The oil fraction (having a boiling point less than 450 °) obtained in the first reaction step (a) in examples 1 and 2 corresponds to the gas phase fraction obtained by separation in the first gas-liquid separation step (b) downstream of the first step. of reaction (a) according to the heavy oil hydrocracking process of petroleum as defined in the present invention.

The distilled oil fraction obtained in the first stripping step (a) in Example 2 was fed to the Ni-Mo / AI2O3 catalyst-fixed bed-reactor column (AI2O3-supported Ni-Mo catalyst). The hydrotreatment was carried out at a reaction temperature of 350 ° C (at the beginning of the operation), under a reaction pressure of 10 MPa and a liquid hourly space velocity (LHSV) of 1 hour "1. [Second reaction step (c)] After the reaction, the yields of the eliquid gas fractions involved were examined and the quantities of C, Η, N, and S were analyzed by elemental analysis.

Results are shown in tables 6 and 7. <table> table see original document page 30 </column> </row> <table>

In Example 2, bottom recycling with boiling point greater than 450 ° C (remaining after vacuum distillation) increased the volume of fractions having a boiling point below 450 ° C to about 90%. By contrast, in Example 2 this fraction undergoes altograu hydrotreating by catalytic reaction in the bed reactor fixed in the second reaction step (c). Accordingly, the production volume of vacuum gas oil having a boiling point greater than 343 ° C (as end product oil) decreases to about 4% and light oil production tended to boiling below 343 ° C, reached about 80% as shown in table 6.

In addition, the light oil fraction with a boiling point lower than 34 3 ° C contains less than 10 ppm denitrogen; and naphtha, kerosene and gas oil contain sulfur in an amount of 60 ppm, 30 ppm and 90 ppm respectively as shown in table 7. This suggests that the process was capable of high degree desulphurization and nitrification. A probable reason for this is that high temperature (450 ° C) hydro cracking in the first reaction step (a) implies that the steric hindrance of these components which present denitrification and sulfurization difficulties is present.

Example 4, and Comparative Example 1 The VR feed material was incorporated with iron (or different type) catalyst and sulfur (as cocatalyst). The amount of sulfur was 1.2 times the amount of iron. The resulting slurry was fed to a hydro-cracking suspended bed reactor at a reaction temperature of 430-450 ° C under a reaction pressure of 10 MPa (under a hydrogen gas atmosphere), and at a reaction time of 60 minutes. The reaction product was analyzed for decoque production (or insoluble toluene production) in order to examine the relationship between iron catalyst type, reaction temperature and coke production.

Results are shown in table 8.

Table 8

<table> table see original document page 31 </column> </row> <table>

Decomposition of heavy oil (such as vacuum distillation residue) is conventionally performed by feeding the distillate to lixed catalytic reaction or fluidized bed catalytic reaction (FCC) through the coking process. The production of coke coking process is as high as 20-50%. Therefore, the effective use of heavy oil by this process was difficult.

In contrast, in the case of hydro cracking by a suspended bed reactor, coke production is low as mentioned above.

It is apparent from Table 8 that decoque production varies depending on the iron catalyst type and reaction temperature. In the case where the amount of spirituatural is 1% of the amount of feedstock VR, the coke yield will be 5.2% or 7.4%. In the case of limonite iron ore (as a catalyst in the first step) (a) according to the present invention) be used in an amount of 0.6% (in terms of iron) of the amount of VR feed material, the coke yield will be additionally low, and will be 1.0% at a temperature temperature of 450 ° C. It has been found that if iron ore limonite is used as a catalyst, coke production can be maintained at 2.4% even if the background oil was recycled. Iron limonite ore used as a catalyst in the first reaction step (a) According to the present invention it has a low cost and is still able to keep the amount of coke formation low. The reason for this is that iron limonite ore is converted to pyrrhotite even at low temperatures (around 250 ° C) and pyrrhotite shows catalytic activity for hydrotreating. This means that there is a catalyst in the early stages of heavy oil pyrolysis and promotes hydrotreating sand while suppressing coke formation due to polycondensation of radicals formed by thermal decomposition.

Example 5

The example demonstrates heavy metal-containing heavy oil hydrocracking using the process flow apparatus shown in Figure 1. (The process flow and apparatus shown in Figure 1 have been explained above).

The VR used for hydro cracking is a heavy oil containing heavy metals. The amount of iron ore-limonite was 0.6% by mass of the amount of VR and the amount of sulfur (cocatalyst) was 1.2 times the amount of iron. The reaction in the first stage of reaction (a) was performed under a reaction pressure of 10 MPa, a reaction temperature of 450 ° C, and a reaction time of 60 minutes. The catalyst in the second reaction step (c) was a Ni-Mo catalyst and the reaction in the second reaction step (c) was performed at a temperature of 350 ° C (at the beginning of the process) under a reaction pressure of 10 MPa and the spatial velocity. (1 hour -1) liquid capacity (LHSV).

