CA2314033C - Process for hydrocracking of petroleum heavy oil - Google Patents

Abstract

A process for hydrocracking of petroleum heavy oil containing heavy metals, such as vacuum residue, which gives desulfurized and denitrified light oil economically in high yields, while suppressing coke formation in prolonged continuous operations. The process consists of a first reaction step, in which a feedstock slurry composed of said petroleum heavy oil and limonite iron ore (as a catalyst) is supplied, together with hydrogen, into a suspended-bed reactor for hydrocracking of said heavy oil, a second reaction step, in which the gas-phase fraction [2]
in the effluent [1] from said reactor is fed into a fixed-bed reactor filled with a Ni-Mo or Co-Mo catalyst for hydrotreating, a distillation step, in which the liquid phase fraction in the effluent from the reactor is distilled, and a recycling step, in which the liquid phase fraction [3] in the effluent [1] from the reactor and the heavy gravity component in the liquid phase fraction [3] are returned to the first reaction step.

Description

PROCESS FOR ;HYDROCRACRING OF PETROLEUM HEAVY OIL
BACKGROUND OF THE INVENTION
1. Field oi: the Invention:
The present invention relates to a process for hydro-cracking of petroleum heavy oil, and more specifically, to a process for hydrocracking of petroleum heavy oil contain-ing heavy metals. In particular, the present invention relates to a process for hydrocracking petroleum heavy oil containing heavy metals (such as atmospheric residue and vacuum residue), thereby yielding light oil.

