EP3331968B1

EP3331968B1 - System and method for hydroconversion of heavy oils by means of reactors with dispersed catalyst or with expanded catalyst bed with introduction of gas at the head of the reactor

Info

Publication number: EP3331968B1
Application number: EP15820648.2A
Authority: EP
Inventors: Luigi Patron
Original assignee: Patron Luigi
Current assignee: Patron Luigi
Priority date: 2015-08-06
Filing date: 2015-10-02
Publication date: 2019-04-10
Anticipated expiration: 2035-10-02
Also published as: EP3331968A1; ITUB20152927A1; WO2017021987A1

Description

Field of application of the invention

The present invention relates to heavy oil hydroconversion systems which employ reactors with dispersed catalyst or with expanded catalyst bed (in which also dispersed catalyst is possibly fed or in which only dispersed catalyst is fed).
More precisely, the present invention refers to a system of the aforesaid type in which the extraction of the reaction liquid through the biphasic effluent is enhanced by means of the introduction of gas at the head of the reactor, in a zone containing biphasic foam above the reaction liquid.
The invention also refers to a method for hydroconversion of heavy oils that can be actuated by means of the aforesaid system.

Review of the prior art

In heavy oils (such as crude oil, bitumen, petroleum from tar sands, shale oils and their residues of atmospheric distillation, of vacuum distillation and of thermal visbreaking), hydrocarbons are present in variable percentage that have a boiling point higher than 540 °C. Such hydrocarbons (containing metals such as nickel, vanadium and iron, and heteroatoms such as S, N and O) constitute a not-totally-distillable heavy fraction of said oils. If subjected to evaporation, said hydrocarbons in fact produce a quantity of carbon residue (expressed as % CCR, i.e. Conradson Carbon Residue - ASTM D189) that is greater the lower their hydrogen content. Hydrocarbons with a limited content of hydrogen, such as 8% by weight, when subjected to evaporation leave a carbon residue that can reach 50% of their weight. Such carbon residue is reduced to 20% for contents of hydrogen around 10% by weight and is mostly absent when the hydrogen content of the hydrocarbon is positioned around 12% by weight.
In order to convert the heavy oils into light products with greater market value, the heavy oils are subjected to a temperature treatment with hydrogen and suitable catalysts by means of which the aforesaid heavy fraction (also termed "carbonaceous fraction") is converted into distillable hydrocarbons. Said treatment is also known as "hydroconversion". The catalysts employed in such treatments are generally defined "hydrogenation" catalysts or "hydroconversion" catalysts. Often, only the fraction of the heavy oil having a boiling point higher than 540 °C is subjected to hydroconversion. The hydroconversion treatment is aimed to obtain products free of carbon residue, which can therefore be fed to subsequent treatments of hydrocracking and hydrotreating, by means of which said products attain quality specifications required by the market or they can be used for other refining processes. The hydrocracking and hydrotreating technologies are well-tested and available on the market, therefore further details thereon will not be provided herein.
The hydroconversion treatment can be carried out in pressurized cylindrical vessels (i.e. "reactors") with distribution of the hydrogen at the base, where also the heavy oil to be converted is introduced. The charge stock to be converted and the hydrogen come into contact in the presence of a hydrogenation catalyst (usually comprising molybdenum) dispersed in the reaction liquid or carried on a solid support, structured in small cylinders or microspheres, constituted by silica and/or alumina. The catalyst deposited on a solid support will also be indicated hereinbelow in the present description with the expression "supported catalyst".
If the hydrogenation catalyst is dispersed in the reaction liquid, the reactor in which the hydroconversion treatment is carried out is defined "with dispersed catalyst". If the hydrogenation catalyst is deposited on a solid support (i.e. if the catalyst is supported), the reactor in which the hydroconversion treatment is carried out is defined "with expanded catalyst bed". A third category of hydroconversion reactors is constituted by reactors with expanded catalyst bed in which dispersed catalyst is also fed. A fourth category of hydroconversion reactors is constituted by reactors with expanded catalyst bed in which only dispersed catalyst is fed.
In the reactors with dispersed catalyst, the catalyst can be introduced into the reactor in various modes, such as by means of an oil-soluble precursor (i.e. a metal compound capable of generating the active species when it is in contact with the charge stock and the hydrogen). The catalyst of dispersed type, also termed "slurry catalyst", remains uniformly and stably dispersed in the reaction liquid from which it can be separated, by way of example, via filtration, centrifugation or settling by decanter. In the patent US 4226742 , a process is described for heavy oil hydroconversion which employs catalysts of dispersed type. In such patent, it is shown that, using dispersed catalyst containing molybdenum, the maximum catalytic effect is attained at concentrations of metal in the reaction liquid around 1000 ppm by weight. As described in EP 2291238 , the dispersed catalyst can be combined with a silica-alumina based compound or a zeolite in order to facilitate the denitrification of the charge stock. At the base of the reactor with dispersed catalyst (also termed "bubble column", "slurry reactor" or "bubble column slurry reactor"), a nozzle grid is present for the uniform distribution of the hydrogen. The commercial application of the hydroconversion systems using reactors with dispersed catalyst has been up to now discouraged owing to the high catalyst consumptions. Being dispersed in the reaction liquid, such catalyst flows out with the latter from the reactor; this involves the need to continuously replenish the reaction liquid with fresh catalyst.
In the reactors with expanded catalyst bed, the solid elements on which the hydrogenation catalyst is deposited are maintained suspended in the reaction liquid by means of a circulation thereof, obtained by means of a pump inside or outside the reactor. Said pump is also defined "ebullating pump". For this reason, the reactors with expanded catalyst bed are also known as "ebullated catalytic bed reactors". Within the reactor, a funnel can be present that is provided with a "downcomer" pipe that collects the liquid in the upper part of the reactor and conveys it downward, being suctioned by the circulation pump. In the patent US 4684456 , it is described that, by suitably adjusting the rate of speed of the pump, the supported catalyst remains suspended in the reaction liquid and at the same time confined within the reactor, so as to prevent the catalyst from flowing out with the reaction liquid. By means of a perforated plate placed at the bottom of the reactor, the hydrogen and the heavy charge stock to be converted are distributed. The solid support (with silica and/or alumina base), as a function of its degree of acidity, can facilitate the removal of the nitrogen and the cracking of the charge stock. The hydroconversion systems which employ reactors with expanded catalyst bed (generally in multiple reaction stages), while being among the most used in industry, are still today characterized by a largely incomplete conversion that involves problems of market placement of the non-converted residue. The above-described considerations are also valid if, in the reactor with expanded catalyst bed, dispersed catalyst is also fed and if, in the reactor with expanded catalyst bed, only dispersed catalyst is fed.