The recycling step (h) was performed, in which the liquid component [9] and the residue of the liquid phase fraction [7] were returned to the pulp preparation tank in the first reaction step (a). reaction (a) to the recycling step (h) was repeated continuously for many months. A stable continuous operation was possible without any problems (such as reduction in catalyst activity and blocking of the catalyst layer), with a decoking yield of 2%.

The operation resulted in a large production of light oil (with high degree of hydrocracking) as follows. The production of light vacuum oil (boiling point greater than 343 ° C) was about 4%, and the production of The light oil (having a boiling point below 343 ° C) was about 80 ° C. The N content in the boiling light oil fraction greater than 343 ° C was less than 10 ppm, and the S content in naphtha, kerosene, and gas oil was 60 ppm, 30 ppm, and 90 ppm, respectively. This means that the operation performed denitrification and desulphurization to a great extent.

Heavy petroleum oil hydrocracking according to the present invention allows continuous operation with minimal coke accumulation and results economically in high light oil production with high degree of desulphurization and denitrification as well as high degree of hydrocracking. .

Example 6

The first reaction step (a) was carried out as follows. Petroleum vacuum residue from South America was used as heavy metal-containing petroleum oil (here VR-B) Limonite ferropulverized ore (as catalyst) was added in a quantity of 1% by mass of the quantity of VR-B.

Sulfur (as cocatalyst) was added at a quantity of 1.2 times (or 1.2 times (in atomic ratio) the amount of iron in the iron ore limonite). In this way the paste of feed material was obtained. When distilled, VR-B resulted in the fractions as shown in Table 9. Incidentally, VR-B was found to contain quinolines, indoles and benzothien in greater amounts than VR.

Table 9

<table> table see original document page 34 </column> </row> <table>

The aforementioned feedstock paste was placed in a 5 liter stirring autoclave. The autoclave was supplied with hydrogen gas for hydro-cracking of said VR-B. Hydro-cracking was performed under a reaction pressure of 100 kgf / cm2 (9.8 MPa) equal to the hydrogen supply pressure, and at a recycling temperature of 450 ° C, and at a reaction time of 60 minutes.

After the reaction was completed, the product gas production and the liquid fraction were examined. Also each liquid fraction was examined for amount of carbon, hydrogen, nitrogen, and sulfur by elemental analysis. Results are shown in tables 10 through 12. In table 10 gas volume, fractions and coke are expressed in terms of mass "relative to the amount of VR-B supplied. This definition applies to the production of tables 13, 14, 16 and 18 given maistarde.

Table 10

<table> table see original document page 35 </column> </row> <table>

It is apparent from tables 9 and 10 that hydro-cracking increased the amount of boiling fractions below 450 ° C, increasing from 7.0% by mass to 55.19% by mass (= 10.97 + 27.26 +16.96) by hydro cracking and keeping the coke production volume at 2.08% by mass.

Table 11

<table> table see original document page 35 </column> </row> <table> It becomes apparent from table 11 that hydro-cracking increased the H / C ratio from 1.4 in VR-B feed material to 1.97 in fraction 1, and to 1.72 in fraction 2. It should also be noted that the sulfur content in fractions 1, 2, and 3 was 0.24%, 0.48%, and 0.52%, respectively. This suggests that hydro cracking performed desulphurization to some extent. Conversely, hydro cracking performed very little nitrification.

Table 12

<table> table see original document page 36 </column> </row> <table>

It should be noted from table 12 that metallic elements in the VR-B feedstock were not detected in all fractions but were detected in THFI. This suggests that the metal elements were removed by adsorption in the suspended catalyst. Example 7

The hydro cracking of the first reaction step (a) was carried out under the same conditions as in example 6 on the heavy gravity fraction (or bottom residue or boiling point greater than 450 ° C) obtained in example 6, which had been charged. together with VR-B feed material. (The amount of heavy gravity component was 74.84% of the amount of VR-B feed material.) The amount of limonite was 1% (Fe terms) and the S / Fe ratio was 1.2.

The production of products is shown in table 13. It should be noted in tables 10 and 13 that the recycling of the fund greatly reduces the amount of fraction -450 ° C. In example 6 (without background recycling) the amount of waste was about 41, 93% by mass (30,87 + 8,98 + 2,08%), whereas in example 7 (with background recycling) the amount of waste becomes about 8,62% by mass (5,4% + 1 , 3% + 1.92%).

In addition, bottom oil recycling increased the production of boiling fractions below 450 ° C from 55.13% to 85.25% by mass. This suggests that a large portion of VR-B feed material can be converted into light distillable fractions. Despite conversion to light oil in such high yields, coke yield was kept less than 2%. Incidentally, distilled oil had properties identical to those of example 5.