2. Description of the Related Arts:
Increasing attention is paid to the cracking technol-ogy for producing light oil, which is in short supply, from heavy oil, which is in oversupply. This technology is becoming important more and more in view of the fact that petroleum reserves are decreasing inevitably.
There have been proposed many processes for thermal cracking or hydrocracking of heavy oil; however, they pose problems in some wary or other with conversion of heavy oil (such as vacuum residue) into light oil. This will be fully explained in the following.
Heavy o:il, such as vacuum residue, tends to contain a large amount of nitrogen compounds and sulfur compounds.
Moreover, it also contains a large amount of organometallic impurities which ax-e extremely detrimental to cracking in the presence of catalysts. Such metallic impurities in-elude mostl~~ those which contain nickel (Ni) and vanadium (V) and other metals, with less frequency. They chemically combine with comparatively polymeric organic compounds such as asphalten.e in heavy oil, and they considerably inhibit the catalytic activity for decomposition and removal of nitrogen-, sulfur-, and oxygen-containing compounds and hydrocracking of heavy organic matter.
There is a known process for thermal cracking of heavy oil, such as vacuum residue, without using catalysts. It is the so-called coking process. This process suffers the disadvantage of yielding a large amount of coke as by-product, the disposal of which is a problem, and also yielding a large amount of gas due to overcracking, thereby decreasing in yields of distillates. Another disadvantage is that the resulting distillate is poor in quality because of high content of aromatic and olefinic components.
There is a hydrocracking process of fixed-bed type, which is performed in a reactor filled with a granular catalyst. The disadvantage of this process is that the catalyst is vulner<3ble to asphaltene and heavy metals, such as V and Ni, in heavy oil when cracking takes place exces-sively. Thus the catalyst decreases in activity and the catalyst layer becomes clogged due to gradual deposition of coke as by-product and heavy metals. This prevents pro-longed continuous operations.
There is another hydrocracking process which is car-ried out in an ebu:Llated-bed reactor by the aid of granular catalyst formed by extrusion. This process brings about such vigorous agitation in the reactor that it poses no problems with increase in pressure loss due to coke accumu-lation. In addition, it permits catalyst charging and discharging during operation and hence it permits prolonged continuous operations while keeping the catalyst activity at a certain level. Its advantage over hydrocracking of fixed-bed type is reduced by the fact that it needs the catalyst to be circulated during operation, thereby causing mechanical damages to pumps etc. This makes its operation difficult. It has another disadvantage of requiring expen-sive catalyst, requiring a high reaction pressure (usually 150-200 kg/cm2), and giving rise to products which are not desulfurized. and denitrified sufficiently.
OBJECT AND SUMMARY OF THE INVENTION
The present invention was completed to address the above-mentioned problems. It is an object of the present invention to provide a process for producing light oil by hydrocracking from petroleum heavy oil, such as vacuum residue, containing heavy metals. The hydrocracking proc-ess of the present invention is characterized by its abil-ity to run continuous operation for a long time without coke accumulation and to yield desulfurized and denitrified light oil with a high degree of cracking economically in high yields.
The above-mentioned object of the present invention is achieved by a process for hydrocracking of petroleum heavy oil which is defined in the appended claims. The process is outlined below.
The pre:>ent invention is directed to a process for hydrocracking of petroleum heavy oil containing heavy met-als, said process comprising the following steps (a) to (h).
(a) a first reaction step of feeding a suspended-bed reactor with a feedstock slurry composed of petroleum heavy oil containing heavy metals, limonite iron ore as a cata-lyst, and sulfur a~~ a cocatalyst, together with hydrogen gas, thereby performing hydrocracking on said heavy oil at a reaction p~°essure~ of 30-160 kg/cmz (2.9-15.7 MPa), at a reaction temperature of 430-455°C, and for a reaction time of 30-180 minutes.
(b) a first gas-liquid separation step of separating the reactor effluent, which has been obtained in said first reaction step (a), into a gas phase fraction and a liquid phase fraction.
(c) a second reaction step of feeding a fixed-bed reactor filled with a Ni-Mo or Co-Mo catalyst with the gas phase fraction obtained by separation from said first gas-liquid separation step, thereby performing hydrotreating at a reaction pressure of 30-160 kg/cmz, at a reaction tem-perature of 310-380°C (at the start of operation), and at a liquid hourly space velocity of 0.3-2 hr'1.
(d) a second gas-liquid separation step of separating the reactor effluent, which has been obtained in said sec-and reaction step (c), into a gas phase fraction and a liquid phase fraction.
(e) a distil:Lation step of separating by distillation of the liquid phase fraction, which has been obtained by separation from said second gas-liquid separation step (d), into prescribed fractions.
(f) a flash separation step of performing flash sepa-ration under lower pressure than that of said first gas-liquid separation atep (b) on the liquid phase fraction , which has been obtained by separation from said first gas-liquid separation step (b), such that the light reaction product in the liquid phase fraction vaporizes to the gas phase.
(g) a solid-liquid separation step of feeding a solid-liquid separ<~tor of settling type with part of the liquid phase fraction containing the medium product, medium to heavy product, heavy product, residue, heavy metals and a catalyst used in said first reaction step (a), which have been obtained by sE~paration from the light product in said flash separation step (f), and a solvent for solid-liquid separation, thereby separating them into a solid component and a liquid component (both at normal temperature), said solid component containing said residue, the catalyst used in said first reaction step (a), and heavy metals, said liquid component containing said medium product, medium to heavy product, and heavy product.
(h) a recycling step of returning to said first reac-tion step (a) the liquid component obtained by separation in said solid-liquid separation step (g) and the liquid phase fraction containing the medium product, medium to heavy product, heavy product, residue, heavy metals and a catalyst used in said first reaction step (a) which have been obtained by separation from the light product in said flash separation step (f), as the remainder left after feeding to said solid-liquid separator of settling type.
BRIEF DESCRIPTION OF THE DRAWINGS
Fig. 1 is a schematic diagram showing an example of the hydrocracking process for petroleum heavy oil pertain-ing to the present invention.
Fig. 2 .is a schematic diagram showing another example of the hydrocracking process for petroleum heavy oil pertaining to the p resent invention.
DESCRIPTION OF THE PREFERRED EMBODIMENTS
The preaent invention is practiced in the following manner, which is given as an example.
The process of: the present invention uses, as a feed-stock, petroleum heavy oil, such as vacuum residue, con-taining heavy meta:Ls. This feedstock is incorporated with limonite iron ore, as a catalyst, and sulfur, as a cocata-lyst, and it is used in the form of slurry.
The feedstock slurry is fed, together with hydrogen gas, to a suspended-bed reactor, and it undergoes a first reaction step (a) i'or hydrocracking of said heavy oil at a reaction pressure of 30-160 kg/cm2, at a reaction tempera-ture of 430-455°C, and for a reaction time of 30-180 min-utes. Reactor eff:Luent [1] containing hydrocracking prod-ucts, which has been obtained in said first reaction step, undergoes a first gas-liquid separation step (b) for sepa-ration into a gas phase fraction [2] containing gaseous products and unreacted hydrogen and a liquid phase fraction [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 a fixed-bed reac-for filled with a Ni-Mo or Co-Mo catalyst. It undergoes a second reaction step (c) for hydrotreating which is carried out at a reaction pressure of 30-160 kg/cmz, at a reaction temperature of 310-380°C (at the start of operation), and at a liquid :hourly space velocity of 0.3-2 hr-1. Reactor effluent [4] containing hydrotreating products which has been obtained in sen d second reaction step undergoes a second gas-liquid separation step (d) for separation into a gas-phase fraction [5] containing unreacted hydrogen gas and a liquid phase fraction [6]. The liquid phase fraction [6] obtained by separation in said second gas-liquid sepa-ration step (d) undergoes a distillation step (e) for sepa-ration into prescribed fractions. The liquid phase frac-tion [3] obtained by separation in said first gas-liquid separation step (b) undergoes a flash separation step (f) for flash separation under lower pressure than that of said first gas -liquid separation step in such a way that the light product in the liquid phase fraction vaporizes to the gas phase. The flash separation step (f) yields the liquid phase fraction [7] containing a medium product, medium to heavy product, heavy product, residue, heavy metals and a catalyst used in said first reaction step [a] which have been separated from the light product. Part [7A] of the liquid phase fraction [7] is fed to a solid-liquid separa-for of settling type, together with a solvent for solid-liquid separation. This solid-liquid separator carries out a solid-liquid separation step (g) for separation into a solid component [8] and a liquid component [9] (both at normal temperature), said solid component containing said residue, the catalyst used in said first reaction step (a), and heavy metals, said liquid component containing said medium product, medium to heavy product, and heavy product.
Finally, a recycling step (h) is carried out to return to said first reaction step (a) the liquid component (9], which has been obtained by separation in said solid-liquid separation step (g), and the remainder of said liquid phase fraction [7] (or t:he liquid phase fraction [7B] other than the liquid phase fraction (7A] supplied to said solid-liquid separator of settling type) ([7B] = all of [7] -[7A] ) .
The foregoing is the way in which the hydrocracking process of the present invention is performed on petroleum heavy oil. The distillation step (e) gives a prescribed fraction (light oi:1) continuously or intermittently.
The present invention produces its effect as explained in the following.
According to t:he present invention, the hydrocracking process for petroleum heavy oil starts with a first reac-tion step (a), in which petroleum :heavy oil incorporated with limonite iron ore as a catalyst undergoes hydrocrack-ing in a suspended-bed reactor. The limonite iron ore is rather poor in catalytic activity for desulfurization and denitrification but has a very high catalytic activity for hydrocracking. Moreover, it is an inexpensive disposable catalyst. The ent:Lre process is run in such a way that intermediate produ<:ts in each step are returned to the first reaction step (a) by the recycling step (h). The intermediate produ<:ts include the liquid component [9]
containing heavy products which is obtained in the first reaction step (a) and its subsequent steps (b) to (f), and the liquid phase fraction [7] containing said heavy prod-ucts as well as used catalyst. The recycled liquid compo-nent [9] and liquid phase fraction [7] pass through the first reaction step (a) so that the heavy products con-tained therein undergo hydrocracking. This hydrocracking is facilitated by t:he used catalyst remaining in the liquid phase fraction [7]. (The used catalyst is still active for hydrocracking and hence it works effectively in the first reaction step (a).) owing to the first reaction step (a) and the recycling step (h), the process of the present invention permits production of light oil in high yields without coke accumulation. In addition, the process of the present invention is economical because the first reaction step employs an inexpensive catalyst which is recycled to it.
The fir:at gas-liquid separatian step is designed to separate the reactor effluent [1] containing hydrotreating products, which has been obtained in said first reaction step (a), into a gals-phase fraction [2] containing gaseous reaction products and unreacted hydrogen and a liquid phase fraction [3] containing liquid reaction products and said catalyst (the same one as used in said first reaction step (a)) and heavy metals. In the ensuing second reaction step (c), the thus separated gas-phase fraction [2] undergoes hydrotreating in a fixed-bed reactor filled with a Ni-Mo or Co-Mo catalyst. It: is to be noted that the gas-phase frac-tion [2] fed to the' second reaction step (c) does not con-tain asphalt~ene produced in the first reaction step (a), and almost all of asphaltene, together with heavy metals, are present in the liquid phase fraction [3] separated in the first gays-liquid separation step (b). Therefore, gas-phase fraction [2] undergoes hydrotreating in the fixed-bed reaction in the absence of heavy metals and asphaltene which deactivate tree catalyst for desulfurization and deni-trification and give rise to coke which causes catalyst deactivation and catalyst layer clagging. The result is the prevention of coke accumulation and catalyst poisoning and the extension of catalyst life, and the prolonged con-tinuous operation. High degree of desulfurization, deni-trification and hydrocracking take place in the second reaction step. The second reaction step (c) is followed by the second c~as-liquid separation step (d), the distillation step (e), the flash separation step (f), the solid-liquid separation step (g), and the recycling step (h), as men-tinned above. The first gas-liquid separation step (b) and the second reaction step (c) are mainly responsible for the prevention of coke accumulation and prolonged continuous operation. As the result, the entire process permits con-tinuous production of highly desulfurized and denitrified light oil.
In summary, a~~cording to the present invention, the process for hydrocracking of petroleum heavy oil accom-plishes hydrocracking of petroleum heavy oil (such as vac-uum residue) containing heavy metals economically in high yields for prolonged continuous operation without coke accumulation. Thus the process permits production of highly desul.furized and denitrified light oil with a high degree of cracking.
A detailed description is made below of the effect produced by the present invention.
The first reaction step (a) employs a catalyst for hydrocracking. Basically, this catalyst should be highly active (for efficient hydrocracking) and readily available at a low prp'.ce. Among inexpensive catalysts are iron-based ones such as iron sulfide, iron oxide, and red mud. Unfor-tunately, they have the disadvantage of being poor in cata-lytic activity. The present inventors extensively studied the effect of iron-based catalysts on hydrocracking of petroleum heavy oil. As the result, it was found that limonite iron ore exhibits high catalytic activity when used for petroleum heavy oil. This finding has led the present inventors to employ limonite iron ore as the cata-lyst for hydrocrac:king in the first reaction step (a). In the first reaction step (a), limonite iron ore (as a cata-lyst for hydrocrac:king of petroleum heavy oil containing heavy metals) is used in combination with sulfur as a co-catalyst. Heavy o:il incorporated with the catalyst and cocatalyst takes on a form of slurry. This feedstock slurry together with hydrogen gas are fed to a suspended-bed reactor. Hydrocracking of heavy oil is carried out at a reaction pressure of 30-160 kg/cm2, a reaction tempera-ture of 430-455°C, and for a reaction time of 30-180 min-utes.
The first reaction step (a) gives the reactor effluent [1] which contains hydrotreating products. This effluent (1] is fed t~o the i:irst gas-liquid separation step (b) for separation into the' gas-phase fraction [2] containing gase-ous reaction products and unreacted hydrogen and the liquid phase fraction [3] containing liquid products and the cata-lyst (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] separated from the liquid phase fraction [3]

