EA040018B1 - CRUDE OIL REFINING METHOD - Google Patents

Description

The present invention relates to a process for processing crude oil.

It has previously been described that crude oil refining can be integrated with downstream chemical plants, such as a pyrolysis/steam cracker, to enhance the formation of high value chemicals through the formation of fuels.

U.S. Pat. No. 3,702,292 describes a scheme for a combined fuel and chemical refinery including a crude distillation facility, a hydrocracking facility, a delayed coking facility, a reformer, an ethylene and propylene production facility, including a pyrolysis/steam cracking unit and a separation unit. pyrolysis products, a catalytic cracking agent, an aromatic product recovery agent, a butadiene recovery agent, and an alkylation agent in an interconnected system to achieve a conversion of crude oil to petrochemicals of about 50% and a conversion of crude oil to fuels of about 50%.

A major disadvantage of conventional means and methods for combining refinery operations with downstream chemical plants for the production of petrochemicals is that such combined processes still produce significant amounts of fuel. In addition, conventional means and methods for combining refinery operations with downstream chemical units have a relatively low propylene yield, based on the weight percent of crude oil.

It is an object of the present invention to provide means and methods for combining refinery operations with downstream chemical plants that have an increased production of petrochemicals from the production of fuels. Another object of the present invention is to provide means and methods for combining refinery operations with downstream olefin synthesis chemical units.

The solution to this problem is achieved in the embodiments of the invention described below in this document and characterized in the claims.

In one aspect, the present invention relates to a process for processing crude oil. This method is also shown in Fig. 1-3, which are further described below.

Accordingly, the present invention provides a process for refining crude oil, comprising the steps of:

(a) subjecting the crude oil to the distillation of the crude oil to form a gas fraction, a bottoms residue and at least one of kerosene or gas oil;

(b) separating the VAT residue into vacuum gas oil and vacuum residue;

(c) subjecting the vacuum gas oil to hydrocracking the vacuum gas oil to form at least liquefied petroleum gas (LPG);

(d) subjecting the vacuum residue to hydrocracking the vacuum residue to form a heavy distillate, liquefied petroleum gas (LPG) and a middle distillate;

(e) subjecting the heavy distillate and/or middle distillate obtained in step (d) to hydrocracking to form liquefied petroleum gas (LPG); and (f) subjecting the LPG obtained in one or more of steps (b) to (e) to olefin synthesis.

In a preferred embodiment of the process according to the present invention, the synthesis of olefins comprises steam cracking and/or catalytic dehydrogenation.

In a preferred embodiment of the process according to the present invention, steam cracking comprises the steps of:

separating C2-C4 paraffins from said gaseous stream, further separating C2-C4 paraffins into individual streams, each of the streams predominantly containing C2 paraffins, C3 paraffins and C4 paraffins, respectively; and each individual stream is fed into its designated furnace section of said steam cracker.

In a preferred embodiment of the process according to the present invention, the reaction conditions in step (e) comprise a temperature of 200 to 600° C., a pressure of 1 to 12 MPa and a hydrogen content of 5 to 30 wt.% relative to the mass of hydrocarbons in the reaction.

In a preferred embodiment, the method according to the present invention further comprises the steps of:

separating the gas fraction into a stream containing LPG and an off-gas stream containing methane and/or hydrogen;

supplying an off-gas stream to the steam cracker to provide fuel for heating the furnace of the steam cracker.

In a preferred embodiment of the process according to the invention, steps (c) and (e) are carried out in the same hydrocracking unit.

The present invention also provides a crude oil refining process comprising the steps of:

(a) subjecting the crude oil to the distillation of the crude oil to form a gas fraction, a bottoms

- 1 040018 tatka and at least one of kerosene or gas oil;

(b) separating the distillation residue into vacuum gas oil and vacuum residue;

(c) subjecting the vacuum gas oil to hydrocracking the vacuum gas oil to form liquefied petroleum gas (LPG);

(d) subjecting the vacuum residue to hydrocracking the vacuum residue to form additional vacuum gas oil, liquefied petroleum gas (LPG) and middle distillate; and (e) subjecting the additional vacuum gas oil to vacuum gas oil hydrocracking.

In a preferred embodiment, the process of the present invention further comprises contacting the middle distillate with hydrogen in the presence of a catalyst under reaction conditions sufficient to form LPG.

In a preferred embodiment of the process according to the present invention, the reaction conditions include a temperature of 200 to 600° C., a pressure of 1 to 12 MPa and a hydrogen content of 5 to 30 wt.% relative to the mass of hydrocarbons.

In a preferred embodiment of the process according to the invention, the catalyst in the contacting step contains a mixture of Ni, W and Mo and one or more of Pd, Rh, Ru, Ir, Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W and V in the form of metal or metal sulfide.

In a preferred embodiment, in the process according to the present invention, the LPG produced by vacuum gas oil hydrocracking, vacuum residue hydrocracking and/or said contacting step is further subjected to olefin synthesis.

In a preferred embodiment of the process according to the present invention, the synthesis of olefins comprises steam cracking and/or catalytic dehydrogenation.

Traditionally, petrochemical products such as propylene are obtained by subjecting crude oil to a distillation process and subjecting certain fractions of the crude oil thus obtained to a refining process. In the context of the present invention, it has been found that the propylene yield of a process for converting crude oil to petrochemicals can be increased by subjecting one or more of the naphtha, kerosene, and gas oil fractions to aromatic ring opening, and then converting the LPG formed from aromatic ring opening to olefins, as compared to subjecting the same crude oil fractions directly to steam cracking. As used herein, the term propylene yield refers to the wt % of propylene obtained from the total weight of crude oil. In addition, due to the initial conversion of hydrocarbon feedstock in the LPG before being subjected to olefin synthesis, the fractionation section downstream of the olefin synthesis unit can be much less complex, since significantly fewer by-products are formed, such as soot oil or cracked distillate, resulting in a significant reduction in capital costs.

Accordingly, the term kerosene and/or gas oil produced by distillation of crude oil in said process means that said one or more fractions of kerosene and gas oil are produced by the crude oil distillation process step of the combined process of the present invention. In addition, the term refinery-derived middle distillate produced in said process means that said refinery-derived middle distillate is produced by a refinery process step of the combined process of the present invention.

The prior art describes methods for the production of petrochemicals from certain hydrocarbon feedstocks, such as certain fractions of crude oil and/or distillates obtained from a refinery.

US 3,839,484 describes a process for producing unsaturated hydrocarbons by pyrolysis of naphtha in a pyrolysis oven, comprising hydrocracking said naphtha to form a mixture of paraffins and isoparaffins, said mixture consisting essentially of hydrocarbons containing from 1 to about 7 carbon atoms per molecule, and pyrolysis of the resulting mixture of paraffins and isoparaffins in the specified pyrolysis furnace. US Pat. No. 3,839,484 also discloses that diesel and distillate fuels boiling in the range of 400-650°F (about 204-343°C) are not further processed.