Table 13

<table> table see original document page 37 </column> </row> <table>

Example 8

The reactor effluent obtained from the first stripping step (a) in example 7 underwent gas-liquid separation (in the first gas-liquid separation step) so that components with boiling point lower than 450 ° C were contained in the gaseous. Thus, gas phase fractionation was obtained. This gas phase was passed through a Ni-Mo / AI2O3 catalyst filled fixed bed reactor column, and the hydrotreatment was performed at a reaction temperature of 350 ° C (at the beginning of the reaction under a reaction pressure of 100kgf / cm2 ( 9.8MPa) and the hourly liquid space velocity of 1 hour "1 (second reaction step (c)).

Table 14 shows the outputs of the liquid gas products, and table 15 shows the results of element analysis.

Table 14 <table> table see original document page 38 </column> </row> <table>

It should be noted that the second reaction step promotes hydro-treatment. Production of fraction 3 decreased from 37.53% to 34.34% by mass. Nitrogen concentrations in fractions 1, 2, 3 were, respectively, 200, 960, 3000 ppm. The sulfur concentrations in fractions 1, 2, 3 were respectively 60, 100, 700 ppm. In other words, the nitrogen and sulfur concentrations in each fraction were higher than those of example 3 where the feed material was VR. Because compared to VR, VR-B contains more quinolines, indoles and benzothiophenes, which are hardly de-nitrified and desulfurized.

Example 9

The reactor effluent obtained from the second bypass step (c) in Example 8 has undergone gas-liquid separation (second gas-liquid separation step) so that a boiling point component of less than 343 ° C is contained in the gas phase. Thus the fasegasosa fraction was obtained. This gas phase fraction was passed through a fixed bed reactor column filled with Ni-Mo / Al 2 O 3 catalyst, and the hydrotreatment was performed at a reaction temperature of 380 ° C under a reaction pressure of 9.8 MPa and at an hourly liquid space velocity of 1 hour -1 (third reaction step (c)). Table 16 shows the production of gas and liquid products and table 17 shows the results of element analysis.

Table 16

<table> table see original document page 39 </column> </row> <table> Table 17

<table> table see original document page 40 </column> </row> <table>

It should be noted that the third reaction step promotes hydro-treatment. The yield of fraction 2 (171-343 ° C) decreased from 37.24% to 31.45%. In contrast, the production volume of fraction 1 (-171 ° C) increased from 13.04% to 18.63% by mass. As shown in Table 17, the nitrogen content was less than 10 ppm in the light oil fraction. with a boiling point below 343 ° C. The sulfur content was 30 and 50 ppm in fractions 1 and 2, respectively. This suggests that the third reaction step achieved a high degree of denitrification sulfurization. Example 10

The first reaction step (a) was performed as follows. VR-B was used as heavy oil containing heavy contendometals. Limonite pulverized iron ore (as catalyst) was added in an amount of 1% of the amount of VR-B. Sulfur (as a catalyst) was added in an amount 1.2 times the iron (or 1.2 times (atomic ratio) the iron in the Limonite iron ore). In this way the paste of feed material was obtained. The aforementioned paste was placed in an autoclaving with 5 liter agitation. The autoclave was supplied with hydrogen gas for hydro-cracking of the mentioned VR-BO hydro-cracking was performed under a pressure of 50 or 30 kgf / cm2, equal to the hydrogen supply pressure, at a reaction temperature of 450 ° C and in a reaction time 60 minutes.

After the reaction has been completed, the gas product production of each fraction of liquid is examined.

The results are shown in table 18. It should be noted in table 18 that hydro cracking resulted in high throughput even though the reaction pressure was 50 or 30 µg / cm2. However, the volume of coke increased slightly compared to the case where the reaction pressure was 100 kgf / cm2. When the reaction pressure was 30kgf / cm2, an oil production of 6.6, 77% by mass (11.12) was obtained. + 25.07 + 30.5 8).

Table 18

<table> table see original document page 41 </column> </row> <table>

Example 11

The first reaction step (a) was performed as follows.

VR-B was used as heavy oil of heavy contendometals. Limonite pulverized iron ore (catalyst) was added in an amount of 1% of that amount of VR-B. Sulfur was added at a quantity of 1.2 times the amount of iron (or 1.2 times (atomic ratio) the amount of iron in the iron ore limonimite). To VR-B was added the liquid phase fraction (obtained in the second gas-liquid separation step in figure 9) in an amount of 20% VR-B.

In this way the feed material paste was obtained.

The aforementioned feedstock paste was placed in a 5 liter agitation tank. The autoclave was supplied with hydrogen gas for hydro-cracking of said VR-B. Hydro cracking was performed under a reaction pressure of 30kgf / cm2 equal to the hydrogen supply pressure at a temperature of 450 ° C and a reaction time of 60 minutes. Once the reaction has been completed, the production volume of product gas and each fraction of liquids are examined. The results are shown in Table 19. It should be noted in Table 19 that a hydrocracking efficiency was achieved with reduced decoction and increased low boiling oil production even under a reaction pressure of 30kgf / cm2 if the Fraction of the liquid phase from the second gas-liquid separator is returned to the first reaction step.