is fed to the fixed-bed reactor filled with a Ni-Mo or Co-Mo catalyst. Hydrotreating is carried out at a reaction pressure of 30-160 kg/cm2, at a reaction temperature of 310-380°C (at. the start of operation), and at a liquid hourly space velocity of 0.3-2 hr~l. The gas-phase frac-tion [2] is composed with hydrogen sulfide gas (HzS), hy-drocarbon gas (C1-f4), light products (CS-171°C), medium product (171-343°C), medium to heavy products (343-450°C), but it contains very little heavy metals and asphaltene which adversely affect the catalytic activity in the second reaction step (c). This may be reasoned as follows. The first reaction step (a) includes decomposition of those compounds foamed byr bonding of heavy metals with organic compound in heavy oil. Heavy metals liberated by decompo-sition are adsorbed by the catalyst and thus separated from organic compounds or hydrocracking products of heavy oil (obtained in the first reaction step (a)). Thus, almost all of heavy metal components in the petroleum heavy oil (as a feedstock) i~~ contained, together with the catalyst, in the liquid phase fraction [3] in the first gas-liquid separation svtep (b). The liquid phase fraction [3] is eventually d:ischarc~ed into the downstream step. In addi-tion, the first reaction step (a) involves decomposition of asphaltene contained abundantly in heavy oil. Thus, the effluent from the first reaction step (a) decreases in asphaltene content. In the first gas-liquid separation step (b), asphaltene (which has a very high boiling point) does not vaporize the gas phase but mostly remains, to-gether with heavy metals, in the liquid phase. It follows, therefore, that the gaseous reaction product in the gas phase fraction, which has entered the second reaction step (c), undergoes hydrocracking in the fixed-bed reactor in the absence of heavy metals (which deactivate the catalyst for desulfurization and denitrification) and asphaltene (which causes coke accumulation, catalyst deactivation, and catalyst layer clogging).
Moreover, the first reaction step (a) involves dealky-lation which. reduces steric hindrance to catalyst active sites by alkali substituents. Thus, the second reaction step (c) permits the Ni-Mo or Co-Mo catalyst to fully ex-hibit its catalytic activity for desulfurization and deni-trification.
The practice mentioned above produces the effect of preventing coke accumulation and catalyst poisoning, ex-tending catalyst life, and permitting stable prolonged continuous operation. The result is the production of hydrotreated pure oil through desulfurization and denitri-fication at a high degree. The hydrotreated oil is accom-panied by hydrocarbon gas, hydrogen sulfide gas, and ammo-nia gas resulting :From hydrotreating.
The second reaction step (c) gives the reactor efflu-ent [4] containing hydrotreated products. This effluent enters the second gas-liquid separation step (d) for sepa-ration into the ga:a-phase fraction [5] containing unreacted hydrogen and the liquid phase fraction (6).
The thus obtained liquid phase fraction (61 enters the distillation step (e) for separation into prescribed frac-tions by distillation. The fractions include naphtha (light product, CS-171°C), kerosene and gas oil (medium product, 171-343°C), and vacuum gas oil (medium to heavy product, 343°C +) .
The liq~zid phase fraction [3] obtained by separation in the first gas-liquid separation step (b) enters the flash separation step (f), which causes the light product to vaporize from the liquid phase fraction [3] into the gas phase.
As mentioned above, the first reaction step (a) em-ploys limonite iron ore as an inexpensive disposable cata-lyst. After the light product has been separated into the gas phase and the :Liquid phase fraction [7] in the flash separation step (f) connected to the liquid-phase line from the first gas-liqu:Ld separation step (b), the catalyst is separated from the liquid phase and discharged, together with residue and heavy metals, from the system, in the solid-liquid separation step (g).
Separat:LOn of the catalyst (solid) from the liquid is usually accomplished by distillation, centrifugation, or filtration. The d:LSadvantage of distillation is a large loss of oil which results from the fact that the solid-liquid mixture is discharged from the bottom of the distil-lation column and hence a large amount of liquid accompa-vied by solids is discharged from the system. The disad-vantage of centrifugation or filtration is high equipment cost and low efficiency for the fluid containing fine par-ticles (tens of micrometers at the largest) as in the case of the present invention. Consequently, the present inven-tion employs the solid-liquid separation step (g) of set-tling type in which a solvent is used for solid-liquid separation. A detailed description follows.
As mentioned above, the flash separation step (f) gives the liquid phase fraction [7] separated from the light product. The liquid phase fraction [7] contains medium products, medium to heavy products, heavy products, residues, catalyst (used in the first reaction step (a)), and heavy metals (c:ontained in petroleum heavy oil as the raw material). If this liquid phase fraction [7] is to be separated into liquid and solid by a solid-liquid separator of settling type, :separation needs a large-scale equipment because the :Liquid phase fraction 1.7] contains fine solid particles which are slow in settling. This problem is solved by supplying the solid-liquid separator of settling type with a aolvent: which has a lower boiling point than the liquid phase fraction [7] and partly dissolves the heavy products and residues. The substance remaining un-dissolved consists of pure solids and pseudo solids having the intermediate properties of solid and liquid. The lat-ter functions as a binder for inorganic solids (such as catalyst, heavy metal adsorbed to catalyst, and coke), thereby promoting flocculation. As the result, solid par-ticles grow and settle in a short time. The resulting precipitate (in the form of sludge) is discharged from the bottom of th.e solid-liquid separator of settling type. The solvent is recovered and recycled. The liquid is dis-charged from. the top of the solid-liquid separator of set-tling type. The discharged liquid contains very few inor-ganic solids; it has its solvent recovered by flash separa-tion. The recovered solvent is recycled to the solid-liquid separation step (g), and the remainder is recycled to the first reaction step (a) .
In other word:., what is accomplished in the solid-liquid separation step (g) is as follows. The solid-liquid separator of settling type is fed with the part [7A) of the liquid phase fraction (7] from the flash separation step (f) which has been separated from light products. This liquid phase fraction [7] contains medium products, medium to heavy products, heavy products, residues, catalyst (used in the first reaction step (a)) and heavy metals. The part [7A] of the liquid phase fraction is fed together with a solvent for solid-7Liquid separation. The separator sepa-rates it into two components - one being a solid component [8] (at normal temperature) containing the residues, the catalyst (used in t:he first reaction step (a)), and heavy metals, the other being a liquid component [9] containing the medium product:>, 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 liquid phase fraction [7] is also recycled to the first reaction 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 to the solid-liquid separator of settling type. [7B] = all of [7] - [7A] ) What is recycled to the first reaction step (a) con-sists of the liquid component [9] and the remainder of the liquid phase fract:Lon [7]. They are mostly heavy products (or heavy distillate) having a boiling point of 450-525°C
under normal pressure. They readily undergo hydrocracking in the first reaction step (a). The remainder of the liq-uid phase fraction [7] contains used catalyst in addition to heavy products, and this used catalyst still possesses catalytic activity for hydrocracking. Therefore, this used catalyst functions effectively in the first reaction step (a), making it possible to reduce the amount of the cata-lyst to be replenished.
The foregoing is the reason far production of light oil in high :yields. The inexpensive catalyst and its recy-cling to the first reaction step (a) contribute to the economical p:roducti.on.
In conclusion, according to the present invention, the process for lzydrocracking of petroleum heavy oil accom-plishes hydrocracki.ng of petroleum heavy oil, such as vac-uum residue, containing heavy metals economically in high yields for prolonged continuous operation without coke accumulation.. Thus the process permits production of highly desulfurized and denitrified light oil with a high degree of cracking.
According to the present invention, the process for hydrocrackin.g of petroleum heavy oil involves the first reaction step (a) employing a suspended-bed reactor in which reaction is .carried out at a reaction pressure of 30-160 kg/cm2, ,3t a reaction temperature of 430-455°C, and for a reaction time of 30-180 minutes. The reason for this is explained in the following.
The reaction pressure should be 30-160 kg/cm2 because reaction at a pressure lower than 30 kg/cm2 [9.80665 x 10°
Pa/(kg/cm2) X 60 kg/cm2 = 2.9 X 106 Pa = 2.9 MPa] permits coke formation in :Large quantities due to low hydrogen partial pressure and reaction at a pressure higher than 160 kg/cmz leads to high production cost without any additional effect. The reaction temperature should be 430-455°C be-cause reaction at a temperature lower than 430°C does not promote hydrocrack:ing but presents difficulties in produc-tion of light oil :in high yields and reaction at a tempera-ture higher than 4!i5°C brings about vigorous pyrolysis and increases the rate of polycondensation, thereby sharply increasing coke formation. The reaction time should be 30-180 minutes becauss~ reaction for less than 30 minutes pre-Bents difficulties in production of oil in high yields and reaction for more than 180 minutes causes hydrocracking to proceed too far, giving rise to a large amount of hydrocar-bon gas, with the :result that the amount of light oil pro-duced decreases and the amount of hydrogen consumption increases more than necessary, which leads to high produc-tion cost.
Reaction in the suspended-bed reactor should prefera-bly be carried out at a reaction pressure of 50-100 kg/cm2, at a reaction temperature of 440-4.50°C, and for a reaction time of 60-120 minutes. Reaction under these conditions keeps coke formation low and permits economical production of light oil in high yields.
The second reaction step (c) employs a fixed-bed reac-for in which reaction is carried out at a reaction pressure of 30-160 kg/cmz, at a reaction temperature of 310-380°C
(at the start of operation), and ar_ a liquid hourly space velocity of 0.3-2.0 hr-1. The reason for this is explained in the following.
The rea<:tion pressure should be 30-160 kg/cmz, which is identical with that in the first reaction step (a). In other words, the liquid phase fraction from the first gas-liquid separation step (b) is not pressurized for economi-cal reason. Reaction at this pressure prevents coke accu-mulation and permits prolonged continuous operation. The reaction temperature should be 310-380°C (at the start of operation) because reaction at a temperature lower than 310°C is slow and incomplete for desulfurization, denitri-fication, and hydrocracking, and reaction at a temperature higher than 380°C yields more hydrocarbon gas (with less oil) due to excessive hydrocracking. The liquid hourly space velocity should be 0.3-2 hr-1 because a space veloc-ity lower than 0.3 hr~l does not improve reactivity for desulfurization and denitrification (which leads to high production cost) and a space velocity higher than 2 hr-1 results in inefficient desulfurization and denitrification.
Reaction in the fixed-bed reactor should preferably be carried out at a reaction pressure of 50-100 kg/cmZ, at a reaction temperature of 330-360°C, and at a liquid hourly space velocity of 0.5-1 hr-1. Reaction under these condi-tions is most efficient and economical for prolonged con-tinuous operation and permits production of highly desul-furized and denitrified light oil.
In the :first reaction step (a), limonite iron ore as the catalyst should be added in an amount of 0.3-2 mass%
(in terms of iron) based on the amount of petroleum heavy oil. With an amount less than 0.3 mass%, the catalyst tends to produce more asphaltene and coke. With an amount more than 2 mass%, the catalyst does not work effectively to produce oil having a boiling point lower than 450°C.
Sulfur as the cocar_alyst should be added so as to convert the limonite iron ore into pyrrhotite which exhibits the catalytic activity. The amount of sulfur should be more than that of iron :in limonite iron ore (in atomic ratio), with the upper limit being three times for economical rea-son.
The limonite iron ore used as a catalyst in the first reaction step (a) should preferably be one which is in the form of fines powder having an average particle size smaller than 2 Vim. A catalyst having a large average particle size has a small effective surface area necessary for efficient contact with. petroleum heavy oil and hence is poor in cata-lytic activity. A catalyst having an average particle size smaller than. 2 ~m has a large effective surface area and hence has high catalytic activity. For this reason, the average particle size of the catalyst should preferably be smaller than. 1 Vim.
Limonite iron ore in the form of fine powder having an average particle size smaller than 2 um can be obtained by mechanical pulverizing of limonite iron ore. This pulver-izing should preferably be carried out by wet process (in a petroleum solvent) rather than by dry process. The powder obtained by dry process (using a pneumatic crusher) coagu-lates remarkably in the feedstock slurry containing heavy oil and catalyst and hence is poor in dispersibility, whereas the powder obtained by wet process (in a petroleum solvent) does not coagulate in the feedstock slurry but exhibit good dispe:rsibility, which leads to high catalytic activity.
Limonite iron ore is known to contain a-iron oxyhydro-xide and a-iron oxide by powder X-ray diffraction. It varies in composition depending on its origin (place and mine lot). The present inventors investigated the effect of the composition on the catalytic activity. It was found that limonite iron ore substantially free of iron oxide has the highest catalytic activity. Therefore, it is desirable to use limonite iron ore substantially free of iron oxide in the first: reaction step (a) .
A mention is made below of the reason why limonite iron ore substantially free of iron oxide has the highest catalytic activity.
In general, any iron compound is sulfurized with sul-fur or sulfur compounds to form iron sulfide called pyr-rhotite (Fel_XS), which is an active species exhibiting the catalytic activity. If conversion into pyrrhotite takes place at a low temperature, it means that the active spe-cies exists before the feedstock for hydrocracking begins to decompose. As the result, pyrolytic radicals are sup-plied sufficient hydrogen. This prevents polymerization of pyrolytic radicals and promotes hydrocracking (meaning high catalytic activity).
Conversion of a-iron oxyhydroxide into pyrrhotite takes place at a lower temperature than a-iron oxide.
Therefore, limonite iron ore substantially free of iron oxide has th.e highest catalytic activity. Incidentally, the temperature of conversion into pyrrhotite is 200°C for a-iron oxyhydroxide and 350°C for a-iron oxide. It is 350°C in the case of natural pyrite (FeS2) as another exam-ple of the iron-based catalyst. The foregoing indicates that a-iron oxyhydroxide converts into pyrrhotite at a very low temperature. The term "limonite iron ore substan-tially free of iron oxide" denotes limonite iron ore con-taining less than :LO mass% of a-iron oxide which can be detected by powder x-ray diffraction.
The recycling step (h) should preferably be carried out such that the i:otal amount of the liquid component and the heavy product Bind the residues in the liquid phase fraction, to be returned to the first reaction step (a) is 10-130 mass% of the' petroleum heavy oil supplied to the first reaction step (a). (This total amount is referred to as BTM recycling amount.) If the BTM recycling amount is less than 10 mass%, there is only a small amount of heavy product and :residues to be returned to the first reaction step (a) for hydrocracking, and this leads to a low yield.
If the BTM recycling amount is more than 130 mass%, the slurry mixed with t:he liquid component and liquid phase fraction decreases in fluidity, and this leads to poor handling properties. In particular, when the liquid compo-nent and liquid phase fraction are supplied to the slurry preparing tank for the feedstock slurry, the slurry ob-tained in then slurry preparing tank decreases in fluidity.
In the solid-liquid separation step (g), it is desir-able to run the solid-liquid separator of settling type at a temperatures of 200-300°C and at a pressure of 20-40 kg/cm2. In this temperature range, the substance consist-ing of heavy product and residues is partly dissolved and the undissolved components (organic components and inor-ganic components) settle rapidly. This means efficient solid-liquid. separation. In this pressure range, the sol-vent for solid-liquid separation remains unevaporated and hence permits stable slurry handling.
There are no :specific restrictions on the petroleum heavy oil containing heavy metals which is used as the feedstock in the present invention. It includes any petro-leum heavy oil containing heavy metals, such as atmospheric residue and vacuum residue. It also includes super-heavy oil containing heavy metals, such as natural bitumen, such as tar sand and oi:l sand. The heavy metals are those which cause catalytic poisoning to the Ni-Mo or Co-Mo catalyst.
They are not specifically restricted. They include, for example, Ni and V.
According to the present invention, hydrocracking of petroleum heavy oi:l is carried out by means of the appara-tus and process flow shown in Fig. 1. A detailed descrip-tion follows.
As shown in Fi.g. 1, the slurry preparing tank (1) is fed with petroleum heavy oil, limonite iron ore (as a cata-lyst), and sulfur (as a cocatalyst). They are made into a feedstock slurry. The feedstock slurry, 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). Hydrocracking of the heavy oil is car-ried out at a reaction pressure of 30-160 kg/cmz, at a reaction temperature of 430-455°C, and for a reaction time of 30-180 minutes. The suspended-bed reactor (3) may be a reactor of bubble column type. [The first reaction step (a)]
The first reaction step (a) gives rise to the reactor effluent containing hydrocracking products. This effluent enters the first gas-liquid separator (4) to be separated into the gas-phase fraction and the liquid phase fraction. The former contains gaseous products and unreacted hydrogen in its high-temperature high-pressure state. The latter contains liquid products, the catalyst used in the first reaction step (a), and heavy metals. [The first gas-liquid separation step (b)]
The gas-phase fraction obtained by separation in the first gas-liquid separation step (b) enters the fixed-bed reactor (5) filled with a Ni-Mo or Co-Mo catalyst, in which hydrocracking takes place at a reaction pressure of 30-160 kg/cm2, at a reaction temperature of 310-380°C (at the start of operation), and at a liquid hourly space velocity of 0.3-2 hr-1. The fixed-bed reactor (5) may be a tubular flow reactor, for example. [The second reaction step (c)]
The reactor effluent (containing hydrotreating products) obtained in the second reaction step (c) enters the second gas-liquid separator (6) for separation into the gas-phase fraction containing unreacted hydrogen and the liquid-phase product. [The second gas-liquid separation step (d)]