In the reactors with dispersed catalyst or with expanded catalyst bed (in which also dispersed catalyst is possibly fed or in which only dispersed catalyst is fed), the hydrogen is generally introduced at the base of the reactor at a superficial velocity of several centimeters per second (measured as hydrogen volume at reaction temperature and pressure, fed in the unit of time and divided by the area of the reactor cross section). The introduction of the hydrogen generates a set of bubbles which, by ascending the reaction liquid, induces the remixing and ensures high heat and mass transfer coefficients both in axial direction and in radial direction of the reactor, even in the absence of agitator or mechanical mixing systems. At the mouth of the outlet duct (mouth positioned on the upper vault of the reactor or at its interior), the hydrogen, the volatile conversion products and the reaction liquid generate an effluent (defined "biphasic") which is sent to a gas-liquid separator, from whose head the gaseous phase exits from which - by means of condensation in one or more stages - the volatile conversion products, as well as the residual hydrogen which is sent to the purification section in order to then be reused, are recovered. The reaction liquid descends to the bottom of the separator; such reaction liquid is constituted by conversion products (mainly with high boiling point) dissolved in the non-converted fraction of charge stock, and reaction-generated solids comprising the sulfides of the metals present in the charge stock, coke, asphaltene resins that are insoluble and solids due to the catalyst.
The sum of the weight flow rates of the hydrocarbons extracted at the head of the separator and of the hydrocarbons, with a boiling point up to 540 °C, extractable from the reaction liquid of the separator bottom, can be taken as a measure of the reactor's hydroconversion capacity, whether the reactor is of the type with dispersed catalyst or is of the type with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed). In normal reaction conditions, the conversion products having a low boiling point (for example less than 300 °C) are mainly extracted through the gas phase of the biphasic effluent at the head of the reactor, without significant accumulation in the reaction liquid present inside the reactor. Differently, the conversion products having a higher boiling point (for example comprised between 300 °C and 540 °C) tend to be accumulated in the reaction liquid within the reactor and are extracted, both through the gaseous phase and through the liquid phase of the biphasic effluent, to an extent that increases with the superficial velocity of the hydrogen introduced at the base of the reactor. The conversion products having a boiling point comprised between 300 °C and 540 °C are also identified with the adjective "high-boiling".
In order to improve the extraction of high-boiling conversion products, one tends to maximize the quantity of hydrogen introduced at the reactor bottom. The superficial velocity of the hydrogen fed to the reactor bottom nevertheless cannot exceed specific values, beyond which the excessive volume occupied by the gas ("gas holdup"), taken away from the reaction liquid volume, prevents an efficient use of the reactor and induces the coalescence of the gas bubbles, a phenomenon to be avoided due to the negative repercussions both on the uniformity of the reaction means (the remixing effect is in fact decreased) and on the removal of the reaction heat (and hence on the control of the temperature). The extraction of the high-boiling conversion products in the conventional reactors, with dispersed catalyst or with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed), therefore cannot be increased beyond a specific limit by operating only on the flow rate of hydrogen introduced at the reactor bottom.
The accumulation, during reaction, of hydrocarbons with boiling point comprised between 300 °C and 540 °C, mainly of maltene nature and well-known to be characterized by low reactivity, indeed negatively affects the capacity of the reactor to convert fresh charge stock. The negative effect of the accumulation of the high-boiling conversion products on the hydroconversion capacity is particularly important in the case of commercial scale reactors where the accumulation advances with the height, progressively reducing the hydroconversion unit capacity of the system (defined as m³ of charge stock converted in one hour per m³ of reaction volume). A unit capacity that is thus reduced limits the convenience of making large-height reactors.
As described in the patent US 8709966 , the extraction of the high-boiling conversion products (and hence the capacity of hydroconversion of the reactor) improves by operating at high reaction temperatures (between 445 °C and 460 °C). Nevertheless, such conditions involve the production of insoluble organic solids (in quantities on average higher than 0.03 kg per kg of treated charge stock). This strongly limits the attainable degree of conversion, with problems of disposal of the non-converted fraction.
The accumulation of the high-boiling conversion products in the reaction liquid is a consequence of an extraction mode that is not suitable for the rate with which said conversion products are generated.
Only partial solutions have up to now been found for improving the extraction of the high-boiling conversion products in the conventional reactors with dispersed catalyst or with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed).
In the patent US 6454932 , relative to a system for hydroconversion of heavy oils with several reaction stages with use of reactors with expanded catalyst bed, in order to increase the recovery of the high-boiling conversion products, one proceeds with stripping with hydrogen the liquid phase present in the gas-liquid separator placed between one reactor and the next. No operation is proposed for improving the extraction of the high-boiling conversion products directly at the reactor, not even at the liquid state.
In the patent US 8236170 , relative to the hydroconversion of heavy oils by means of the use of dispersed catalyst, the extraction of the high-boiling conversion products is improved due to the combined effect of the temperature and of the high superficial velocity of the hydrogen introduced at the base (bottom) of the reactor. The increase of the superficial velocity of the gas, however, modifies the flow regime of the liquid phase of the reactor towards a condition that produces coalescence of the gas bubbles. In order to mitigate the effects produced by the increase of the superficial velocity of the gas introduced into the lower end of the reactor, a liquid recirculation reactor is employed in which a ratio is maintained between the liquid flow velocity and the superficial velocity of the gas of 0.2 or higher. The high liquid recirculation flow rate (necessary for accelerating the gas bubbles in order to reduce the residence time thereof in the reactor, so as to prevent coalescence) increases plant complexity and hence cost for an additional section of separation of the liquid and recovery of the gas.
In the Italian patent 1415850 , owned by the same Applicant, a hydroconversion system is described that improves the extraction of the high-boiling conversion products by using, alongside the conventional extraction via biphasic effluent, a second extraction mode which employs reaction liquid drawn directly from the bubbling zone of a hydroconversion reactor with dispersed catalyst. US 2015/0210940 discloses a system and process for hydroconversion of heavy oils.