The term crude oil as used herein refers to crude oil extracted from geological formations. The term crude oil should also be understood to include crude oil that has undergone oil-water separation and/or oil-gas separation and/or desalting and/or stabilization. Any crude oil is suitable as a starting material for the method of this invention, including Arabian heavy (Arabian Heavy), Arabian light (Arabian Light), other grades of Persian Gulf oil, Brent (Brent), North Sea oil, North African and West African oil, Indonesian oil, Chinese oil and mixtures thereof, but also shale oil, tar sands, gas condensates and bio-derived oils. The crude oil used as feedstock for the process of the present invention is preferably a conventional oil having an API gravity greater than 20 as determined by ASTM D287. More preferably, the crude oil used in the process of the present invention is a light crude oil having an API gravity greater than 30. Most preferably, the crude oil used in the process of the present invention contains Arabian light crude oil. Arabian light crude oil typically has an API gravity in the range of 32-36°

API and sulfur content in the range of 1.5-4.5 wt%.

The term petrochemicals as used herein refers to chemicals derived from crude oil that are not used as fuel. Petrochemicals include olefins and aromatics, which are used as feedstock for the production of chemicals and polymers. High value petrochemicals include olefins and aromatics. Typical high value olefins include, but are not limited to, ethylene, propylene, butadiene, butylene-1, isobutylene, isoprene, cyclopentadiene, and styrene. Typical high value aromatics include, but are not limited to, benzene, toluene, xylene, and ethylbenzene.

The term fuel as used herein refers to products derived from crude oil used as an energy carrier. Unlike petrochemicals, which are a collection of well-defined compounds, fuels are usually complex mixtures of various hydrocarbon compounds. Fuels typically obtained from refineries include, but are not limited to, gasoline, jet fuel, diesel fuel, fuel oil, and petroleum coke.

The term crude oil distillation gases or gas fraction as used herein refers to the fraction obtained from the distillation of crude oil that is gaseous at ambient temperature. Accordingly, the gas fraction obtained by distillation of crude oil contains mainly C1-C4 hydrocarbons and may further contain impurities such as hydrogen sulfide and carbon dioxide. In this description, other oil fractions obtained by distillation of crude oil are referred to as naphtha, kerosene, gas oil and still bottoms. The terms naphtha, kerosene, gas oil and bottoms are used herein as having their conventional meaning in the field of petroleum refining processes; see Alfke et al. (2007) Oil Refining, Ullmann's Encyclopedia of Industrial Chemistry, and Speight (2005) Petroleum Refinery Processes, Kirk-Othmer Encyclopedia of Chemical Technology. In this regard, it should be noted that there may be overlaps between different crude oil distillation fractions due to the complex mixture of hydrocarbon compounds contained in crude oil and the technical limitations of the crude oil distillation process. Preferably, the term naphtha as used herein refers to an oil fraction obtained from the distillation of crude oil having a boiling point range of about 20-200°C, more preferably about 30-190°C. Preferably, the light naphtha is a fraction having a boiling point range of about 20-100°C, more preferably about 30-90°C. The heavy naphtha preferably has a boiling point range of about 80-200°C, more preferably about 90-190°C. Preferably, the term kerosene as used herein refers to a petroleum fraction obtained from the distillation of crude oil having a boiling point range of about 180-270°C, more preferably about 190-260°C. Preferably, the term gas oil as used herein refers to a petroleum fraction obtained from the distillation of crude oil having a boiling point range of about 250-360°C, more preferably about 260350°C. Preferably, the term bottoms as used herein refers to an oil fraction obtained from the distillation of crude oil having a boiling point greater than about 340°C, more preferably greater than about 350°C.

As used herein, the term refinery refers to a section of a petrochemical plant complex for the chemical conversion of crude oil into petrochemicals and fuels. In this regard, it should be noted that an olefin synthesis plant such as a steam cracker is also considered to be a refinery. In this specification, the various hydrocarbon streams generated by refineries or produced during refinery operation are referred to as refinery generated gases, refinery light distillate, refinery middle distillate, refinery heavy distillate. Accordingly, the distillate formed by the refinery is obtained by chemical conversion followed by separation, for example by distillation or by extraction, which is different from the crude oil fraction. The term refinery gases refers to the fraction of refinery products that is gaseous at ambient temperature. Accordingly, the gas stream generated by the refinery may contain gaseous compounds such as LPG and methane. Other components contained in the refinery gas stream may be hydrogen and hydrogen sulfide. As used herein, the terms light distillate, middle distillate, and heavy distillate have their conventional meanings in the field of petroleum refining processes; see Speight, J. G. (2005) loc.cit. In this regard, it should be noted that there may be overlaps between the various distillation fractions of crude oil due to the complex mixture of hydrocarbon compounds contained in the product stream formed during the operation of the refinery and the technical limitations of the distillation process used to separate the various fractions. . Preferably, the light distillate generated by the refinery is a hydrocarbon distillate produced during operation of the refinery having a boiling point range of about 20-200°C, more preferably about 30-190°C. The light distillate is often relatively rich in aromatic hydrocarbons having a single aromatic ring. Preferably, the middle distillate produced by the refinery is a hydrocarbon distillate produced during the operation of the refinery, having a boiling point range of about 180-360°C, more preferably about 190350°C. The middle distillate is often relatively rich in aromatic hydrocarbons having two aromatic rings. Preferably, the heavy distillate generated by the refinery is a hydrocarbon distillate produced during the operation of the refinery having a boiling point of greater than about 340°C, more preferably greater than about 350°C. The heavy distillate is often relatively rich in hydrocarbons having condensed aromatic rings.

The term aromatic hydrocarbons or aromatic compounds is very well known in the art. Accordingly, the term aromatic hydrocarbon refers to a cyclically conjugated hydrocarbon with a stability (due to delocalization) that is significantly greater than that of a hypothetical localized structure (eg, the Kekule structure). The most common way to determine the aromaticity of a given hydrocarbon is to observe diatropicity in the 1H NMR spectrum, for example the presence of chemical shifts ranging from 7.2-7.3 ppm for the benzene ring protons.

The terms naphthenic hydrocarbons or naphthenes or cycloalkanes as used herein have their established meaning and respectively refer to types of alkanes having one or more rings of carbon atoms in the chemical structure of the molecules.

As used herein, the term olefin has the conventional meaning. Accordingly, an olefin refers to an unsaturated hydrocarbon compound containing at least one carbon-carbon double bond. Preferably, the term olefins refers to a mixture containing two or more of ethylene, propylene, butadiene, butylene-1, isobutylene, isoprene and cyclopentadiene.

One of the petrochemical products formed in the process of the present invention is a mixture of benzene, toluene and xylenes (BTX or BTX). The term LPG as used in this document is the common abbreviation for liquefied petroleum gas. LPG usually consists of a mixture of C2-C4 hydrocarbons, i.e. mixtures of C2, C3 and C4 hydrocarbons. Preferably, the product formed in the process of the present invention also contains valuable aromatic hydrocarbons such as ethylbenzene. Accordingly, the present invention preferably provides a process for preparing a mixture of benzene, toluene, xylenes and ethylbenzene (BTXE, BTEX). The resulting product may be a physical mixture of different aromatic hydrocarbons, or may be further separated directly, such as by distillation, to obtain various purified product streams. Such a purified product stream may include a benzene product stream, a toluene product stream, a xylene product stream, and/or an ethylbenzene product stream.

The term BTX as used herein refers to a mixture of benzene, toluene and xylenes.

As used herein, the term C# hydrocarbons, where # is a positive integer, is used to describe all hydrocarbons having # carbon atoms. In addition, the term C#+ hydrocarbons is used to describe all hydrocarbon molecules having # or more carbon atoms. Accordingly, the term C5+ hydrocarbons is used to describe a mixture of hydrocarbons having 5 or more carbon atoms. The term C5+ alkanes respectively refers to alkanes containing 5 or more carbon atoms.