The liquid phase product obtained by separation in the second gas-liquid separation step (d) enters the distillation column (7) for fractional distillation to give desired fractions, such as naphtha (C5-171°C), kerosene and gas oil (171-343°C), and vacuum gas oil (343-525°C). [The distillation step (e)] Incidentally, the gas-phase fraction obtained in the second reaction step (c) enters the gas purification step (8) and the purified gas is partly used as a 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 obtained by separation in the first gas-liquid separation step (b) enters the flash separator (9), such as gas-liquid separator, for separation of light products from the liquid phase fraction under lower pressure than that of the first gas-liquid separation step.
[The flash separation step (f)]
The flash separation step (f) gives (after separation from the light products) the liquid phase fraction containing products such as heavy products, residues, the catalyst (used in the first reaction step (a)), and heavy metals. A first part of the liquid phase fraction is fed, together with a solvent for solid-liquid separation, to the solid-liquid separator of settling type (10), so that said first part of the liquid phase fraction is separated into a solid component and a liquid component (both at normal temperature), said solid component containing said residue, the catalyst used in the first reaction step (a), and heavy metals, said liquid component containing said medium product, medium to heavy product, and heavy product. [solid-liquid separation step (g)]
The liquid component obtained by separation in the solid-liquid separation step (g) and the remaining part of the liquid phase fraction other than said first part of the liquid phase fraction is returned to the slurry preparing tank(1) or the suspended-bed reactor (3) in the first reaction step (a).
[The recycling step (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 formation and increasing yields of oil having a low boiling point. The reason for this is that the liquid phase fraction obtained in the second gas-liquid separation step is a highly hydrotreated component having a high capability of hydrotreating.
According to the present invention, the process for hydrocracking of petroleum heavy oil is accomplished by means of the apparatus and process flow shown in Fig. 2. In the case where the heavy oil contains a large amount of quinolines, indoles, benzothiophenes, etc. which are hardly denitrified and desulfurized, it is desirable that the second reaction step (c) and the second gas-liquid separation step that follows in the second reaction step (c) should be followed by a third reaction step. The second gas-liquid separator separates the oil produced in the second reaction step into a liquid phase fraction and a gas-phase fraction, the former containing high-boiling fractions (such as quinolines, indoles, and benzothiophenes), and the latter having a comparatively low boiling point. The gas-phase fraction is denitrified and desulfurized in the third reaction step, so that good quality oil can be obtained.
To be concrete, the third reaction step is carried out as follows. The gas-phase fraction obtained in the second separation step is fed to the fixed-bed reactor 11 filled with a Ni-Mo or Co-Mo catalyst. Hydrotreating in the fixed-bed reactor is carried out at a reaction pressure of 30-160 kg/cm2 (preferably 50-100 kg/cm2), at a reaction temperature of 310-420°C (preferably 340-390°C) at the start of operation, and at a liquid hourly space velocity of 0.3-2 hr-1. (preferably 0.5-1 hr-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 which takes place in third gas-liquid separator 12. The liquid phase fraction is distilled to be separated into desired fractions. The gas-phase fraction is purge gas or recycled to 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 may be returned to the first reaction step so as to reduce coke production to improve yields of :low-boiling oil. The reason for this is that the liquid phase fraction obtained in the second gas-liquid separation atep is a highly hydrotreated component having a high capability of hydrotreating.
EXAMPLES
The present invention will be described with reference to the following examples, which are not intended to re-strict the scope thereof.
Example 1 The first reaction step (a) was carried out as.follows.
In this example, vacuum residue (VR for short) of Middle East origin was used as the petroleum heavy oil containing heavy metals. This VR was incorporated with pulverized limonite iron ore as a catalyst in an amount of 0.6 mass%
(in terms of iron) based on the amount of VR and also with sulfur as a cocata:Lyst in an amount of 1.2 times the amount of iron (1.2 times" in atomic ratio, the amount of iron in limonite iron ore). Thus there was obtained a feedstock slurry. Incidenta:Lly, the VR has the composition of frac-tion as shown in Table 1.
Table 1 Composition of fractions, wt%
on feedstock VR