Object of the invention

The object of the present invention is to overcome the aforesaid drawbacks and to indicate an alternative solution to that illustrated in the abovementioned Italian patent 1415850 in order to improve the extraction of the high-boiling conversion products through the biphasic effluent that flows out from the head of the reactor, with a mode applicable both to reactors with dispersed catalyst and to reactors with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed).
In other words, the objective of the present invention is to reduce the accumulation of high-boiling conversion products in the reactors with dispersed catalyst and in the reactors with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed) during a heavy oil hydroconversion process.

Summary of the invention

The present invention regards a system for hydroconversion of heavy oils which employs a reactor with dispersed catalyst or with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed), in which the conversion products are obtained from the biphasic effluent that flows out at the head of the reactor.
The hydrogen introduced at the bottom of the reactor generates a set of bubbles that ascend the reaction liquid. At the surface that delimits the reaction liquid, there is the separation of the bubbles and the consequent degassing of the liquid. Due to the foaming properties of the heavy oils (a consequence of the presence of heteroatoms with surfactant effect, such as S, N and O, mainly in the fraction with boiling point higher than 540 °C), the gas bubbles, being recombined, produce a flow of biphasic foam that by ascending lifts the reaction liquid towards the head of the reactor where the outlet duct is placed. Charge stocks lacking fractions with boiling point higher than 540 °C do not produce biphasic foam to an extent sufficient for lifting the liquid towards the outlet duct.
At the mouth of the outlet duct, as a consequence of the acceleration that the gas undergoes due to the narrowing of the flow section, there is the entrainment of the reaction liquid outside the reactor. Thus, a biphasic effluent is generated in which the gaseous phase (comprising hydrogen and conversion products at the gas state and at the vapor state) entrains the liquid phase (constituted by reaction liquid) dispersed therein. The entrainment is also verified with regard to the solids suspended in the reaction liquid.
The quantity (kg) of dispersed (and hence entrained) reaction liquid in one m³ of biphasic effluent, as a consequence of the low value of the density of the gas with respect to that of the liquid, is provided, in a first approximation, by the value of the density (kg/m³) of the biphasic effluent. The density of the biphasic effluent, in the outlet duct of the reactor, is several times lower than that of the biphasic foam which, within the reactor, ascends towards the outlet duct. More precisely, the density of the biphasic effluent is for example from 4 to 6 times lower than that of the biphasic foam. It can be inferred that a considerable part of the reaction liquid present in the biphasic foam falls into reaction and only a part thereof (which contains the conversion products at the liquid state) flows out of the reactor as liquid component of the biphasic effluent. The presence within the reactor, above the reaction liquid, of biphasic foam that contains significantly more reaction liquid than the biphasic effluent, offers the possibility to introduce gas at the biphasic foam in order to counteract the fall of the reaction liquid and force the outflow thereof in the biphasic effluent, without altering the flow regime of the underlying reaction liquid.
The extraction of the reaction liquid through the biphasic effluent, forced by the introduction of gas above the reaction liquid, is fully exploited in the present invention in order to ensure that the capacity of extraction of the conversion products present at the liquid state in the reactor can be increased without having to increase the reaction temperature or the flow rate of gas introduced at the base (bottom) of the reactor, with the drawbacks and limits that this would involve.
The extraction of the reaction liquid through the biphasic effluent, enhanced by the entrainment resulting from the introduction of gas at the head of the reactor, allows:

a) improving the hydroconversion unit capacity of the reactor and rendering it independent from the height thereof;
b) with reference to a reactor with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed), obtaining an extraction of the high-boiling conversion products suitable for operating the reactor in a single reaction stage, with high degree of conversion, so as to exceed the conversion limits that are normally encountered as a consequence of the precipitation of the asphaltenes in the final reaction stages;
c) with reference to a reactor with dispersed catalyst or to a reactor with expanded catalyst bed (in which only the silica-alumina support is present) in which dispersed catalyst is also fed or to a reactor with expanded catalyst bed in which only dispersed catalyst is fed and is present, reduce the formation of solids during reaction in order to operate in conditions of high accumulation of dispersed catalyst in the reaction liquid, minimizing the replenishment rate thereof i.e. the consumption with respect to the processed charge stock.

More in detail, object of the present invention, is a system for hydroconversion of heavy oils in a single reaction stage comprising:

a reactor, or several reactors in parallel, with dispersed catalyst or with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed);
a first line for feeding the reactor with heavy oil comprising at least 10% by weight of hydrocarbons having, at atmospheric pressure, a boiling point higher than 540 °C;
when the reactor is of the type with dispersed catalyst, a second line for feeding the reactor with a first gas including hydrogen,
when the reactor is of the type with expanded catalyst bed, said heavy oil is premixed with a first gas including hydrogen;
a third line for feeding the reactor with a hydroconversion catalyst,
the reactor containing:
- at least at a lower part of the reactor, reaction liquid traversed by said first gas (also defined "bubbling liquid"), and
- at least at an upper part of the reactor placed above the reaction liquid traversed by said first gas, a biphasic foam including at least one liquid phase and one gaseous phase, said biphasic foam originating a biphasic effluent when exiting from the head of the reactor;
a separator suitable for separating the liquid phase from the gaseous phase of said biphasic effluent;
a fourth line for drawing said biphasic effluent from the upper part of the reactor, and for introducing the same in the separator;
a first flash-distillation stage, preferably with atmospheric distillation;
a fifth line for drawing the liquid phase from the separator and for introducing the same in said first stage;
a second stage for concentrating via vacuum distillation;
a sixth line for drawing, from said first stage, a first residue thereof, and for introducing said first residue in said second stage;
a seventh line for drawing, from said second stage, a second residue thereof, and for introducing said second residue in the reactor,

an eighth line for introducing a second gas in the reactor at a zone thereof containing said biphasic foam (i.e. at said upper part of the reactor);
a densimeter suitable for measuring the density of said biphasic foam in said upper part of the reactor, in an intermediate position between the gas introduction at said eighth line and said drawing at said fourth line. The densimeter therefore measures the density of the biphasic foam following the introduction of said gas in the reactor through the eighth line and the introduction of said second residue through the seventh line.

Incidentally, the abovementioned stages of flash-atmospheric distillation and of concentration via vacuum distillation are known and therefore further details will not be provided thereon.
Further innovative characteristics of the present system invention are described in the dependent claims.
According to one aspect of the invention, the hydroconversion system comprises a nozzle grid by means of which said second gas can be introduced into the upper part of the reactor.
According to another aspect of the invention, the hydroconversion system comprises a ninth line for drawing, from said seventh line, solids generated during reaction.
According to another aspect of the invention, the hydroconversion system also comprises:

a third stage for centrifugation or decanting;
a tenth line for partial drawing of said first residue from said sixth line and introduction thereof in said third stage;
an eleventh line for drawing, from said third stage, solids generated under reaction;
a twelfth line for drawing, from said third stage, said first residue that at least partially lacks said solids, and introducing the same in said sixth line.

Incidentally, the abovementioned stage for centrifugation or decanting is known and therefore further details will not be provided thereon.
Another object of the invention is a method for hydroconversion of heavy oils in a single reaction stage (usable, by way of example, with the system that is the object of the invention) comprising the following steps:

a) feeding, at the bottom, at least one reactor with dispersed catalyst or with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed) with:
- heavy oil comprising at least 10% by weight of hydrocarbons having, at atmospheric pressure, a boiling point higher than 540 °C;
- a first gas including hydrogen;
- a hydroconversion catalyst,
generating:
- at least at a lower part of the reactor, reaction liquid traversed by said first gas (also defined "bubbling liquid"), and
- at least at an upper part of the reactor placed above the reaction liquid traversed by said first gas, a biphasic foam including at least one liquid phase and one gaseous phase, said biphasic foam originating a biphasic effluent when exiting from the head of the reactor;
b) drawing said biphasic effluent from the upper part of the reactor and separating the liquid phase from the gaseous phase thereof;
c) subjecting the liquid phase to flash-distillation, preferably atmospheric;
d) subjecting to concentration, via vacuum distillation, a first residue of the flash-atmospheric distillation;
e) introducing in the reactor a second residue of the concentration via vacuum distillation,