The method of the present invention includes the distillation of crude oil, which includes the separation of different fractions of crude oil based on differences in boiling points. As used herein, the term crude oil distillation unit refers to a fractionation column that is used to separate crude oil into fractions by fractional distillation; see Alfke et al. (2007) loc.cit. Preferably, the crude oil is processed in an atmospheric distillation unit to separate the gas oil and lighter fractions from the higher boiling components (atmospheric bottoms or bottoms). In the present invention, it is not necessary to send the bottoms to a vacuum distillation unit for further fractionation of the bottoms, and it is possible to process the bottoms as a single fraction. However, in the case of a relatively heavy crude oil being fed, it may be advantageous to further fractionate the bottoms with a vacuum distillation unit to further fractionate the bottoms.

- 4 040018 vacuum gas oil and vacuum residue fraction. If vacuum distillation is used, the vacuum gas oil fraction and the vacuum residue fraction can be processed separately in downstream refineries. For example, the vacuum residue fraction can be specifically directed to solvent deasphalting prior to further processing. Preferably, the term vacuum gas oil as used herein refers to a petroleum fraction obtained from the distillation of crude oil having a boiling point range of about 340-560°C, more preferably about 350-550°C. Preferably, the term vacuum residues as used herein refers to an oil fraction obtained from the distillation of crude oil having a boiling point greater than about 540°C, more preferably greater than about 550°C.

The term aromatic ring opening unit refers to a refinery in which an aromatic ring opening process is carried out. Aromatic ring opening is a specific hydrocracking process particularly suited to the conversion of feedstocks that are relatively rich in aromatics having a boiling point in the kerosene, gas oil and optionally vacuum gas oil boiling range to produce LPG and, depending on the specifics of the process and / or technological conditions, light distillate (gasoline obtained from aromatic (ARO) compounds). Such an aromatic ring opening (ARO) process is described in, for example, US Pat. No. 3,256,176 and US Pat. No. 4,789,457. target products from the unconverted material, and may also include the possibility of recycling the unconverted material to one or both of the reactors. The reactors can operate at a temperature of 200-600°C, preferably 300-400°C, a pressure of 3-35 MPa, preferably 5-20 MPa, together with 5-20 wt.% hydrogen (relative to hydrocarbon feedstock), while this hydrogen can flow in parallel with the hydrocarbon feed or countercurrent to the direction of flow of the hydrocarbon feed, in the presence of a bifunctional catalyst active in both hydrogenation-dehydrogenation and ring rupture, wherein said aromatic ring saturation and ring cleavage can occur. The catalysts used in such processes contain one or more elements selected from the group consisting of Pd, Rh, Ru, Ir, Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W and V, in the form of a metal or metal sulfide supported on an acidic solid such as alumina, silica, silica-alumina and zeolites. In this regard, it should be noted that the term applied to as used herein includes any conventional method for preparing a catalyst in which one or more elements are combined with a catalyst support. By selecting one or a combination of catalyst compositions, operating temperature, operating space velocity and/or partial pressure of hydrogen, the process can be directed towards complete saturation and subsequent rupture of all rings, or towards keeping one aromatic ring unsaturated and subsequent rupture of all rings except one. In the latter case, the ARO process produces a light distillate (ARO gasoline) that is relatively rich in hydrocarbon compounds having a single aromatic and/or naphthenic ring. In the context of the present invention, it is preferable to use an aromatic ring opening process optimized to open all aromatic rings and thus produce LPG from a light distillate that is relatively rich in single aromatic ring hydrocarbons. However, also in a mode in which all aromatic rings are opened, the ARO process can still generate small amounts of distillates, which are preferably recycled to refineries capable of processing and upgrading said distillates into petrochemicals or intermediates, which can be further refined into petrochemical products. Another aromatic ring opening process (ARO process) is described in US 7513988. Accordingly, the ARO process may include saturation of the aromatic ring at a temperature of 100500°C, preferably 200-500°C, more preferably 300-500°C, a pressure of 2- 10 MPa, together with 5-30 wt.%, preferably 10-30 wt.% hydrogen (relative to hydrocarbon feedstock) in the presence of a catalyst for the hydrogenation of aromatic compounds and ring rupture at a temperature of 200-600°C, preferably 300-400°C, pressure 1-12 MPa, together with 5-20 wt.% hydrogen (relative to hydrocarbon feedstock) in the presence of a ring breaking catalyst, while said saturation of the aromatic ring and ring breaking can be carried out in one reactor or in two consecutive reactors. The aromatic hydrogenation catalyst may be a conventional hydrogenation/hydrotreatment catalyst such as a catalyst containing a mixture of Ni, W and Mo on a refractory support, typically alumina. The ring breaking catalyst contains a transition metal or metal sulfide component and a support. Preferably, the catalyst contains one or more elements selected from the group consisting of Pd, Rh, Ru, Ir, Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W and V, in the form metal or metal sulfide deposited on an acidic solid such as alumina, silica, silica-alumina and zeolites. By selecting one or a combination of catalyst compositions, operating temperature, operating space velocity and/or

- 5 040018 partial pressure of hydrogen, the process can be directed towards complete saturation and subsequent rupture of all rings, or to the preservation of one aromatic ring unsaturated and subsequent rupture of all rings except one. In the latter case, the ARO process produces a light distillate (ARO gasoline) that is relatively rich in hydrocarbon compounds having a single aromatic ring. In the context of the present invention, it is preferable to use an aromatic ring opening process optimized to open all aromatic rings and thus produce LPG from a light distillate that is relatively rich in single aromatic ring hydrocarbons. However, also in a mode in which all aromatic rings are opened, the ARO process can still generate small amounts of distillates, which are preferably recycled to refineries capable of processing and upgrading said distillates into petrochemicals or intermediates, which can be further refined into petrochemical products. Other examples of aromatic ring opening processes to produce LPG are described in US 7,067,448 and US 2005/0101814.

By directing the aromatic ring opening process towards full saturation and then breaking all rings, or alternatively keeping one aromatic ring unsaturated and then breaking all but one ring, the ratio of olefins formed to aromatics formed in the process of the present invention can be adjusted in this way that it is possible to obtain a neutral hydrogen balance, depending on the hydrogen balance of the raw material. In the case of a hydrogen-rich feedstock such as shale oil, production of aromatics is (almost) not required to obtain a hydrogen balanced process as a whole.

The aromatic ring-opening hydrocarbon feedstock used in the process of the present invention preferably contains kerosene and gas oil obtained by distillation of crude oil in said process and a middle distillate produced by said process formed from a refinery. More preferably, the aromatic ring-opening hydrocarbon feedstock used in the process of the present invention contains naphtha, kerosene and gas oil obtained by distillation of crude oil in said process, and refinery-derived light distillate and refinery-derived middle distillate produced in said process.

The LPG produced in the process, which is directed to the synthesis of olefins, preferably contains LPG, which is part of the gas fraction formed by the distillation of crude oil, and LPG, which is part of the gases produced by the refinery.