-171C 171-232C 232-343C 343-450C +450C

Feedstock - - - 4.4 95.6 VR

The feedstock slurry was placed in a 5-liter agitated autoclave. The autoclave was supplied with hydrogen gas for hydrocracking of said VR (vacuum residue or heavy oil).

Hydrocracking was carried out at a reaction pressure of 10 MPa (102 kg/cmz), which is equal to the hydrogen supply pressure, at a reaction temperatura_ of 450°C, and for a reaction time of 60 minutes.
After the reaction was complete, yield of product gas and each liquid fraction are examined. Also, each liquid fraction was examined for the amount of carbon, hydrogen, nitrogen, and sulfur by elemental analysis. The results are shown in Tables 2 to 4.
Table 2 Yield and composition, wt%
on feedstock VR

Fraction-1Fraction-2Fraction+450C

Gas -171C 171-343C343-450CHexaneAsphal-Coke solubletene Liquid 6.fi8.4 28.7 18.5 27.4 9.7 1.5 product In Table 2, the yield of gas, liquid fractions, and coke (collectively referred to as fractions etc. hereinaf-ter) is expressed in terms of ratio (mass%) of the amount of fractions etc. to the amount of VR supplied. This defi-nition applies to the yield in Tab:Les 5, 6, 8, 10, 13, 14, 16, 18, and 19 given later.
It is apparent: from Tables 1 and 2 that the amount of fractions (or distillate of vacuum distillation) having a boiling point lower than 450°C increased from 4.4 wt% to 55.6 wt% [= 8.4 wt% (fraction-1) + 28.7 wt% (fraction-2) +
18.5 wt% (fraction-3)), by hydrocracking while keeping the yield of coke at 1.5 wt%. Incidentally, wt% (or mass%) will be expressed simply as % hereinafter.

Table 3 Content Atomic Aromat-of ratio elements"
(wt%) C H N S (H/C) icily (fa) Feedstock 84.5 10.2 0.3 4.3 1.4 0.36 VR

Fraction-184.8 14.4 0.2 0.2 2.0 0.04 Liquid Fraction-2E14.5 11.9 0.3 1.9 1.7 0.25 product Fraction-385.6 10.9 0.3 2.8 1.5 0.33 It is apparent. from Table 3 that hydrocracking in-creased the H/C rat:io from 1.4 in the feedstock VR to 2.0 in fractin-1 (with a boiling point lower than 171°C) and 1.7 in fractin-2 (with a boiling point of 171-343°C). It is also noted that the sulfur content in fraction-1, frac-lion-2, and fraction 3 was 0.2%, 1.95%, and 2.8%, respec-lively. This suggests that hydrocracking performed desul-furization to some extent. By contract, hydrocracking performed denitrification only a little.
Table 4 Content ppm of metallic element, V' Ni Na Ca Mg Fe Feedstock 120 90 170 20 <10 10 VR

Liquid 450C <'1 <1 <1 <1 <1 <1 -productTHFI 3200 840 860 270 550 394000 It is noted from Table 4 that metallic elements in the feedstock VR were not detected in all the distillates but were detected in THFI. This suggests that metallic ele-ments were removed by adsorption on the suspending catalyst.
Incidentally, THFI denotes those components which are in-soluble in tetrahydrofuran.

Example 2 Hydrocra eking of the first reaction step (a) was per-formed under the same condition as in Example 1 on the heavy fraction (or bottom or distillation residue having a boiling point higher than 450°C) obtained in Example l, which had been charged together with the feedstock VR.
(The amount of the heavy fraction was 80% of the amount of feedstock VR.) Inc:identally, the amount of pulverized limonite iron ore as the catalyst was 0.6 mass% (in terms of iron) of the amount of feedstock VR, and the amount of sulfur as the cocat:alyst was 1.2 times the amount of iron, as in the case of Example 1.
After the reaction was complete, yield of each frac-tion of gas .and liquid was examined. The results are shown in Table 5.
Table 5 Yield and composition, wt% on feedstock VR
Fraction-1 Fraction-2 Fraction -3 +450°C
Gas -.~71°C 171-343°C 343-450°C Hexane Asphal- Coke soluble tene Product 9.6 9.0 36.8 41.7 1.6 0.7 2.4 It is noted from Tables 2 and 5 that bottom recycling greatly decreased t:he amount of +450°C fraction. In Exam-ple 1 (without bottom recycling), the amount of fraction remaining as residue of vacuum distillation (or the amount of +450°C fraction or the amount of fraction having a boil-ing point higher than 450°C) was about 40% (= 27.4% + 9.7%
+ 1.5%), whereas in Example 2 (with bottom recycling), the amount of +450°C fraction was about: 5% (= 1.6% + 0.7% +
2.4%) .
In add ition, bottom recycling raised the yield of fractions having a boiling point lower than 450°C from about 55% to about 90%. This suggests that a large portion of feedstock VR can be converted into light distillable fractions.
Despite conversion into light oil in such high yields, the yield of coke was only 2%.
Incidentally, the distilled oil was identical in prop-erties (amount of <:, H, N, S, and metallic elements, and H/C ratio) with that in Example 1.
Example 3 The oil fraction (having a boiling point lower than 450°C) obtained in the first reaction step (a) in Examples 1 and 2 corresponds to the gas-phase fraction obtained by separation in the i:irst gas-liquid separation step (b) placed downstream t:he first reaction step (a), according to the process for hydrocracking of petroleum heavy oil as defined in the present invention.
The distilled oil fraction obtained in the first reac-tion step (a) in Example 2 was fed to a fixed-bed reaction column filled with Ni-Mo/A1z03 catalyst (Ni-Mo catalyst supported on A1203). Hydrotreating was carried out at a reaction temperature of 350°C (at the start of reaction), at a reaction pressure of 10 MPa, and at a liquid hourly space velocity (LHSV) of 1 hr-1. [The second reaction step (c)] After the reaction, yields of the gas and liquid fractions evolved were examined. They were also examined for the amount of C, H, N, and S by elemental analysis.
The results are shown in Tables 6 and 7.
Table 6 Yield and composition, wt%
on feedstock VR