the reactor operating at a temperature comprised between 330 °C and 430 °C, and at a pressure comprised between 10 MPa and 30 MPa;
said second gas being introduced into the upper part of the reactor in step a), and said second residue being introduced into the reactor in step e), with respective flow rates such that the density of the biphasic foam, measured following the introduction of said second gas in the upper part of the reactor, is comprised between 100 kg/m³ and 500 kg/m³;
in step a), the second gas being introduced into the upper part of the reactor at a superficial velocity comprised between 0.1 cm/s and 50 cm/s.
By "surface speed of the second gas" it is intended the ratio between the flow rate with which, in step a), the second gas is introduced in the reactor, measured at the same temperature and at the same pressure of the latter, expressed in cm³ per second, and the area of the cross section of the reactor expressed in cm².
The superficial velocity of the second gas is therefore expressed in cm per second;
in step b), said biphasic effluent being originated with a superficial velocity of the gas equal to at least 10 times the liquid flow velocity.
By "superficial velocity of the gas" it is intended the sum of:
- the "superficial velocity of the first gas", i.e. the ratio between the flow rate with which, in step a), the first gas is introduced in the reactor, measured at the same temperature and at the same pressure of the latter, expressed in cm³ per second, and the area of the cross section of the reactor expressed in cm²
  and
- the superficial velocity of the second gas.
The superficial velocity of the gas is therefore expressed in cm per second. By "liquid flow velocity" it is intended the ratio between the flow rate with which, in step b), the liquid phase of the biphasic effluent is drawn from the upper part of the reactor, expressed in cm³ per second, and the area of the cross section of the reactor expressed in cm². The liquid flow velocity is therefore expressed in cm per second;
in step a), said second gas being introduced into the upper part of the reactor at a temperature such that the temperature of the biphasic foam is comprised between 330 °C and 430 °C;
in step e), said second residue being introduced in the reactor at a unit flow rate at least equal to 0.5 x Vs x H, where Vs is the space velocity with which, in step a), said heavy oil is introduced in the reactor and H is the height of the reactor.
By "unit flow rate of said second residue" it is intended the ratio between the flow rate with which said second residue is introduced in the reactor, expressed in m³ per hour, and the area of the cross section of the reactor expressed in m². The unit flow rate is therefore expressed in m per hour.
By "space velocity of said heavy oil" it is intended the ratio between the flow rate with which said heavy oil is introduced in the reactor, expressed in m³ per hour, and the volume of the reactor expressed in m³. The space velocity is therefore expressed in hours^-1.
The height H of the reactor is expressed in m.

Further innovative characteristics of the present method invention are described in the dependent claims.
According to another aspect of the invention, in step b), said biphasic effluent being originated with a superficial velocity of the gas equal to at least 15 times the liquid flow velocity.
According to another aspect of the invention, in step e), said second residue being introduced in the reactor at a unit flow rate at least equal to Vs x H.
According to another aspect of the invention, the superficial velocity of the gas is preferably higher than 2 cm/s, and still more preferably higher than 5 cm/s.
According to another aspect of the invention, in step e), solids generated under reaction are drawn from said second residue before the same is introduced in the reactor. The solids generated under reaction are removed by drawing a fraction of said second residue. Also removed with said drawing, in the same proportion with respect to the charge stock, is the possible dispersed catalyst contained therein.
According to another aspect of the invention, between step d) and step e), the hydroconversion method comprises the following steps:

f) centrifuging or decanting a fraction of the first residue where the solids generated under reaction are in suspension, so as to allow a drawing of said solids;
g) drawing and removing said solids generated under reaction;
h) subjecting said first residue lacking said solids to said concentration via vacuum distillation.

Analogous to that said in reference to the preceding aspect of the invention, with said drawing (of solids generated under reaction), also the possible dispersed catalyst contained therein is removed in the same proportion with respect to the charge stock.
According to another aspect of the invention, the accumulation factor is not less than 25.
If the hydroconversion method does not comprise steps f) to h) (i.e. if, in step e), the solids generated under reaction are drawn from said second residue before the same is introduced in the reactor), by "accumulation factor" it is intended the ratio between the flow rate with which, in step a), the heavy oil is introduced in the reactor, expressed in m³ per hour, and the flow rate with which, in step e), a fraction of said second residue is drawn to purge the solids, expressed in m³ per hour.
If the hydroconversion method comprises steps f) to h), by "accumulation factor" it is intended the ratio between the flow rate with which, in step a), the heavy oil is introduced in the reactor, expressed in m³ per hour, and the flow rate with which, in step f), said fraction of the first residue is drawn in order to be decanted or centrifuged, expressed in m³ per hour.
The accumulation factor is therefore dimensionless.
According to another aspect of the invention, in step a), when the reactor is of the type with expanded catalyst bed in which dispersed catalyst is also fed, the expanded catalyst bed only contains silica-alumina support, lacking catalytically-active metals for the purpose of hydrogenation, or it partially or completely lacks said silica-alumina support.
I.e.:

according to another aspect of the invention, in step a), when the reactor is of the type with expanded catalyst bed in which dispersed catalyst is also fed, the expanded catalyst bed contains silica-alumina support, lacking catalytically-active metals for the purpose of hydrogenation;
according to another aspect of the invention, in step a), when the reactor is of the type with expanded catalyst bed in which dispersed catalyst is also fed, the expanded catalyst bed only contains silica-alumina support, lacking catalytically-active metals for the purpose of hydrogenation;
according to another aspect of the invention, in step a), when the reactor is of the type with expanded catalyst bed in which dispersed catalyst is also fed, the expanded catalyst bed lacks silica-alumina support.

According to another aspect of the invention, in step a), when the reactor is of the type with dispersed catalyst or it is of the type with expanded catalyst bed in which dispersed catalyst is also fed, the catalyst bed only containing silica-alumina support, or the reactor is of the type with expanded catalyst bed in which only dispersed catalyst is fed (given that the catalyst bed lacks supported catalyst and silica-alumina support), the dispersed catalyst has molybdenum base and the replenishment of said dispersed catalyst is less than 100 ppm of metallic molybdenum with respect to the charge stock being fed.
As is known, by "dispersed catalyst with molybdenum base" it is intended molybdenum fed to the reactor through an oil-soluble compound thereof which, under reaction, generates the active species concordantly identified as "molybde-nite", i.e. "molybdenum sulfide".
As is known, by "replenishment of the dispersed catalyst" it is intended the replenishment made necessary by the removal of dispersed catalyst, which takes place together with the removal of the solids generated under reaction by said first or second residue, in accordance with the abovementioned aspects of the invention.
According to another aspect of the invention, when the reactor is of the type with expanded catalyst bed (into which dispersed catalyst is also possibly fed), the reactor operates at a degree of conversion not less than 95%.