Preferably, the method of the present invention includes the steps of:

(a) subjecting the crude oil to the distillation of the crude oil to form a gas fraction, bottoms and one or more fractions from kerosene and gas oil;

(b) subjecting the bottoms to bottoms upgrade to form LPG and a bottoms upgrade liquid effluent;

(c) subjecting the bottoms upgrade liquid effluent and one or more fractions selected from the group consisting of kerosene and gas oil to aromatic ring opening to produce LPG; and (d) subjecting the LPG formed in the combined process to olefin synthesis.

By specifically subjecting the bottoms to bottoms upgrade to form LPG and the bottoms upgrade liquid effluent, and by subjecting said bottoms upgrade liquid effluent to aromatic ring opening, the yield of propylene in the process of the present invention can be further increased. In addition, crude oil can be upgraded into petrochemicals, in particular propylene, to a much greater extent.

As used herein, the term bottoms upgrader refers to a refinery suitable for carrying out a bottoms upgrade process, which is a process of fracturing hydrocarbons contained in bottoms and/or heavy distillate generated by the refinery into lower boiling hydrocarbons; see Alike et al. (2007) loc.cit. Commercially available equipment includes a delayed coker, a fluidized bed coker, an FCC bottoms unit, a Flexicoker type coker, a visbreaking unit, or a catalytic hydrovisbreaking unit. Preferably, the bottoms upgrader may be a coker or a bottoms hydrocracker. The coker is a refinery that converts bottoms to LPG, light distillate, middle distillate, heavy distillate and petroleum coke. In this process, thermal cracking of hydrocarbon molecules with long chains in the composition of the bottoms of the bottoms into molecules with shorter chains occurs.

The feedstock for refining bottoms preferably contains bottoms and heavy distillate obtained in the process. Such heavy distillate may contain steam cracker heavy distillate, such as soot oil and/or cracked distillate, but may also contain bottoms upgrade heavy distillate, which may be recycled to full recovery. However, a relatively small bituminous pitch stream may be removed from the process.

Preferably, bottoms upgrading used in the process of the present invention is bottoms hydrocracking.

By selecting bottoms hydrocracking over other bottoms upgrades, the propylene yield and carbon efficiency of the process of the present invention can be further improved.

The bottoms hydrocracking plant is a refinery that is suitable for bottoms hydrocracking process, which is LPG bottoms conversion process, light distillate, middle distillate and heavy distillate. Bottoms hydrocracking processes are well known in the art; see, for example, Alfke et al. (2007) loc.cit. Accordingly, there are three main types of reactors used in commercial hydrocracking: a fixed bed (refrigerate bed) reactor, a fluidized bed reactor, and a slurry (flow) reactor. Bottoms hydrocracking processes in fixed bed reactors are capable of processing contaminated streams such as atmospheric bottoms and vacuum bottoms to form light and medium distillates that can be further processed to form olefins and aromatics. Catalysts used in fixed bed bottoms hydrocracking processes typically contain one or more elements selected from the group consisting of Co, Mo and Ni on a refractory support, typically alumina. In the case of heavily contaminated feedstocks, the catalyst in fixed bed bottoms hydrocracking processes can also be replenished to some extent (moving bed). Technological conditions typically include a temperature of 350-450°C and a pressure of 2-20 MPa g. Fluidized-bed bottoms hydrocracking processes are also well known and, among other things, are distinguished by the fact that the catalyst is continuously changed, allowing highly contaminated feedstock to be processed. Catalysts used in fluidized bed bottoms hydrocracking processes typically contain one or more elements selected from the group consisting of Co, Mo and Ni on a refractory support, typically alumina. The small particle size of the catalysts used effectively increases their activity (see similar compositions in forms suitable for use in a fixed bed). These two factors allow fluidized bed hydrocracking processes to achieve significantly higher yields of light products and higher levels of hydrogen addition compared to fixed bed hydrocracking units. Technological conditions typically include a temperature of 350-450°C and a pressure of 5-25 MPa g. Slurry bottoms hydrocracking processes are a combination of thermal cracking and catalytic hydrogenation to achieve high yields of distilled products from highly contaminated bottoms feed. In the first, liquid stage, the thermal cracking and hydrocracking reactions occur simultaneously in a fluidized bed under process conditions that typically include a temperature of 400-500° C. and a pressure of 15-25 MPa g. VAT residues, hydrogen and catalyst are introduced at the bottom of the reactor, and a fluidized bed is formed, the height of which depends on the feed rate (flow rate) and the required conversion. In these processes, the catalyst is continuously changed to achieve stable conversion levels in the operating cycle. The catalyst may be an unsupported metal sulfide that forms in situ within the reactor. In practice, the additional costs associated with fluidized bed and slurry phase reactors are only justified when high conversion of highly contaminated heavy hydrocarbon streams such as vacuum gas oils is required. Under these conditions, the limited conversion of very large molecules and the difficulties associated with catalyst deactivation make fixed bed processes relatively unattractive in the process of the present invention. Accordingly, fluidized bed and slurry reactors are preferred due to their increased light and medium distillate yield compared to fixed bed hydrocracking. As used herein, the bottoms upgrade liquid effluent refers to the product generated by the bottoms upgrade, excluding gaseous products such as methane and LPG, and the heavy distillate generated by the bottoms upgrade. The heavy distillate produced by the bottoms upgrade is preferably recycled to the bottoms upgrade until it is completely exhausted. However, it may be necessary to remove a relatively small bituminous pitch stream. From a carbon efficiency point of view, a bottoms hydrocracker is preferred over a coker because the latter produces significant quantities of petroleum coke that cannot be converted into high value petrochemicals. From a combined process hydrogen balance point of view, it may be preferable to choose a coking unit over a bottoms hydrocracking unit, since the latter consumes significant amounts of hydrogen. In addition, given the capital and/or operating costs, it may be advantageous to choose a coker unit over a bottoms hydrocracking unit. In case the bottoms are further fractionated using a vacuum distillation unit to separate the bottoms into a vacuum gas oil fraction and a vacuum residue fraction, it is preferable to subject the vacuum gas oil to vacuum gas oil hydrocracking, and the vacuum bottoms to vacuum bottoms hydrocracking, while the heavy distillate formed by hydrocracking vacuum residues, subsequently subjected to vacuum gas oil hydrocracking. In the event that the present invention involves vacuum distillation, the vacuum gas oil thus obtained is preferably fed to an aromatic ring opener along with one or more other hydrocarbon streams that are relatively rich in aromatics and that have boiling points in the range of kerosene and gas oil. . Such hydrocarbon streams which are relatively rich in aromatics and which have boiling points in the range of kerosene and gas oil may be selected from the group consisting of kerosene, gas oil and middle distillate. Vacuum hydrocracking residues are preferably subjected to slurry bottoms hydrocracking as described herein above.

Preferably at least 50 wt.%, more preferably at least 60 wt.%, even more preferably at least 70 wt.%, particularly preferably at least 80 wt.%, in particular preferably at least 90 wt. % and most preferably at least 95 wt.% of the combined kerosene and gas oil formed using the distillation of crude oil in the way is subjected to hydrocracking. Accordingly, preferably less than 50 wt.%, more preferably less than 40 wt.%, even more preferably less than 30 wt.%, particularly preferably less than 20 wt.%, particularly preferably less than 10 wt.% and most preferably less than 5 wt.% crude oil is converted into fuel in the process of the present invention.

As used herein, the term olefin synthesis plant refers to a plant in which an olefin synthesis process is carried out. This term includes any process for the conversion of hydrocarbons to olefins, including, without limitation, non-catalytic processes such as pyrolysis or steam cracking, catalytic processes such as propane dehydrogenation or butane dehydrogenation, and a combination of the two, such as catalytic steam cracking.