Naphtha3<;eroseneGas oil Vacuum gas oil Gas AsphalteneCoke product13.428.3 15.5 37.4 4.4 0.8 2.4 Table 7 Content Atomic Aromat-of ratio elements C, wt% H, wt% N, S, (H/C) icity ppm ppm (fa) Feedstock 84.5 10.2 3000 43000 1.4 0.36 VR

Naphtha 85.2 14.0 <10 60 2.0 0.09 Liquid KeroseneF35.7 12.9 <10 30 1.8 0.17 product Gas oil 87.3 12.3 <10 90 1.7 0.22 In Example 2, the recycling of bottom having a boiling point higher than 450°C (remaining after vacuum distilla-tion) increased this yield of fractions having a boiling point lower than 4'.>0°C up to about 90%. By contrast, in Example 3, this fr<~ction undergoes hydrotreating to a high degree by catalytic: reaction in the fixed-bed reactor in the second reaction step (c). Therefore, the yield of the vacuum gas oil having a boiling point higher than 343°C (as the final product oil) decreased to about 4% and the yield of the light oil having a boiling point lower than 343°C
reached about 80%, as shown in Table 6.

Moreover, the fraction of light oil having a boiling point lower 'than 343°C contains less than 10 ppm of nitro-gen; and naphtha, kerosene, and gas oil contain sulfur in an amount of 60 ppm, 30 ppm, and 90 ppm, respectively, as shown in Tab:Le 7. This suggests that the process was capa-ble of denit:rificat:ion and desulfurization to a high degree.
A probable re=_ason for this is that hydrocracking at a high temperature (450°C) in the first reaction step (a) brings about dealky:Lation which removes steric hindrance from those componf~nts which present difficulties in desulfuriza-tion and den:itrification.
Example 4 and Comparative Example 1 The feedstock VR was incorporated with an iron cata-lyst (of dif:Eerent kind) and sulfur (as a cocatalyst). The amount of su:Lfur wa.s 1.2 times the amount of iron. The resulting slurry wa.s fed to a suspended-bed reactor for hydrocracking at a reaction temperature of 430-450°C, at a reaction pressure of 10 MPa (under a hydrogen gas atmos-phere), and :Eor a reaction time of 60 minutes. The reac-tion product was analyzed for coke yield (or yield of tolu-ene insolublsss) so as to examine the relation among the kind of iron catalyst, reaction temperature, and coke yield.
The results are shown in Table 8.

Table 8 Coke yield, wt%, based on feedstock VR

Without With bottom bottom oil oil rec clin recycling 430C 440C 445C 450C __450C

Natural pyrite- - 5.2 - -(1 ) Natural pyrite- - 7.4 - -(2) Limonite 1.1 - - 1.5 2.4 iron ore The decomposition of heavy oil (such as vacuum distil-lation residue) is conventionally accomplished by feeding the distillate to fixed-bed catalytic reaction or fluid-ized-bed catalyst reaction (FCC) through the coker process.
The coke yield in t:he coker process is as high as 20-50%.
Therefore, t:he effective use of heavy oil by this process was difficult.
By contrast, in the case of hydrocracking by a sus-pended-bed reactor, the coke yield is low as mentioned above.
It is apparent form Table 8 that the coke yield varies depending on the kind of iron catalyst and the reaction temperature. In the case where the amount of natural py-rite is 1% of the amount of feedstock VR, the coke yield was 5.2% or '7.4%. By contrast, in the case where the limo-rite iron ore (as the catalyst in the first reaction step (a) according to the present invention) is used in an amount of 0.6% (in terms of iron) of the amount of feed-stock VR, the coke yield is further low, and it is 1% at a reaction tem;peratuz:e of 430°C and 1.5% at a reaction tem-perature of 450°C. It was found that in the case where limonite iron ore is used as the catalyst, the coke yield can be kept .at 2.4% even though the recycling of bottom oil is performed.
The limonite iron ore used as the catalyst in the first reaction step (a) according to the present invention is inexpensive and yet capable of keeping low the amount of coke formation. The reason for this is that limonite iron ore is converted into pyrrhotite even at a low temperature (about 250°C) and pyrrhotite exhibits the catalytic activ-ity for hydrotreati.ng. This means that the catalyst exists in the early stage of pyrolysis of heavy oil and it pro-motes the hydrotreating reaction, while suppressing the coke formation due to polycondensation of radicals formed by thermal decomposition.
Example 5 The example demonstrates hydrocracking of petroleum heavy oil containing heavy metals that employs the appara-tus and process flow shown in Fig. 1. (The apparatus and process flow shown in Fig. 1 have been explained above.) VR used for hydrocracking is petroleum heavy oil con-taining heavy metals. The amount of limonite iron ore was 0.6 mass% of the amount of VR and the amount of sulfur (as the cocatalyst) was 1.2 times the amount of iron. Reaction in the first reaction step (a) was carried out at a reac-tion pressure of 10 MPa, at a reaction temperature of 450°C, and for a re,sction time of 60 minutes. The catalyst in the second reaction step (c) was Ni-Mo catalyst, and reaction in the second reaction step (c) was carried out at 350°C
(at the start of processing), at a reaction pressure of 10 MPa, and at a liquid hourly space velocity (LHSV) of 1 hr-1.
The recycling step (h) was carried out in which the liquid component [9] and the residue of the liquid phase fraction [7] were returned to the slurry preparing tank in the first reaction step (a).
The first reaction step (a) to the recycling step (h) were repeated continuously for several months.
Stable, continuous operation was possible without any trouble (suc:h as decrease in catalytic activity and clog-ging of catalyst layer), with the amount of coke production being 2%.
Operation gave light oil (with a high degree of hydro-cracking) in high yields as follows. The yield of vacuum light oil (having a boiling point higher than 343°C) was about 4%, and the yield of light oil (having a boiling point lower than 343°C) was about 80%. The content of N in the light oil fraction having a boiling point higher than 343°C was less than. 10 ppm, and the content of S in naphtha, kerosene, and gas oil was 60 ppm, 30 ppm, and 90 ppm, re-spectively. This means that the operation accomplished denitrification and desulfurization to a great extent.
Hydrocracking of petroleum heavy oil according to this example permits prolonged continuous operation with a mini-mum of coke accumu.Lation and gives light oil economically in high yields, with a high degree of desulfurization and denitrification and also a high degree of hydrocracking.
Example 6 The first reaction step (a) was carried out as follows.
Vacuum reside of petroleum of South America origin was used as the petroleum heavy oil containing heavy metals (refer-red to as VR-B hereinafter). Pulverized limonite iron ore (as the catalyst) was added in an amount of 1 wt~ of the amount of VR-B. Sulfur (as the cocatalyst) was added in an amount of 1.2 times the amount of iron (or 1.2 times (in atomic ratio) the amount of iron in limonite iron ore). In this way there was obtained the feedstock slurry. When distilled, VR-B gave fractions as shown in Table 9. Inci-dentally, VR-B was found to contain quinolines, indoles and benzothiophenes in larger quantities than VR.
Table 9 Composition of fraction, wt% on feedstock VR-B
-171°I:, 171-232°C 232-343°C 343-450°C
+450°C
Feedstock VR-B - - 0.33 l 6.67 93.00 The above-mentioned feedstock slurry was placed in a 5-liter agitated autoclave. The autoclave was supplied with hydrogen gas for hydrocracking of said VR-B. Hydro-cracking was carried out at a reaction pressure of 100 kg/cm2 (9.8 IHPa), which is equal to the hydrogen supply pressure, at a reaction temperature of 450°C, and for a reaction time of 60 minutes.
After the reaction was complete, the yield of product gas and each. liquid fraction are examined. Also, the each liquid fraction was examined for the amount of carbon, hydrogen, nitrogen, and sulfur by elemental analysis. The results are chown in Tables 10 to 12. In Table 10, the yield of gas, fractions, and coke is expressed in terms of wt% of the amount of VR-B supplied. This definition ap-plies to the yield in Tables 13, 14, 16 and 18 given later.
Table 10 Yield and composition, wt%
on feedstock VR-B