Brief description of the figure

Further objects and advantages of the present invention will be clearer from the following detailed description of an embodiment thereof and from the enclosed figure, given as a mere non-limiting example, which schematically represents a system for hydroconversion of heavy oils in a single reaction stage, which employs a reactor with dispersed catalyst or with expanded catalyst bed (in which also dispersed catalyst is possibly fed or in which only dispersed catalyst is fed), in which the extraction of the reaction liquid through the biphasic effluent is enhanced by means of the introduction of gas in a zone containing biphasic foam at the head of the reactor, above the reaction liquid.
The scale and proportions of the various depicted elements do not necessarily correspond with actual scale and proportions.
Hereinbelow in the present description, the figure will also be illustrated with reference to elements not shown in the same.

Detailed description of several preferred embodiments of the invention

The figure shows a system for hydroconversion of heavy oils, provided with a single reaction stage, comprising a cylindrical reactor 4 of the type with dispersed catalyst or of the type with expanded catalyst bed (in which dispersed catalyst is also possibly fed or in which only dispersed catalyst is fed) provided with a system for introducing gas in a zone containing biphasic foam above the reaction liquid which, as will be illustrated hereinbelow in the present description, produces an entrainment of reaction liquid in the biphasic effluent.
Even if the system represented in the figure shows only one operating reactor, the reaction stage can be constituted by several reactors, like the reactor 4, in parallel.
The reactor 4 is fed with heavy oil through a line 1 and with hydrogen or gas containing hydrogen, through a line 2. Preferably, the heavy oil fed through the line 1 comprises at least 10% by weight of hydrocarbons having a boiling point higher than 540 °C so as to have sufficient foaming properties to generate a biphasic foam above the reaction liquid (traversed by hydrogen or by the gas containing hydrogen).
If the reactor 4 is of the type with expanded catalyst bed (in which also dispersed catalyst is possibly fed or in which only dispersed catalyst is fed), the hydrogen is premixed with the feeding of the heavy oil. In such case, the two fluids are distributed at the base of the reactor by means of a perforated plate (not shown in the figure) that supports the catalyst. The hydrogenation catalyst is deposited on a solid support, for example in the form of small cylinders or microspheres. The reactor 4 is fed with the supported catalyst in the upper part thereof through a line not shown in the figure. The reactor 4 is continuously or periodically fed with the supported catalyst to compensate for the spent catalyst that is withdrawn from the lower part of the reactor 4 through a line not shown in the figure.
If the reactor 4 is of the type with dispersed catalyst, the hydrogen is introduced at the base thereof by means of a nozzle distributor (not shown in the figure).
The reactor 4 is fed with the dispersed catalyst at the bottom thereof by means of a line 3, from which the catalyst admixes with the reaction liquid. The reactor 4 is fed with the catalyst to compensate for the quantity of catalyst that is removed with the purge of the solids. The catalyst can be introduced as is or by means of an oil-soluble precursor, i.e. a compound of metal (or metals) soluble in hydrocarbons, capable of generating the active species when it is in contact with the reaction liquid and the hydrogen. Catalysts are preferred with molybdenum base or molybdenum and iron base, possibly comprising silica-alumina or a zeolite compound.
The line 3 visible in the figure is therefore only present if the reactor 4 is of the type with dispersed catalyst or is of the type with expanded catalyst bed in which dispersed catalyst is also fed or is of the type with expanded catalyst bed in which only dispersed catalyst is fed.
The reactor 4 preferably operates at a temperature comprised between 330 °C and 430 °C, and at a pressure comprised between 10 MPa and 30 MPa. Under reaction conditions, in the upper part of the reactor 4, above the reaction liquid, a biphasic foam is produced which lifts reaction liquid towards the outlet where it generates a biphasic effluent that by means of a line 5 is fed to a gas-liquid separator 6 operating at the same pressure of the reactor 4. At the head of the separator 6, a gaseous flow 7 is obtained from which, via condensation, the light conversion products are recovered along with the excess hydrogen which, after a purification treatment, is recycled to the reactor 4. The reaction liquid, which constitutes the liquid phase present in the biphasic effluent, is collected, due to its density, at the bottom of the separator 6 together with the solids produced under reaction (such as coke, insoluble asphaltene resins and sulfides of the metals brought by the heavy oil). If the reactor 4 is of the type with dispersed catalyst (or of the type with expanded catalyst bed in which dispersed catalyst is also fed or in which only dispersed catalyst is fed), in the liquid at the separator 6 bottom there is also a quantity of dispersed catalyst with a concentration close to that of reaction. The separator 6 bottom liquid, with the solids produced by the reaction in suspension (and possibly the catalyst if the reactor 4 is fed with dispersed catalyst), is first sent, by means of a line 8, to a stage of flash-atmospheric distillation 9 from which the most volatile conversion products 10 are recovered, and subsequently sent, by means of a line 11, to a stage of concentration via vacuum distillation 12, with the extraction of the high-boiling conversion products 13 with a final boiling point of 540 °C, possibly lowered in order to obtain the quality specifications (% CCR and % insoluble asphaltenes in n-pentane, first of all) required for the subsequent treatments of hydrocracking and hydrotreating (not shown in the figure). The residue of the vacuum distillation is recycled to the reactor 4 by means of a line 14.
From the same line 14, by means of a line 15, a stream is derived that is used for removing the solids generated by the reaction and accumulated in the reaction liquid. This constitutes a first mode of removal from the hydroconversion system of the solids produced by the reaction. A second removal mode will be illustrated hereinbelow in the present description with reference to the lines 18, 20 and 21, and to the stage 19.
The flow rate of reaction liquid which, through the biphasic effluent, reaches the bottom of the separator 6 in order to feed the extraction of the high-boiling conversion products, depends on the superficial velocity of the hydrogen introduced at the base of the reactor 4. As stated above, said superficial velocity of the hydrogen cannot however be increased beyond a specific value. A hindered capacity of extraction of reaction liquid via biphasic effluent strongly limits the capacity of hydroconversion of the reactors with dispersed catalyst, as well as of the reactors with expanded catalyst bed of the prior art. In order to increase the aforesaid flow rate of reaction liquid (i.e. in order to increase the capacity of extraction of the reaction liquid via biphasic effluent), the reactor 4 is provided with a gas entrainment system adapted to facilitate the outflow of reaction liquid from the reactor 4 at the line 5 (i.e. the outlet duct of the reactor 4).
With reference to the figure, it is possible to observe that in the reactor 4, above the level of the reaction liquid, at the zone containing the biphasic foam, preferably at a height close to the base of the upper vault of the reactor 4 or within the vault itself, a line 16 for introducing gas is positioned, preferably by means of a nozzle grid. The gas preferably but not necessarily comprises hydrogen, and still more preferably hydrogen drawn before the purification treatment and/or recycled hydrogen and/or hydrocarbons at the gas state.
Above the gas introduction point, a densimeter 17 is installed which detects the density of the biphasic foam.
Owing to the entrainment of the reaction liquid following the introduction of gas, the transfer of reaction liquid into the biphasic effluent exiting at the head of the reactor 4 increases in proportion to the flow rate of gas introduced and in proportion to the density of the biphasic foam measured by the densimeter 17.
In particular, the flow rate of reaction liquid transferred into the biphasic effluent in relation to the flow rate of gas that enters into the vault (equal to the sum of the flow rate of the gas introduced at the line 2 and the flow rate of the gas introduced at the line 16) can vary from 0.5 kg to 5 kg of liquid per kg of gas, as a function of the density measured by the densimeter 17, in turn connected to the foaming properties of the charge stock.