Very often, the olefin synthesis process involves pyrolysis or steam cracking. As used herein, the term steam cracking refers to a petrochemical process that breaks saturated hydrocarbons into smaller, often unsaturated hydrocarbons such as ethylene and propylene. In steam cracking, gaseous hydrocarbon feedstocks such as ethane, propane and butanes or mixtures thereof (gas cracking) or liquid hydrocarbon feedstocks such as naphtha or gas oil (liquid cracking) are diluted with steam and briefly heated in a furnace in the absence of oxygen. Typically, the reaction temperature is 750-900° C. and the reaction can only be carried out for a very short time, typically with a residence time of 50-1000 ms. Preferably, a relatively low process pressure is chosen, from atmospheric to 175 kPa g. Preferably, hydrocarbon compounds such as ethane, propane and butanes are cracked separately in their respective furnaces to ensure optimum cracking conditions. After reaching the cracking temperature, the gas is rapidly cooled to stop the reaction in a pipeline heat exchanger or inside a quench manifold using quench oil. Steam cracking results in the slow deposition of coke, a form of coal, on the walls of the reactor. Removing the coke requires disconnecting the furnace from the process, and then a stream of steam or steam/air mixture is passed through the furnace coils. This converts the solid layer of non-volatile carbon into carbon monoxide and carbon dioxide. After this reaction is complete, the furnace is returned to service. The products obtained by steam cracking depend on the composition of the feedstock, the ratio between the amounts of hydrocarbon and steam, and the cracking temperature and residence time in the furnace. Light hydrocarbon feedstocks such as ethane, propane, butane or light naphtha produce product streams rich in polymerization grade lighter olefins including ethylene, propylene and butadiene. Heavier hydrocarbons (crude and heavy naphtha and gas oil fractions) also produce products rich in aromatics.

To separate the various hydrocarbon compounds obtained as a result of steam cracking, the cracked gas is sent to a fractionation unit. Such fractionators are well known in the art and may include a so-called gasoline distillation column in which heavy distillate (soot) and middle distillate (cracked distillate) are separated from light distillate and gases. In the downstream optional quench tower, most of the light distillate produced by steam cracking (pyrolysis gasoline or pyrogasoline) can be separated from the gases by condensation of the light distillate. Subsequently, the gases may be subjected to several compression stages, with the remainder of the light distillate being separated from the gases between the compression stages. Also acid gases (CO ₂ and H ₂ S) can be removed between compression stages. At the next stage, the gases obtained using pyroli

- 8 040018 for, can be partially condensed in the stages of the cascade cooling system approximately to the state in which only hydrogen remains in the gas phase. The various hydrocarbon compounds can subsequently be separated by simple distillation, with ethylene, propylene and C4 olefins being the most important high value chemicals produced by steam cracking. The methane produced from steam cracking is typically used as a fuel gas, and the hydrogen can be separated and recycled to processes that consume hydrogen, such as hydrocracking processes. Acetylene obtained from steam cracking is preferably selectively hydrogenated to ethylene. The alkanes contained in the cracked gas can be recycled to the olefin synthesis process.

Preferably, the synthesis of olefins used in the present invention includes the pyrolysis of ethane and the dehydrogenation of propane. By converting one or more fractions selected from naphtha, kerosene and gas oil obtained by distillation of crude oil in the process, and formed by the refinery light distillate and / or formed by the refinery middle distillate obtained in the process, in the LPG, propane, which is part of LPG can be subjected to propane dehydrogenation to form propylene and hydrogen, resulting in an increased propylene yield compared to pyrolysis, since the propane dehydrogenation process essentially produces no undesirable by-products.

By choosing olefin synthesis involving propane dehydrogenation, the overall hydrogen balance of the combined process can be improved. Another advantage of integrating the dehydrogenation process into a combined process is that a high purity hydrogen stream is produced that can be used as a feedstock for aromatic ring opening or other upstream process units without costly purification.

The term propane dehydrogenation plant as used herein refers to a petrochemical process plant in which a propane feed stream is converted to a product containing propylene and hydrogen. Accordingly, the term butane dehydrogenation plant refers to a process plant for the conversion of butane feedstock to C4 olefins. Together, the dehydrogenation processes of lower alkanes such as propane and butanes are described as the lower alkane dehydrogenation process. Processes for the dehydrogenation of lower alkanes are well known in the art and include oxidative dehydrogenation processes and non-oxidative dehydrogenation processes. In the oxidative dehydrogenation process, process heat is provided by the partial oxidation of the lower alkane(s) in the feedstock. In the non-oxidative dehydrogenation process, which is preferred in the context of the present invention, process heat for the endothermic dehydrogenation reaction is provided by external heat sources such as hot flue gases from fuel gas combustion or steam. In the process of non-oxidative dehydrogenation, the process conditions typically include a temperature of 540700°C and an absolute pressure of 25-500 kPa. For example, the Oleflex (UOP) process allows the dehydrogenation of propane to form propylene and the dehydrogenation of (iso)butane to form (iso)butylene (or mixtures thereof) in the presence of a catalyst containing platinum supported on alumina in a moving bed reactor; see, for example, US 4,827,072. The STAR (Uhde) process allows for the dehydrogenation of propane to form propylene or the dehydrogenation of butane to form butylene in the presence of a promoted platinum catalyst supported on a zinc aluminate spinel substrate; see, for example, US 4926005. The STAR process has recently been improved by applying the principle of oxydehydrogenation. In the secondary adiabatic zone in the reactor, part of the hydrogen from the intermediate product is selectively converted with the addition of oxygen to form water. This shifts the thermodynamic equilibrium towards higher conversion and provides a higher yield. In addition, the external heat required for the endothermic dehydrogenation reaction is partly provided by the exothermic hydrogen conversion. The Catofin process (Lummus) uses multiple fixed bed reactors operating on a cyclic basis. The catalyst is activated alumina impregnated with 18-20 wt.% chromium; see, for example, EP 0192059 A1 and GB 2162082 A. The Catofin process has the advantage of being able to handle impurities that would poison the platinum catalyst. The products obtained from the butane dehydrogenation process depend on the nature of the butane feedstock and the butane dehydrogenation process used. In addition, the Catofin process allows the dehydrogenation of butane to form butylene; see, for example, US 7622623.

Preferably, the gas fraction formed by the crude oil distillation plant and the gases produced by the refinery are subjected to gas separation to separate different components, for example, to separate methane from LPG.

As used herein, the term gas separation plant refers to a refinery that separates various compounds contained in gases generated from a crude oil distillation unit and/or in gases derived from a refinery. Compounds that can be separated into separate streams in a gas separation plant include ethane, propane, butanes, hydrogen, and fuel gas primarily containing methane.

- 9 040018

Any conventional method suitable for the separation of these gases in the context of the present invention may be used. Accordingly, the gases may be subjected to multiple compression stages, whereby acid gases such as CO2 and H2S may be removed between the compression stages. In the next stage, the produced gases can be partially condensed in the stages of a cascade cooling system to approximately the state in which only hydrogen remains in the gas phase. The various hydrocarbon compounds may subsequently be separated by distillation.