Table 19

<table> table see original document page 42 </column> </row> <table>

Added to the liquid phase from the second liquid gas separator (Effect of the invention)

The present invention provides a heavy petroleum oil hydrocracking process containing heavy metals. This process allows a vacuum distillation residue (eg, heavy oil from heavy metals containing heavy metals) to be converted to light oil from oil with a minimum of coke buildup in prolonged continuous operation. The process results in a high hydrocracked denitrified yeastsulfurized oil.

Claims (6)

1. Heavy metal containing heavy oil hydro-cracking process, characterized in that it comprises the following steps (a) to (h): (a) a first reaction step for feeding a suspended bed reactor with a composite feedstock of heavy oil from heavy contendometals, iron limonite ore, and sulfur, together with hydrogen gas, thereafter hydrating the said heavy oil under a shear pressure of 50-100 kgf / cm2 at a reaction temperature of -440-450 ° C and at a reaction time of 60-120 minutes, (b) a first gas-liquid separation step to separate the reactor effluent obtained in said first reaction step (a) for a fraction (c) a second reaction step for feeding a fixed bed reactor filled with Ni-Mo or Co-Mo catalyst with the gas phase fraction obtained by separating from said first sep step hydration treatment under a reaction pressure of 50-100 kgf / cm2, at a temperature of 330-360 ° C (at the beginning of the operation), at an hourly liquid velocity of 0.5-1 (d) a second gas-liquid separation step to separate the reactor effluent that was obtained from said second reaction step (c), for a gas phase fraction and for a liquid phase fraction; distillation to separate by distillation the liquid phase fraction, which was obtained by separating from said second gas-liquid separation step (d) for the prescribed fractions, (f) a flash separation step for performing flash separation under reduced pressure in the fraction. faseliquid, which was obtained by separating from said first gas-liquid separation step (b), so that the light reaction product in the liquid phase fraction passes gas gaseous (g) a solid-liquid separation step p For feeding a solid-liquid type precipitating separator with part of the liquid phase fraction, operated at a temperature of 200-300 ° C and under a pressure of 20-40kgf / cm2, containing medium products, medium weight, heavy products, residue, heavy metals and catalysts used in said first reaction step (a), which were obtained by separation from the light product in the flash separation step (f), and a solid-liquid separation solvent, thereby separating them into a solid component and to a liquid (both at normal temperature), said solid component containing the citidose, the catalyst used in said first step (a), and heavy metals, said liquid component containing said medium, heavyweight and heavy products; and (h) a recycling step for returning to the first reaction step (a) the liquid component obtained by separation in said solid-liquid separation step (g), and the liquid phase fraction containing the medium products, the average products for separation. heavy products, the heavy products, and residue, heavy metals and catalyst used in said first reaction step (a), which were obtained by separation from the light product in said flash separation step (f), as the remainder left after said solid separator -liquid of the type of precipitation has been fed. Process according to Claim 1, characterized in that the limonite iron ore with catalyst in the first reaction step (a) is added by an amount of 0.3-2% by weight (in terms of iron) of the amount of heavy oil of petroleum and sulfur as a catalyst be added in an amount (at the atomic ratio) of 1-3 times the amount of iron contained in the limonite iron nominium. Process according to Claim 1 or -2, characterized in that the first reaction step ferrone limonite ore (a) is used in fine powder form having an average particle diameter of less than 2 µm, which is prepared. mechanically spraying in an oil solvent. Process according to any one of claims 1 to 3, characterized in that the limonite iron ore in the first reaction step (a) is one which does not contain iron oxide. Process according to any one of Claims 1 to 4, characterized in that the total amount of the liquid component, the heavy product, and the liquid phase fraction residue is 10-130% by mass of the amount of the petroleum heavy oil supplied to the former. reaction step (a). Process according to any one of claims 1 to 5, characterized in that it comprises the following steps: i) a third reaction step in which the gas phase fraction obtained in the second gas-liquid separation step (d) is subjected to hydro cracking by means of a Ni-Mo or Co-Mo catalyst under a reaction pressure of 50-100 kgf / cm2, a reaction temperature of 340-390 ° C (at the beginning of operation), and a spatial speed (ii) a third gas-liquid separation step in which the reactor effluent obtained in the third stripping step is separated into a gas phase and a liquid phase; iii) a distillation step in which the liquid phase fraction obtained by separation in the third liquid gas separation step is distilled, and (iv) a recycling step, in which the liquid phase fraction obtained in the second separation step is returned to the first reaction step (a).