The ratio between the diameter of the reactor 4 and the diameter of the outlet duct of the biphasic effluent is another parameter that determines the degree of the entrainment.
The temperature of the gas introduced by means of the line 16 is such that the temperature of the biphasic foam at the head of the reactor 4 is preferably comprised between 330 °C and 430 °C. At the same time as the introduction of the gas at the line 16, the flow rate of the residue of the vacuum distillation recycled at the bottom of the reactor 4 (at the line 14) is increased. The correct balance, between said recycled flow rate at the reactor bottom and the flow rate with which the gas at the line 16 is introduced, is verified when the density measured by the densimeter 17 is preferably comprised between 50 kg/m³ and 500 kg/m³, and still more preferably between 100 kg/m³ and 500 kg/m³. The flow rate of reaction liquid which is transferred to the biphasic effluent, and consequently to the bottom of the separator 6, corresponds with the flow rate of residue of the vacuum distillation fed to the reactor 4 bottom summed with the flow rate of the products generated from the conversion of the charge stock, present in liquid form in the reaction liquid.
The residue of the vacuum distillation circulated into reaction constitutes the vehicle through which the gases, introduced at the lines 2 and 16, transfer the conversion products to the liquid phase of the biphasic effluent, mainly high-boiling conversion products, present at the liquid state into reaction, to be subsequently recovered outside the reactor 4.
Operating a reactor, with dispersed catalyst or with expanded catalyst bed, of height H meters fed with heavy oil at a space velocity Vs (given by the ratio between the flow rate of the fed charge stock expressed in m³ per hour and the volume in m³ of the reactor), in order to counteract the accumulation of high-boiling conversion products into reaction liquid by means of the introduction of entrainment gas at the line 16, the residue of the vacuum distillation is fed to the bottom of the reactor 4 (at the line 14) with a unit flow rate (given by the ratio between the flow rate of the vacuum residue expressed in m³ per hour and the area of the cross section of the reactor expressed in m²) preferably at least equal to 0.5 x Vs x H, and still more preferably at least equal to Vs x H. With respect to the known hydroconversion systems, the hydroconversion unit capacity of the system thus assumes a value independent of the height of the reactor 4 and adaptable to Vs (of course within the limits allowed by the hydroconversion kinetics depending on the nature of the treated charge stock).
From a practical standpoint, the hydroconversion unit capacity is therefore no longer negatively affected by the height of the reactor 4 but is preserved even with the increase of the latter.
Whatever the flow rate of the residue of the vacuum distillation recycled at the bottom of the reactor 4 is, in order to continue to operate with the liquid-foam interface (and hence allow the degassing of the reaction liquid within the reactor), it is necessary to prevent, due to the flow rate of liquid at the reactor bottom, the bubbling liquid from being extended over the entire height of the reactor 4. In order to maintain the level of the reaction liquid positioned within the reactor 4, the superficial velocity of the gas uG (expressed in cm per second) exiting from the reactor 4 (given by the sum of the superficial velocity of the gas introduced at the base of the reactor 4 at the line 2 and the superficial velocity of the entrainment gas introduced in the reactor 4 above the level of the reaction liquid at the line 16) is preferably at least 10 times, and still more preferably at least 15 times, the velocity uL (likewise expressed in cm per second) of the reaction liquid exiting from the reactor 4 (at the line 5). In other words, preferably uG > 10 x uL, and still more preferably uG > 15 x uL.
Numerically uL is given by the ratio between the flow rate of reaction liquid exiting from the reactor 4 (which is recovered at the bottom of the separator 6, at the line 8) expressed in cm³ per second, and the area of the cross section of the reactor 4 expressed in cm². The condition uG > 10 uL, and still more the condition uG > 15 uL, involves an entrainment of liquid into the transport duct 5 of the biphasic effluent not less than the liquid coming from the reactor, which maintains the liquid-foam interface within the reactor itself.
At the line 16, flow rates of gas are introduced such to involve values of uG preferably greater than 2 cm/s, and still more preferably greater than 5 cm/s. By way of example, flow rates of gas are introduced such to involve values of uG comprised between 7 and 12 cm/s in order to extract the liquid conversion products generated by a reactor of height equal to 30 meters, fed with charge stock at a space velocity of 0.25 h^-1.
The superficial velocity of the gas at the line 16 is preferably comprised between 0.1 cm/s and 50 cm/s.
If the reactor 4 is of the type with expanded catalyst bed (in which also dispersed catalyst is possibly fed or in which only dispersed catalyst is fed), the possibility of increasing the quantity of high-boiling conversion products present in the biphasic effluent, operating on the flow rate of entrainment gas introduced at the line 16 and on the flow rate of the residue of the vacuum distillation recycled at the bottom of the reactor 4, allows operating the reactor 4 in a single reaction stage, avoiding the placement of several reactors in series one to the other in order to facilitate the extraction of the conversion products. Placing the non-converted charge stock of a first reactor in a second reactor, and possibly the non-converted charge stock exiting from the second reactor in a third reactor, involves conditions of increasing instability of the reaction liquid going from the first reactor to the final reactor. Said instability is due to the insolubility (and hence to the precipitation) of the asphaltene fraction which is accumulated with the progress of the reaction, and is the cause which limits the conversion degree obtainable in the reactors with expanded catalyst bed to levels that often do not exceed 70%. However, with the ability to operate the reactor as a single reaction stage, such conversion limit is overcome. The enhanced extraction of the reaction liquid, object of the present invention, is the factor that allows operating the reactors of the type with expanded catalyst bed in a single reaction stage, without penalizing capacity.
If it was of interest to limit the quantity of entrainment gas introduced at the top of the reactor at the line 16, a second flow of reaction liquid (not shown in the figure) can be extracted from a zone of the reactor containing bubbling liquid free of solids. The drawn liquid is sent to a degasser in order to obtain the reaction liquid from which the volatile conversion products are extracted via flash and distillation, and subsequently the high-boiling conversion products are extracted via vacuum distillation. In such case, the introduction of the gas at the head of the reactor 4 at the line 16 is carried out by respecting the above-indicated density limits of the biphasic foam (preferably between 50 kg/m³ and 500 kg/m³, and still more preferably between 100 kg/m³ and 500 kg/m³) and ensuring a ratio between the superficial velocity of the gas uG and the velocity of the liquid uL exiting at the head of the reactor 4 preferably of at least 10, and still more preferably of at least 15. The vacuum distillation extracts the conversion products contained both in the liquid phase present in the biphasic effluent from the head of the reactor 4, and in the liquid obtained directly from the bubbling zone. The unit flow rate of the residue of the vacuum distillation recycled at the bottom of the reactor 4 is, also in such case, tied to the height H of the reactor 4 and to the space velocity Vs with which the charge stock is fed.
The unit flow rate of the residue of the vacuum distillation recycled at the bottom of the reactor 4 is preferably greater than 0.5 x Vs x H, and still more preferably greater than Vs x H.