Preferably, the synthesis of olefins used in the present invention includes the pyrolysis of butanes. By choosing the pyrolysis of butanes over, for example, the dehydrogenation of butanes, the propylene yield of the process of the present invention can be further improved.

Preferably, the process of the present invention further comprises subjecting a portion of the liquid bottoms upgrade effluent and optionally the light distillate formed by aromatic ring opening to fluid catalytic cracking to form light distillate, middle distillate and olefins, and subjecting said middle distillate , formed by catalytic cracking with a fluid catalyst, opening the aromatic ring.

By subjecting the bottoms upgrade liquid effluent to fluid catalytic cracking, the hydrogen consumption of the process of the present invention can be reduced compared to a process in which the heavy distillate generated by the bottoms upgrade is recycled to said bottoms upgrade to full production. In addition, by selecting a process including fluid catalytic cracking, the light distillate formed by aromatic ring opening can be more effectively upgraded into petrochemicals.

As used herein, the term fluid catalytic cracking unit or FCC unit refers to a refinery for converting high boiling high molecular weight hydrocarbon fractions of crude oil into low boiling hydrocarbon fractions and olefin gases. In an FCC unit, cracking usually occurs with a very active zeolite-based catalyst in a vertical or vertically inclined tube with a short contact time called a riser. The preheated feed is sprayed into the base of the riser through the feed nozzles where it contacts the very hot fluidized catalyst. Preferred process conditions used for fluid catalytic cracking typically include a temperature of 425-700° C. and a pressure of 10-800 kPa g. The hot catalyst vaporizes the feedstock and catalyzes a cracking reaction that breaks down the high molecular weight hydrocarbons into lighter components including LPG, light distillate and middle distillate. The catalyst/hydrocarbon mixture travels up the riser for a few seconds, after which the mixture is separated by cyclones. Catalyzed hydrocarbons are sent to the main distillation column (a component of the FCC unit for separation into fuel gas, LPG, light distillate, middle distillate and heavy distillate). The spent catalyst is separated from cracked hydrocarbon vapors and sent to a desorber where it is contacted with steam to remove hydrocarbons remaining in the catalyst pores. The spent catalyst then enters a fluidized bed regenerator where air (or in some cases air plus oxygen) is used to burn off the coke to restore catalyst activity and also provide the necessary heat for the next reaction cycle, since cracking is an endothermic reaction. The regenerated catalyst then flows to the bottom of the riser, repeating the cycle. The method of the present invention may include several FCC units operating under different process conditions, depending on the hydrocarbon feed and the composition of the target product. As used herein, the term Low Severity FCC or Refining FCC refers to an FCC process that is optimized to produce a light distillate that is relatively rich in aromatics (FCC gasoline). Because most conventional refineries are optimized for gasoline production, typical FCC process operating conditions can be considered to represent a low severity FCC. Preferred process conditions used for FCC refinery typically include a temperature of 425-570°C and a pressure of 10-800 kPag. As used herein, the term High Severity FCC or Petrochemical FCC refers to an FCC process that is optimized for the production of olefins. High stringency FCC processes are known in the art and are described in EP 0909804 A2, EP 0909582 A1 and US 5846402, among others. 10-800 kPag In the present invention, FCC with high rigidity is preferred.

The process of the present invention may require the removal of sulfur from certain crude oil fractions to prevent catalyst deactivation in downstream refinery processes such as catalytic reforming or fluid catalytic cracking. This hydrodesulfurization process is carried out in a hydrodesulfurization unit (FIDS) or

- 10 040018 hydrotreating unit; see Alike (2007) loc. cit. Typically, the hydrodesulfurization reaction proceeds in a fixed bed reactor at elevated temperatures of 200-425°C, preferably 300-400°C, and elevated pressures of 1-20 MPa g, preferably 1-13 MPa g, in the presence of a catalyst containing elements, selected from the group consisting of Ni, Mo, Co, W and Pt, with or without promoters, supported on alumina, the catalyst being in sulfide form.

In addition, the yield of propylene can be further improved by converting the butenes and ethylene produced in the process of the present invention to propylene using a metathesis process. The process of metathesis is well known in the art and, inter alia, is described in US 4575575.

In yet another aspect, the invention also relates to a process plant suitable for carrying out the method of the invention. This process plant and the method carried out in said process plant are shown in FIG. 1-3.

Accordingly, the present invention also provides a process unit for converting crude oil to petrochemicals, comprising a crude oil distillation unit (10), comprising an inlet for crude oil (100) and at least one outlet for kerosene and/or gas oil fractions (310);

an aromatic ring opening unit (26) comprising an inlet for hydrocarbon feedstock (302) supplied to open the aromatic ring and an outlet for LPG (222); and an olefin synthesis unit (30) comprising an inlet for LPG (200) formed by a combined petrochemical process unit and an outlet for olefins (500) in which said hydrocarbon feedstock supplied to open the aromatic ring contains:

kerosene and/or gas oil formed by the crude oil distillation unit (10); and a refinery-derived middle distillate from said combined petrochemical process plant.

This aspect of the present invention is shown in FIG. 1.

The crude oil distillation unit (10) preferably further comprises a gas fraction outlet (230). LPG (222) formed by opening an aromatic ring and LPG (220) included in the gas fraction and LPG from a refinery formed in a combined process can be combined to form an input for LPG (200) formed by a combined petrochemical technological installation. The crude oil distillation unit (10) preferably further comprises an outlet for a naphtha fraction. The naphtha formed by the crude oil distillation unit may be combined with kerosene and/or gas oil formed by the crude oil distillation unit and subsequently subjected to aromatic ring opening. In addition, fractions (310) of kerosene and/or gas oil and optionally naphtha from a crude distillation unit may be combined with fractions (320) of refinery-derived middle distillate and optionally refinery-derived light distillate from the combined petrochemical technological installation, with the formation of input for hydrocarbon feedstock (302) supplied to the opening of the aromatic ring.

As used herein, the term inlet for X or outlet for X, where X is a given hydrocarbon fraction or the like, refers to an inlet or outlet for a stream containing said hydrocarbon fraction or the like. In the event that the outlet for X is directly connected to the downstream refinery containing the inlet for X, said direct connection may also contain additional installations, such as heat exchangers, separation and / or purification plants to remove unwanted compounds contained in the specified stream, etc.

Where more than one feed stream is supplied to a refinery in the context of the present invention, said feed streams may be combined to form a single refinery inlet or may form separate refinery inlets.

In addition, the installation (10) of the distillation of crude oil may contain a conclusion for VAT residues (400). The process plant of the present invention may further comprise a bottoms upgrade unit (40), comprising an inlet for bottoms (400) and heavy distillate generated by the refinery (401), and an outlet for LPG (223) formed by bottoms upgrading, and an outlet for the liquid effluent (326) to upgrade bottoms. This aspect of the present invention is shown in FIG. 2. The bottoms upgrade liquid effluent (326) contains a light distillate and a middle distillate, which preferably undergo aromatic ring opening.

The process unit of the present invention may further comprise a fluidized catalytic cracking unit (60) comprising an inlet for a portion of the bottoms upgrade liquid effluent (326) and optionally a light distillate (328) formed by aromatic ring opening, and an outlet for fractions (324) of middle distillate and optionally light distillate formed by catalytic cracking with

- 11 040018 fluid catalyst, and for olefins (540) formed by fluid catalytic cracking. This aspect of the present invention is shown in FIG. 2.