Fraction-1Fraction-2Fraction+450C

Gas-.t71C 171-343C343-450CHexaneAsphal-Coke solubletene Product4.40'10.97 27.26 16.96 30.87 8.98 2'08 ~ ~

It is apparent from Tables 9 and 10 that hydrocracking increased th~a amount of fractions having a boiling point lower than 4.'50°C increased from 7.00 wt% to 55.19 wt% (_ 10.97 + 27.25 + 16.96) by hydrocracking, while keeping the yield of coke at 2.08 wt%.
Table 11 Content Atomic Aromat-of ratio elements, (wt%) C H N g (H/C) icity (fa) Feedstock 85.50 10.10 0.93 0.90 1.402 0.38 VR-ES

Fraction-184.74 13.97 0.05 0.24 1.965 0.11 Liquid Fraction-286.02 12.39 0.39 0.48 1.717 0.24 product Fraction-3815.3511.46 1.05 0.52 1.582 0.32 It is apparent from Table 11 that hydrocracking in-creased the :11/C ratio from 1.4 in the feedstock VR-B to 1.97 in fractin-1 and 1.72 in fractin-2. It is also noted that the sulfur content in fractions-1, -2, and -3 was 0.24%, 0.48%, and 0.52%, respectively. This suggests that r-.

hydrocracking performed desulfurization to some extent. By contract, hydrocrac:king performed denitrification only a little.
Table 12 Content of metallic elements, ppm V Ni Na Ca Mg Fe Feedstock 5:i 65 17 36 <5 163 VR-B

Liquid 450C <~~ <1 <1 <1 <1 <1 -productTHFI 830 650 150 130 155 160000 It is noted from Table 12 that metallic elements in the feedstoc:k VR-B were not detected in all the fractions but were detected i.n THFI. This suggests that metallic elements were removed by adsorption on the suspending cata-lyst.
Example 7 Hydrocra eking of the first reaction step (a) was per-formed under the same condition as in Example 6 on the heavy gravity fraction (or bottom or distillation residue having a boiling point higher than 450°C) obtained in Exam-ple 6, which had been charged together with the feedstock VR-B. (The amount of the heavy gravity component was 74.84% of the amount of feedstock VR-B.) The amount of limonite was 1 wt% (in terms of Fe), and the S/Fe ratio was 1.2.
The yie:Ld of products is shown in Table 13. It is noted from Tables 10 and 13 that bottom recycling greatly decreased the amount of +450°C fraction. In Example 6 (without bottom rec:ycling), the amount of residues was about 41.93 wt% (30.87 + 8.98 + 2.08), whereas in Example 7 (with bottom oil recycling), the amount of residue was 8.62 wt% (5.4 + 1.3 + 1.92).
In addition, bottom oil recycling raised the yield of fractions having a boiling point lower than 450°C from 55.13 wt% to 85.25 wt%. This suggests that a large portion of feedstock vR-B can be converted into light distillable fractions. Despite conversion into light oil in such high yields, the yield of coke was kept lower than 2 wt%.
Incidentally, the distilled oil was identical in prop-erties with that in Example 5.
Table 13 Yield and composition, wt%
on feedstock VR-B

_ Fraction-1Fraction-2Fraction_ -3 +450C

Gas HexaneAsphal--171 171-343C343-450C Coke C

solubletene liquid 7.7214.85 32.87 37.53 5.40 1.30 1.92 product ~ ~ ~

Example 8 The reactor effluent obtained from the first reaction step (a) in Example 7 underwent gas-liquid separation (the first gas-liquid separation step) such that the components having a boiling point lower than 450°C was contained in the gas phase. Thus there was obtained gas phase fraction.
This gas phase fra<:tion was passed through a fixed-bed reactor column filled with Ni-Mo/A1z03 catalyst, and hy-drotreating was carried out at a reaction temperature of 350°C (at the start. of reaction), at a reaction pressure of 100 kg/cm2 (9.8 MPa.), and at a liquid hourly space velocity of 1 hr-1. (The second reaction step (c)) Table 14 shows; the yields of the gas and liquid prod-ucts, and Table 15 shows the results of elemental analysis.
Table 14 Yield and composition, wt%
on feedstock VR-B

Fraction-1Fraction-2Fraction -3 +450C

Gas -171C 171-343C343-450C HexaneAsphal-Coke soluble tens Product8.6113.04 37.24 34.34 5.40 1.30 1.92 ~ ~

Table 15 Content Atomic Aromat-of ratio elements C (wt%)H (wt%)N (ppm)S (ppm)(HOC) icily (fa) Feedstock 85.50 10.10 9300 9000 1.402 0.38 Fraction-184.74 15.23 200 60 2.156 0.07 Liquid Fraction-287.00 12.89 960 100 1.780 0.15 product Fraction-387.27 12.36 3000 700 1.701 0.24 It is noted that the second reaction step promoted hydrotreating. Thos yield of fraction-3 decreased from 37.53 wt~ to 34.34 wt~. The concentration of nitrogen was 200 ppm in fraction-l, 960 ppm in fraction-2, and 3000 ppm in fraction-3. ThES concentration of sulfur was 60 ppm in fraction-1, 100 ppm in fraction-2, and 700 ppm in fraction-3. In other words" the concentration of nitrogen and sul-fur in each fraction was higher than in Example 3 in which the feedstock was VR. This is because as compared with VR, VR-B contains more quinolines, indoles, and benzothiophenes which are hardly denitrified and desulfurized.
Example 9 The reactor effluent obtained from the second reaction step (c) in Examples 8 underwent gas-liquid separation (the second gas-liquid :separation step) such that the component having a boiling point lower than :343°C was contained in the gas phase. Thus there was obtained agas phase fraction.
This gas phase fraction was passed through a fixed-bed reactor column fil:Led with Ni-Mo/A1203 catalyst, and hy-drotreating 'was carried out at a reaction temperature of 380°C, at a reaction pressure of 9.8 MPs, and at a liquid hourly space veloc_Lty of 1 hr~i. (The third reaction step (c)) Table 16 shows the yields of the gas and liquid prod-uct, and Table 17 shows the results of elemental analysis.
Table 16 Yield and composition, wt%
on feedstock VR-B

+450C
Fraction-1Fraction-2Fraction HexaneAsphal-Gas _171C 171-343C343-450C Coke solubletene Product9.1118.63 31.45 34.34 5.40 1.30 1.92 ~ ~ ~

Table 17 Content Atomic Aromat-of ratio elements C (wt%)H (wM/o) N S (ppm)(H/C) icily (ppm) (fa) Feedstock 85.50 10.10 9300 9000 1.402 0.38 Fraction-184.32 15.51 <10 30 2.209 0.06 Product Fraction-286.85 13.12 <10 50 1.814 0.13 It is noted that the third reaction step promoted hydrotreating. The yield of fraction-2 (171-343°C) de-creased from 37.24 wt% to 31.45 wt%. Conversely, the yield of fraction-1 (-17:L°C) increased from 13.04 to 18.63 wt%.
As shown in Table 17, the nitrogen content was lower than 10 ppm in light oil fraction having a boiling point lower than 343°C. The content of sulfur was 30 ppm and 50 ppm in fraction-1 and fraction-2, respectively. This sug-gests that the third reaction step achieves a high degree of denitrification and desulfurization.
Example 10 The first: reaction step (a) was carried out as follows.
VR-B was used as the petroleum heavy oil containing heavy metals. Pulverized limonite iron ore (as the catalyst) was added in an amount of 1 wt~ of the amount of VR-B. Sulfur (as the cocatalyst) was added in an amount of 1.2 times the amount of iron (or 1.2 times (in atomic ratio) the amount of iron in limonite iron ore). In this way there was ob-tained the feedstoc:k slurry.
The above-mentioned feedstock slurry was placed in a 5-liter agitated autoclave. The autoclave was supplied with hydrogen gas f:or hydrocracking of said VR-B. Hydro-cracking was carried out at a reaction pressure of 50 or 30 kg/cm2, which is equal to the hydrogen supply pressure, at a reaction t.=_mperat:ure of 450°C, and for a reaction time of 60 minutes.
After the reaction was complete, the yield of product gas and each liquid fraction are examined. The results are shown in Table 18. It is noted from Table 18 that hydro-cracking gave a high oil yield even though the reaction pressure is 50 kg/c:m2 or 30 kg/cm2. However, the coke yield slightly increased more than in the case where the reaction pressure is 100 kg/cm2. When the reaction pres-sure was 30 :kg/cmz, the oil yield was 66.77 wt% (11.12 +
25.07 + 30.58) Table 18 Yield and composition, wt%
on feedstock VR-B