From the flows at the lines 7, 10, 13, the products of hydroconversion of the charge stock are obtained. The solids generated by the reaction are accumulated in the reaction liquid as solid phase in suspension (present in the different sections that constitute the hydroconversion system). As previously mentioned, said solids are removed by drawing, from the line 14, a fraction of the liquid suspension, residue of the vacuum distillation (where the solids are concentrated), and purging such fraction at the line 15. The liquid phase present in such stream 15 corresponds to the non-converted charge stock which, exiting as is from the hydroconversion system, determines the difference with respect to the 100% of the attainable conversion degree. In order to minimize the purge of non-converted charge stock and hence increase the conversion degree, it is necessary to reduce the formation of solids under reaction. In the stream at the line 15, the sulfides of the metals brought by the charge stock 1 are present, together with possibly the catalyst fed in dispersed form and the "fines" generated by the supported catalyst. In addition to such inorganic solids, solids of organic nature are usually present, but in much greater quantity; these are constituted by insoluble asphaltene resins and by coke, no longer convertible, to be removed. The production of insoluble resins and coke can be reduced until it is eliminated by lowering the reaction temperature in order to facilitate the hydrogenation reactions and at the same time prevent the undesired reactions that via dehydrogenation lead to the formation of insoluble resins and coke. The system of extraction of the conversion products at the liquid state via entrainment gas and vacuum recycling, since it assures a suitable extraction of the conversion products, even operating at low reaction temperature, is an enabling factor for operating the hydroconversion of a specific charge stock in the most favorable low reaction temperature conditions, without having to otherwise suffer unsustainable capacity reduction.
For a specific charge stock, a pre-established catalyst and a defined pressure of hydrogen at the reactor, in the hydroconversion system, object of the invention, the reaction temperature can be limited in a manner such to involve a formation of total solids, inorganic and organic, that is limited within 0.003 kg per kg of fed charge stock. One such production level of solids allows operating the purge at the line 15 at flow rates less than 5% of the charge stock feed flow rate, determining a conversion degree at least equal to 95%, regardless of whether reactors employed are of the type with dispersed catalyst or of the type with expanded catalyst bed.
Also visible in the figure is a second mode with which the removal of the solids produced by the reaction occurs, by deriving the liquid suspension that contains such solids before the concentration step. From the line 11, which feeds the stage 12 of vacuum extraction of the high-boiling conversion products, a flow 18 is derived that is subjected to centrifugation or settling by decanter at a stage indicated with the reference number 19. The centrifuge or decanter allows separating the solids (collected at a line 20 in order to be purged) from the liquid phase (drawn at a line 21) which is rejoined to the flow at the line 11 in order to feed the vacuum extraction of the high-boiling conversion products.
Operating the removal of the solids by following this second mode, the liquid purge stream at the line 15 can be omitted, thus eliminating the purge of non-converted charge stock.
By operating the purge of the solids in suspension phase, at the line 15, or in solid form, at the line 20, the solids produced by the reaction, along with the dispersed catalyst fed to the reactor 4 (if the latter is of the type with dispersed catalyst or is of the type with expanded catalyst bed in which dispersed catalyst is also fed or in which only dispersed catalyst is fed), are accumulated in the hydroconversion system.
Incidentally, by asserting that the purge of the solids is carried out "in suspension phase" it is intended that the solids are suspended in the liquid flow at the line 15. From the latter, part of the liquid flows out that comes from the bottom of the vacuum distillation, which contains the solids in suspension (i.e. in suspended or dispersed form).
By asserting that the purge of the solids is carried out "in solid form" it is intended that the solid phase, separated by centrifuge or decanter, is removed at the line 20.
Whether the purge of the solids is operated in suspension at the line 15, or whether the solids are removed in solid form at the line 20, it is possible to define an "accumulation factor". In particular, by operating the purge of the solids in liquid suspension at the line 15, the accumulation factor is given by the ratio between the flow rate of charge stock being fed at the line 1 and the flow rate of the stream at the line 15. Analogously, by operating the purge of the solids at the line 20, the accumulation factor is given by the ratio between the flow rate of charge stock being fed at the line 1 and the flow rate of the liquid suspension at the line 18.
For a specific flow rate of feed charge stock, the flow rate of the stream 15 or the flow rate of the stream 18 (depending on the mode used for removing the solids) is minimized for the purpose of maximizing the aforesaid accumulation factor. High values of the accumulation factor, even if not involving significant advantages in terms of attainable conversion degree (given that this is in any case higher than 95%, and close to 100% when one proceeds with the removal of the solids in solid form at the line 20), nevertheless offer the great advantage of minimizing the replenishment of catalyst in the reactors with dispersed catalyst, in the reactors with expanded catalyst bed in which dispersed catalyst is also fed and in the reactors with expanded catalyst bed in which only dispersed catalyst is fed. The system of accumulation-removal of solids described herein, in suspension, via stream 15 or in solid form, at the line 20, involves concentrations of catalyst in the reaction liquid that are increased with respect to the metered catalyst amounts, referred to the charge stock being fed.
By operating the removal of the solids at the line 20, with an accumulation factor at the line 18 equal, by way of example, to 25, a metering of catalyst, e.g. molybdenum, of 50 ppm, referred to the charge stock being fed, involves a concentration of catalyst in the line 18 of 1250 ppm corresponding with a concentration of catalyst in the reaction liquid within the reactor 4 (which in terms of solids is less concentrated than the flow at the line 18) of around 1000 ppm. Operating then with accumulation factors at the line 18 of at least 25 (allowed by low formation of solids under reaction), catalyst concentrations are obtained into reaction equal to at least 20 times the metering of dispersed catalyst referred to the charge stock being fed. A specific value of the accumulation factor, detected at the line 18, is associated with a numerically larger value of the accumulation factor detected at the line 15. Hereinbelow in the present description, the accumulation factor will always be that detected at the line 18.
If the reactor 4 is of the type with dispersed catalyst or of the type with expanded catalyst bed in which dispersed catalyst is also fed or is of the type with expanded catalyst bed in which only dispersed catalyst is fed, by following the system of accumulation and removal of solids described above, even if operating at low catalyst replenishment (i.e. with low catalyst consumption), the concentration of the dispersed catalyst in the reaction liquid reaches levels such to ensure the attainment of the maximum catalytic effect, hence rendering uselessly expensive the simultaneous presence of metals in supported form catalytically-active for the hydrogenation (such as molybdenum, chromium, vanadium) when reactors employed are of the type with expanded catalyst bed in which dispersed catalyst is also fed. In such case, the expanded catalyst bed is occupied by the structured support (small cylinders or microspheres), constituted only by silica-alumina or another acid reaction material (like the zeolites), in order to support the denitrification of the charge stock. The structured solid acid reaction material, confined inside the reactor, no longer performing the function of support of the catalytically-active metals for the purpose of hydrogenation, can be present in significantly reduced quantity or be totally eliminated, with consequent recovery of useful reaction volume and reduction of the formation of "fines" to be removed. If the structured solid material is totally eliminated, the reactor of the type with expanded catalyst bed is only fed with dispersed catalyst. The acid reaction component (silica, alumina or zeolite compound), if of interest, can be fed in dispersed form. In such case, the recirculation of the reaction liquid by means of the "ebullating pump" can be limited or omitted.
In the following claims, with the expression "reactor with expanded catalyst bed" it is intended a reactor with expanded catalyst bed in which dispersed catalyst is also fed or a reactor with expanded catalyst bed in which only dispersed catalyst is fed (i.e. in which the silica-alumina structured solid material is absent).