The fluid catalytic cracker (60) may also contain a methane outlet (720). In addition, the process unit containing the fluid catalytic cracking unit (60) may contain an outlet for a portion (327) of the bottoms upgrade liquid effluent stream that has not been subjected to fluid catalytic cracking, which can be recycled to unit (40) refinement of vat residues. In the event that a portion of the light distillate formed by opening the aromatic ring is subjected to fluidized catalytic cracking, the process unit comprises an aromatic ring opening unit (26) having a light distillate outlet (328) that is combined with a portion of the liquid effluent stream (326 ) refinement of bottom residues. The fluidized catalytic cracking unit (60) may further comprise a heavy distillate outlet (410) from the fluidized catalytic cracking, which can be recycled to the bottoms upgrade unit (40) to further upgrade said heavy distillate; see also FIG. 3.

Preferably, the process plant of the present invention further comprises: a gas separation plant (50) comprising an inlet for gases (200) formed in the combined process, an outlet for ethane (240) and a outlet for propane (250);

an ethane cracker (31) comprising an ethane inlet (240) and an ethylene outlet (510); and a propane dehydrogenation unit (32) comprising a propane inlet (250) and a propylene outlet (520). This aspect of the present invention is shown in FIG. 3.

The gas separation unit (50) may additionally comprise an outlet for methane (701). The ethane cracker (31) may further comprise an ethane cracking hydrogen outlet (810) and an ethane cracking methane outlet (710). The propane dehydrogenation unit (32) may further comprise a hydrogen outlet (820) produced by propane dehydrogenation.

The gas separation unit (50) may additionally contain an outlet for butanes (260), while the specified process unit additionally contains a butane cracking unit (34) containing an inlet for butanes (260) and an outlet for olefins (530), and an outlet for BTX ( 600). BTX is preferably included in the pyro-gasoline formed by said butane cracker (34). This aspect of the present invention is shown in FIG. 3. The butane cracking unit (34) may further comprise a hydrogen outlet (830) obtained by butane cracking and a methane outlet (730) obtained by butane cracking.

In this regard, it should be noted that the cracked gas produced in different cracking units can be combined and sent to a single separation zone to separate the olefins and BTX contained in the cracked gas from other components.

The present invention also provides the use of the process plant of the present invention for the conversion of crude oil into petrochemicals containing olefins and BTX.

Another preferred feature of the present invention is that all undesired products, such as petrochemicals that are not of high value, can be recycled to an appropriate plant for the conversion of such an undesirable product or to a target product (for example, a high value petrochemical) or to a product which is suitable as raw material for another installation. This aspect of the present invention is shown in FIG. 3.

In the method and process plant of the present invention, all of the methane produced is collected and preferably subjected to a separation process to produce fuel gas. Said fuel gas is preferably used to provide process heat in the form of hot flue gases resulting from the combustion of the fuel gas or via steam generation. Alternatively, methane may be steam reformed to produce hydrogen. In addition, undesirable by-products formed, for example, during steam cracking, can be recycled. For example, soot oil and cracked distillate from steam cracking can be recycled to an aromatic ring opener.

The various units operating in the process or process unit of the present invention are further combined by feeding hydrogen produced in certain processes, such as olefin synthesis, as a feed stream to processes in which hydrogen is needed as a feedstock, such as hydrocracking. In the event that the process and plant are net consumers of hydrogen (i.e. during start-up of the process or plant, or because all hydrogen-consuming processes consume more hydrogen than all hydrogen-producing processes produce), additional methane or fuel gas may need to be reformed , in addition to the fuel gas generated by the method or process plant of the present invention.

The following reference numbers are used in FIG. 1-3:

- 12 040018

- installation of distillation of crude oil;

- installation of aromatic ring opening;

- plant for the synthesis of olefins;

- ethane cracking unit;

- installation of propane dehydrogenation;

- butane cracking unit;

- unit for upgrading bottoms, preferably a hydrocracking unit for bottoms;

- installation of separation of gases;

- installation of catalytic cracking with a fluidized catalyst;

100 - crude oil;

200 - LPG formed in a combined way;

220 - gas fraction and LPG formed in an oil refinery, obtained in a combined method;

222 - LPG obtained by opening the aromatic ring;

223 - LPG obtained by upgrading bottoms;

230 - gas fraction;

240 - ethane;

250 - propane;

260 - butanes;

302 - hydrocarbon raw material supplied to the opening of the aromatic ring;

310 - kerosene and/or gas oil and, optionally, naphtha, formed by the installation of the distillation of crude oil;

320 shows refinery-derived middle distillate and optionally refinery-derived light distillate from a combined petrochemical process unit;

324 - middle distillate and optionally light distillate formed by fluid catalytic cracking;

326 - liquid effluent for upgrading bottoms;

327 is the portion of the bottoms upgrade liquid effluent that is not subjected to fluid catalytic cracking;

328 - light distillate formed by opening the aromatic ring;

400 - vat residues;

401 - heavy distillate formed by the refinery;

410 is a heavy distillate formed by fluid catalytic cracking;

500 - olefins;

510 - ethylene formed by ethane cracking;

520 - propylene formed by the propane dehydrogenation unit;

530 - C4 olefins formed by butane cracking;

540 - olefins formed by fluid catalytic cracking;

600 - BTK;

701 - methane obtained by gas separation;

710 - methane obtained by cracking ethane;

720 - methane obtained using catalytic cracking with a fluid catalyst;

730 - methane obtained by cracking butanes;

810 - hydrogen obtained by cracking ethane;

820 - hydrogen obtained by propane dehydrogenation;

830 - hydrogen obtained by cracking butanes.

While the invention has been described in detail for purposes of illustration, it should be understood that such details are provided solely for that purpose and that changes may be made by those skilled in the art without departing from the spirit and scope of the invention as defined by the claims.

It should also be noted that the invention relates to all possible combinations of features described herein, in particular, those combinations of features that are present in the claims are preferred.

It should be noted that the term containing does not exclude the presence of other elements. However, it should also be understood that a description of a product containing certain components also describes a product consisting of those components. Likewise, it should also be understood that a description of a process containing certain steps also describes a process consisting of those steps.

The present invention will now be described in more detail using the following non-limiting examples.

- 13 040018

Comparative example 1.

The experimental data presented here were obtained using process flow simulations in Aspen Plus. The kinetics of the steam cracker was strictly taken into account (using software for calculating the composition of the products of the steam cracker). The following steam cracker furnace conditions were used: ethane and propane furnaces: COT (coil outlet temperature) = 845°C and steam to oil ratio = 0.37, C4 furnaces and liquid feed furnaces: coil outlet temperature = 820°C and the ratio of water vapor to oil = 0.37.

To open the aromatic ring, a reaction scheme was used in which all aromatic, naphthenic, and paraffinic compounds were converted into LPG. The composition of propane dehydrogenation products was obtained based on the literature data. The bottoms hydrocracking unit and the FCC unit were modeled based on data presented in open sources.

In Comparative Example 1, Arabian light crude oil was distilled in an atmospheric distillation unit. All fractions, except VAT residues, were subjected to steam cracking. Fractions sent to the steam cracker included LPG, naphtha, kerosene and gas oil. The separation limit for VAT residues was 350°C. The total share of crude oil sent to the steam cracker was up to 52% by weight of crude oil. In the steam cracker, the above crude oil fractions were cracked in furnaces. The results are shown in the table as shown below.