Product Fraction-1Fraction-2Fraction -3 Gas . +450C
' i Hexane Asphal--171 171-343343-450 Coke C C C tene soluble (100 4.4010.97 27.26 16.96 30.87 8.98 2.08 k) (50k) 4.2911.42 25.92 26.42 19.00 10.55 4.21 (30k) 3.9911.12 25.07 30.58 17.89 7.94 4.98 Example 11 The fir st react:ion step (a) was carried out as follows.
VR-B was used as the petroleum heavy oil containing heavy metals. Pulverized limonite iron ore (as the catalyst) was added in an amount of 1 wt% of the amount of VR-B. Sulfur (as the cocatalyst) was added in an amount of 1.2 times the amount of iron (or 1.2 times (in atomic ratio) the amount of iron in limonite iron ore). To VR-B was added the liq-uid phase fraction (obtained in the second gas-liquid sepa-ration step in Example 9) in an amount of 20 wt% of VR-B.
In this way there was obtained the feedstock slurry.
The above-mentioned feedstock slurry was placed in a 5-liter agitated autoclave. The autoclave was supplied with hydrogen gas for hydrocracking of said VR-B. Hydro-cracking was carried out at a reaction pressure of 30 kg/cm2, which is equal to the hydrogen supply pressure, at a reaction temperature of 450°C, and for a reaction time of 60 minutes.

After the reaction was complete, the yield of product gas and each liquid fraction are examined. The results are shown in Table 19. It is noted from Table 19 that effi-cient hydroc:rackincf was accomplished, with a decreased yield of coke and an increased yield of low-boiling oil, even at a reaction pressure of 30 kg/cm2 if the liquid phase fraction from the second gas.-liquid separator is returned to the first reaction step.
Table 19 Yield and composition, wt%
on feedstock VR-B

i ti +450C

Gas Fraction-1ract rac HexaneAsphal-Coke -_171C on- on - solubletene Product*4.1113.13 26.12 32.14 16.857.11 1.74 Product3.9911.12 25.07 30.58 17.697.94 4.98 * added to the liquid-phase from the second gas-liquid separator.
[Effect of the invention] The present invention provides a process for hydroc:racking of petroleum heavy oil containing heavy metals. This process permits vacuum distillation residue (as an example of petroleum heavy oil containing heavy metals) to be converted into light oil by hydrocrack-ing with a minimum of coke accumulation for prolonged con-tinuous operation. It gives desulfurized and denitrified light oil with a high degree of hydrocracking.

Claims (12)

1. A process for hydrocracking of petroleum heavy oil containing heavy metals, said process comprising the following steps (a) to (h):
(a) a first reaction step of feeding a suspended-bed reactor with a feedstock slurry composed of petroleum heavy oil containing heavy metals, limonite iron ore, and sulfur, together with hydrogen gas, thereby performing hydrocracking on said heavy oil at a reaction pressure of 30-160 kg/cm2, at a reaction temperature of 430-450°C, and for a reaction time of 30-180 minutes;
(b) a first gas-liquid separation step of separating the reactor effluent, which has been obtained in said first reaction step (a), into a gas phase fraction and a liquid phase fraction;
(c) a second reaction step of feeding a fixed-bed reactor filled with a Ni-Mo or Co-Mo catalyst with the gas phase fraction obtained from said first gas-liquid separation step, thereby performing hydrotreating at a reaction pressure of 30-160 kg/cm2, at a reaction temperature of 310-380°C at the start of operation, and at a liquid hourly space velocity of 0.3-2 hr-1;
(d) a second gas-liquid separation step of separating the reactor effluent, which has been obtained in said second reaction step (c), into a gas phase fraction and a liquid phase fraction;
(e) a distillation step of separating by distillation the liquid phase fraction, which has been obtained by separation from said second gas-liquid separation step (d), into prescribed fractions;
(f) a flash separation step of performing flash separation under reduced pressure on the liquid phase fraction, which has been obtained by separation from said first gas-liquid separation step (b), such that the light reaction product in the liquid phase fraction moves to the gas phase;
(g) a solid-liquid separation step of feeding a solid-liquid separator of settling type with part of the liquid phase fraction containing medium product, medium to heavy product, heavy product, residue, heavy metals, and the limonite iron ore used in said first reaction step (a), which have been obtained by separation from the light product in said flash separation step (f), and a solvent for solid-liquid separation, thereby separating them into a solid component and a liquid component, both at normal temperatures, said solid component containing said residue, the catalyst used in said first reaction step (a), and heavy metals, said liquid component containing said medium product, medium to heavy product, and heavy product;
(h) a recycling step of returning to said 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 product, medium to heavy product, heavy product, and residue, heavy metals, and catalyst used in said first reaction step (a) which have been obtained by separation from the light product in said flash separation step (f), as the remainder left after feeding to said solid-liquid separator of settling type.

2. A process for hydrocracking of petroleum heavy oil as defined in claim 1, wherein the reaction in the bed reactor in the first reaction step (a) is carried out at a reaction pressure of 50-100 kg/cm2, at a reaction temperature of 440-450°C, and for a reaction time of 60-120 minutes.

3. A process for hydrocracking of petroleum heavy oil as defined in claim 1 or 2, wherein limonite iron ore as the catalyst in the first reaction step (a) is added in an amount of 0.3-2 mass% in terms of iron of the amount of petroleum heavy oil and sulfur as the cocatalyst is added in an amount, in atomic ratio, of 1-3 times the amount of iron contained in limonite iron ore.

4. A process for hydrocracking of petroleum heavy oil as defined in any one of claims 1 to 3, wherein limonite iron ore in the first reaction step (a) is used in the form of fine powder having an average particle diameter smaller than 2 µm which is prepared by mechanically pulverizing in a petroleum solvent.

5. A process for hydrocracking of petroleum heavy oil as defined in any one of claims 1 to 4, wherein said limonite iron ore in the first reaction step (a) is one which contains substantially no iron oxide.

6. A process for hydrocracking of petroleum heavy oil as defined in any one of claims 1 to 5, wherein the total amount of the liquid component, heavy product, and residue in liquid phase fraction is 10-130 mass% of the amount of petroleum heavy oil supplied to the first reaction step (a).

7. A process for hydrocracking of petroleum heavy oil as defined in any one of claims 1 to 6, wherein the solid-liquid separator in the solid-liquid separation step (g) is operated at a temperature of 200-300°C and at a pressure of 20-40 kg/cm2.

8. A process for hydrocracking of petroleum heavy oil as defined in any one of claims 1 to 7, wherein reaction in the fixed-bed reactor in the second reaction step (c) is carried out at a reaction pressure of 50-100 kg/cm2, at a reaction temperature of 330-360°C at the start of operation, and at a liquid hourly space velocity of 0.5-1 hr-1.

9. A process for hydrocracking of petroleum heavy oil as defined in any one of claims 1 to 8, which further comprises the following steps:
(i) a third reaction step in which the gas-phase fraction obtained in the second gas-liquid separation step (d) undergoes hydrocracking by the aid of a Ni-Mo or Co-Mo catalyst at a reaction pressure of 30-160 kg/cm2, at a reaction temperature of 310-420°C at the start of operation, and at a liquid hourly space velocity of 0.3-2 hr-1;

(ii) a third gas-liquid separation step in which the reactor effluent obtained in the third reaction step is separated into the gas phase and the liquid phase; and (iii) a distillation step in which the liquid phase fraction obtained by separation in the third gas-liquid separation step is distilled.

10. A process for hydrocracking of petroleum heavy oil as defined in claim 9, wherein reaction in the third reaction step is carried out at a reaction pressure of 50-100 kg/cm2, at a reaction temperature of 340-390°C (at the start of operation), and at a liquid hourly space velocity of 0.5-1 hr-1.

11. A process for hydrocracking of petroleum heavy oil as defined in any one of claims 1 to 9, which further comprises the following step:
(i) a recycling step in which the liquid phase fraction obtained in the second separation step is returned to the first reaction step (a).

12. A process for hydrocracking of petroleum heavy oil as defined in any one of claims 1 to 4, wherein said limonite iron ore in the first reaction step (a) contains less than 10 mass% of iron oxide which can be detected by powder x-ray diffraction.