Claims

Method for hydroconversion of heavy oils in a single reaction stage comprising the following steps:
a) feeding, at the base, at least one reactor (4) of the type with dispersed catalyst or of the type with expanded catalyst bed with:
• heavy oil comprising at least 10% by weight of hydrocarbons having, at atmospheric pressure, a boiling point higher than 540 °C;

• a first gas including hydrogen;

• a hydroconversion catalyst,
generating:
• at least at a lower part of said reactor (4), reaction liquid traversed by said first gas, and

• at least at an upper part of said reactor (4) placed above said reaction liquid traversed by said first gas, a biphasic foam including at least one liquid phase and one gaseous phase, said biphasic foam originating a biphasic effluent upon exiting from the head of said reactor (4),

b) drawing said biphasic effluent from said upper part of the reactor (4) and separating the liquid phase from the gaseous phase of said biphasic effluent;

c) subjecting said liquid phase to flash-atmospheric distillation;

d) subjecting to concentration, via vacuum distillation, a first residue of said flash-atmospheric distillation;

e) introducing, at the base, in said reactor (4) a second residue of said concentration via vacuum distillation,
the hydroconversion method being characterized in that , in step a), a second gas is introduced into said upper part of the reactor (4), at a zone of said reactor (4) containing said biphasic foam,
in addition:
• said reactor (4) operating at a temperature comprised between 330 °C and 430 °C, and at a pressure comprised between 10 MPa and 30 MPa;

• said second gas being introduced into said upper part of the reactor (4) in step a), and said second residue being introduced into said reactor (4) in step e), with respective flow rates such that the density of said biphasic foam measured following the introduction of said second gas in said upper part of the reactor (4), is comprised between 100 kg/m³ and 500 kg/m³;

• in step a), said second gas being introduced into said upper part of the reactor (4) at a superficial velocity comprised between 0.1 cm/s and 50 cm/s;

• in step b), said biphasic effluent being originated with a superficial velocity of the gas equal to at least 10 times the liquid flow velocity;

• in step a), said second gas being introduced into said upper part of the reactor (4) at a temperature such that the temperature of said biphasic foam is comprised between 330 °C and 430 °C;

• in step e), said second residue being introduced into said reactor (4) at a unit flow rate at least equal to 0.5 x Vs x H, where Vs is the space velocity with which, in step a), said heavy oil is introduced into said reactor (4) and H is the height of said reactor (4).
Hydroconversion method according to claim 1, characterized in that, in step b), said biphasic effluent is originated with a superficial velocity of the gas equal to at least 15 times the liquid flow velocity.
Hydroconversion method according to claim 1, characterized in that, in step e), said second residue being introduced into said reactor (4) at a unit flow rate at least equal to Vs x H.
Hydroconversion method according to claim 1, characterized in that the accumulation factor is not less than 25.
Hydroconversion method according to claim 4, characterized in that, in step a), when said reactor (4) is of the type with expanded catalyst bed in which dispersed catalyst is also fed, said expanded catalyst bed contains silica-alumina support, lacking catalytically-active metals for the purpose of hydrogenation.
Hydroconversion method according to claim 5, characterized in that, in step a), when said reactor (4) is of the type with expanded catalyst bed in which dispersed catalyst is also fed, said expanded catalyst bed only contains silica-alumina support, lacking catalytically-active metals for the purpose of hydrogenation.
Hydroconversion method according to claim 4, characterized in that, in step a), when said reactor (4) is of the type with expanded catalyst bed in which dispersed catalyst is also fed, said expanded catalyst bed lacks silica-alumina support.
Hydroconversion method according to claim 4, characterized in that, in step a), when said reactor (4) is of the type with dispersed catalyst or it is of the type with expanded catalyst bed in which dispersed catalyst is also fed, said catalyst bed only containing silica-alumina support, or it is of the type with expanded catalyst bed in which only dispersed catalyst is present, said dispersed catalyst contains molybdenum and the replenishment of said dispersed catalyst is less than 100 ppm of metallic molybdenum with respect to the charge stock being fed.
System for hydroconversion of heavy oils in a single reaction stage comprising:
• a reactor (4) of the type with dispersed catalyst or of the type with expanded catalyst bed;

• a first line (1) for feeding said reactor (4) with heavy oil comprising at least 10% by weight of hydrocarbons having, at atmospheric pressure, a boiling point higher than 540 °C;

• when said reactor (4) is of the type with dispersed catalyst, a second line (2) for feeding said reactor (4) with a first gas including hydrogen,
when said reactor (4) is of the type with expanded catalyst bed, said heavy oil is premixed with a first gas including hydrogen;

• a third line (3) for feeding said reactor (4) with a hydroconversion catalyst, said reactor (4) containing:
- at least at a lower part of said reactor (4), reaction liquid traversed by said first gas, and

- at least at an upper part of said reactor (4), placed above the reaction liquid traversed by said first gas, a biphasic foam including at least one liquid phase and one gaseous phase, said biphasic foam originating biphasic effluent upon exiting from the head of said reactor (4);

• a separator (6) suitable for separating the liquid phase from the gaseous phase of said biphasic effluent;

• a fourth line (5) for drawing said biphasic effluent from said upper part of the reactor (4), and for introducing said biphasic effluent in said separator (6);

• a first flash - distillation stage (9);

• a fifth line (8) for drawing said liquid phase from said separator (6) and for introducing said liquid phase in said first stage (9);

• a second stage (12) for concentrating via vacuum distillation;

• a sixth line (11) for drawing, from said first stage (9), a first residue of said first stage (9), and for introducing said first residue in said second stage (12);

• a seventh line (14) for drawing, from said second stage (12), a second residue thereof, and for introducing said second residue in said reactor (4) at the base thereof,
the hydroconversion system being characterized in that it also comprises:
• an eighth line (16) for introducing a second gas in said reactor (4) at a zone thereof containing said biphasic foam;

• a densimeter (17) suitable for measuring the density of said biphasic foam in the upper part of said reactor (4), in an intermediate position between the gas introduction at said eighth line (16) and said drawing at said fourth line (5).
Hydroconversion system according to claim 9, characterized in that it comprises a nozzle grid by means of which said second gas can be introduced into the upper part of said reactor (4).