The products that were derived from crude oil were divided into petrochemicals (olefins and BTEX, where BTEK is an abbreviation meaning BTX + ethylbenzene) and other products (hydrogen, methane and heavy ends containing C9 tars, cracked distillate, carbon black and vat residues). The total amounted to 100% of the total crude oil as bottoms were also taken into account. From the composition of the crude oil products, the carbon efficiency was determined as:

(Total Carbon Mass Content of Petrochemicals) / (Total Carbon Mass Content of Crude Oil).

For a comparative example, the yield of propylene was 8 wt.% of the total amount of crude oil.

Example 1

Example 1 is identical to the comparative example, except for the following:

fractions of naphtha, kerosene and gas oil (separation limit 350°C) from the distillation of crude oil were subjected to aromatic ring opening, which was carried out under process conditions to open all aromatic rings and convert the remaining paraffins and naphthenes in the LPG (intermediate product). This LPG was separated into ethane, propane and butane fractions. The ethane and butane fractions were subjected to steam cracking. The propane fraction was dehydrogenated to propylene (with an ultimate selectivity of 90%).

The table below shows the total composition of the products from the steam cracker (cracked light ends, ethane and butane) and from the propane dehydrogenation plant, in wt.% of the total amount of crude oil. The table also contains the atmospheric distillation residue fraction.

In example 1, the propylene yield was 29% by weight of the total crude oil.

Example 2

Example 2 is identical to example 1 except for the following:

first, bottoms were upgraded in a bottoms hydrocracking unit to produce gases, light distillate and middle distillate. The gases obtained by hydrocracking bottoms were subjected to steam cracking. The light distillate and middle distillate obtained by bottoms hydrocracking were subjected to aromatic ring opening, which was carried out under process conditions to open all aromatic rings and convert the remaining paraffins and naphthenes to LPG (intermediate product). This LPG was separated into ethane, propane and butane fractions. The ethane and butane fractions were subjected to steam cracking. The propane fraction was dehydrogenated to propylene (with an ultimate selectivity of 90%).

In addition, the heavy portion of the effluent from the cracker (C9 resins, cracked distillate and soot) was recycled to the bottoms hydrocracker. The maximum conversion in the bottoms hydrocracking unit is close to full (bituminous pitch from the bottoms hydrocracking unit was 2 wt.% of crude oil).

The table below shows the total composition of crude oil products from the steam cracker (cracked light ends, naphtha and LPG) and from the propane dehydrogenation plant, in wt.% of the total amount of crude oil. The composition of the products also contains bituminous pitch from the hydrocracking unit (2 wt.% of crude oil).

In example 2, the propylene yield was 55% by weight of the total crude oil.

Example 3

Example 3 is identical to example 2 except for the following:

The hydrocracking unit did not crack to full capacity, but the conversion to gas, light distillate and middle distillate fractions was 37.5%. The rest of the product from the installation

- 14 040018 hydrocracking bottoms (heavy distillate and bottoms) were sent to the FCC unit for the formation of light fractions and FCC naphtha. The light ends were sent to a steam cracker where the olefins in the light ends were separated from the LPG. This LPG was separated into ethane, propane and butane fractions. The ethane and butane fractions were subjected to steam cracking. The propane fraction was dehydrogenated to propylene (with an ultimate selectivity of 90%). FCC naphtha was subjected to aromatic ring opening under process conditions to open all aromatic rings and convert the remaining paraffins and naphthenes to LPG (intermediate). This LPG was separated into ethane, propane and butane fractions. The ethane and butane fractions were subjected to steam cracking. The propane fraction was dehydrogenated to propylene (with an ultimate selectivity of 90%). The light cycle oil (LCO) from the FCC unit was recycled to the bottoms hydrocracking unit.

The table below shows the total composition of crude oil products from the steam cracker (cracked light ends, naphtha and LPG) and from the propane dehydrogenation plant, in wt.% of the total amount of crude oil. The product formulation also contains bituminous pitch from a hydrocracking unit and coke from an FCC unit (together 4 wt% of crude oil).

In Example 3, the propylene yield was 51% by weight of the total crude oil.

Comparative Example Example 1 Example 2 Example 3 Petrochemicals (wt.% crude oil) Ethylene 15% 15% 29% 27% Propylene 8% 29% 55% 51% Butadiene 2% 0% 1% 1% 1-butene 1% 1% 2% 4% Isobutene 1% 0% 1% 2% Isoprene 0% 0% 0% 0% Cyclopentadiene 1% 0% 0% 0% Benzene 4% 0% 1% 1% Toluene 2% 0% 0% 0% xylene 1% 0% 0% 0% Ethylbenzene 1% 0% 0% 0% Other components (wt.% crude oil) hydrogen 1% 2% 4% 4% methane 7% 3% 6% 6% Heavy components 56% 48% 0% 0% Bituminous pitch RHC and coke FCC 0% 0% 2% 4% Carbon efficiency 38.0% 47.6% 92.5 88.6

CLAIM

Claims (1)

A combined process for the conversion of crude oil to petrochemicals, in which crude oil distillation, aromatic ring opening and olefin synthesis are carried out, and this method includes the stages in which the hydrocarbon feed is subjected to aromatic ring opening to form liquefied petroleum gas (LPG) and subjected to LPG formed in a combined process comprising at least LPG obtained by opening an aromatic ring to olefin synthesis, wherein said hydrocarbon feedstock contains: kerosene and/or gas oil obtained by distillation of crude oil in the specified method; and formed by the refinery middle distillate obtained in the specified method, and in the specified method, the following steps are carried out: subjecting the crude oil to the distillation of the crude oil to form a gas fraction, bottoms and one or more fractions from kerosene and gas oil; subjecting the bottoms to bottoms upgrading, which is bottoms hydrocracking, to form LPG and bottoms upgrading liquid effluent containing middle distillate; subjecting a portion of the bottoms upgrade liquid effluent and one or more fractions selected from the group consisting of kerosene and gas oil to aromatic ring opening to produce LPG; and subjecting the LPG formed in the combined process, comprising at least said LPG formed by aromatic ring opening and said LPG formed by refining bottoms, to olefin synthesis, said olefin synthesis comprising ethane pyrolysis, propane dehydrogenation and butane dehydrogenation; subjecting another portion of the bottoms upgrade liquid effluent to fluid catalytic cracking to form light distillate, middle distillate and olefins; and subjecting said middle distillate formed by fluid catalytic cracking to aromatic ring opening; moreover, at least 50 wt.% from the combined fractions of kerosene and gas oil obtained by distillation of crude oil in the specified method, subjected to the opening of the aromatic ring; moreover, the opening of the aromatic ring is carried out in at least one reactor operating at a temperature of 200-600°C, a pressure of 3-35 MPa, together with 5-20 wt.% hydrogen relative to the hydrocarbon feed, while the specified hydrogen flows countercurrent to the direction of flow of the hydrocarbon raw materials, in the presence of a bifunctional catalyst active both in relation to hydrogenation-dehydrogenation and ring rupture, while said saturation of the aromatic ring and ring rupture is carried out, while the bifunctional catalyst used contains one or more elements selected from the group consisting of Pd , Rh, Ru, Ir, Os, Ga, In, Mo, W, and V, in the form of a metal or metal sulfide, supported on an acidic solid selected from the group consisting of silica and silica-alumina.