WO2019220471A1

WO2019220471A1 - Hydroconversion of heavy oils at improved hydrogenation rate and evaporation capacity

Info

Publication number: WO2019220471A1
Application number: PCT/IT2018/000091
Authority: WO
Inventors: Luigi Patron
Original assignee: Luigi Patron
Priority date: 2018-05-16
Filing date: 2018-06-27
Publication date: 2019-11-21
Also published as: IT201800005448A1

Abstract

The invention relates to the hydroconversion of heavy oils by means of a slurry reactor where the bubbling liquid is replaced by a foam obtained at a high gas velocity/liquid velocity ratio, which facilitates the evaporation of the conversion products. The foam state of the mixture of gas and reaction liquid increases the hydrogenation rate. The combination of the two effects improves the productivity of the reactor.

Description

Hydroconversion of heavy oils at improved hydrogenation rate and evaporation capacity

Field of application of the invention

The present invention relates to the hydroconversion of heavy oils by means of slurry reactors.

Review of the background art

In hydroconversion systems, regardless of the type of reactor employed, be it an ebullated catalytic bed reactor or a slurry bubble column reactor, the distribution of hydrogen gas at the base of the reactor produces, with the reaction liq- uid, a mixture of gas bubbles and liquid (bubbling liquid), in bubbling fluid dynamic regime or bubbling regime. Such a bubbling liquid, when the gas is introduced at the base of the reactor at a superficial velocity of 5 - 5.5 cm/s (cm³/s divided by the reaction section in cm², or cm/s), is characterized by a unit fraction e_g of volume occupied by gas bubbles approximately equal to 1/3, and therefore by a unit fraction of volume occupied by the liquid equal to 2/3 (see, for example, Figure 5, US patent 5308476).

The hydroconversion catalysts contain one or more transition metals which have the aim of specifically promoting hydrogenation. If the catalyst employed is of the dispersed type (slurry reactor), such metals are introduced into the reac- tion in the form of oil-soluble compounds which, by reaction with the sulfur generally released by the charge itself, are transformed into the respective sulfides, finely dispersible in the reaction liquid. For example, if the oil-soluble catalyst consists of a molybdenum compound, by reaction with the sulfur, possibly released by the charge itself, tetravalent molybdenum, or molybdenite, is generated (S=Mo=S), which is thus found finely dispersed in the reaction liquid (ELSEVIER - Journal of catalysis, 308 (2013) 189-200, G. Bellussi et al.:“Hydroconversion of heavy residues in slurry reactors”). When dissolved hydrogen is present in the reaction liquid, molybdenite is converted into molybdenum dihydrosulfide (HS-Mo-SH) with the simultaneous reduction of molybdenum in the di-valent state. The higher the concentration of molecular hydrogen dissolved in the reaction liquid is, the more the S=Mo=S vs HS-Mo-SH equilibrium is shifted to the right. The hydrogen sums to the unsaturated hydrocarbons present by means of the hydrogenated molybdenum di-hydrosulfide which is thus the hydrogenating agent determining the hydrogenation rate. When the two hydrogen atoms are transferred to an unsaturated bond, the molybdenum di-hydrosulfide returns molybdenite. Molybdenite, S=Mo=S, if it remains as such and is not constantly reduced to molybdenum di-hydrosulfide by the hydrogen dissolved in the reaction liquid, catalyzes the reverse dehydrogenation reaction, in addition to generating radical precursors of the formation of coke, unlike the molybdenum di-hydrosulfide which, on the contrary, acts as hydrogenation agent and as radical scavenger, by means of the -SH function. The greater the concentration of molecular hydrogen dissolved in the reaction liquid is, the more likely will hydrogenation prevail over dehydrogenation, and the capture of radicals equally prevail over the generation thereof, so that the reduced form of molybdenum HS-Mo-SH will prevail, to the detriment of molybdenite S=Mo=S. The prevalence of the catalytic form HS-Mo-SH may occur if the hydrogen, that diffuses from the gaseous phase in a determined volume of reaction liquid, can equal the amount of hydrogen which, in that volume of reaction liquid, is consumed for hydrogenating unsaturated structures and forming gases and light products generated by thermal cracking. The hydrogen fed at the base of the reactor gives origin to the bubbles which form the gaseous phase of the bubbling liquid. The hydrogen, to react with the unsaturated hydrocarbons, must first diffuse from the gaseous phase to the liquid one, where it will then be dissolved. Once the temperature is set, the flow of hydrogen diffusing in the reaction liquid is proportional to the partial pressure of the hydrogen and to the area of the gas- liquid interface. In a gas-liquid bubbling column characterized by a unit fraction e_g of volume occupied by gas bubbles (“gas holdup”) having an average diameter d_s (expressed in cm), the specific surface of the gas-liquid bubbles interface a_s (expressed in cm² of surface per cm³ of liquid), for geometrical reasons, takes the value 6 e_g/d_s (“Process Biochemistry” 40 (2005) 2263-2283,“Bubble column reactors”, N. Kantarci et al., page 2275 (5)). This is a limitation which, once the gas distribution is optimized, inevitably affects the hydrogenation capacity of the system, whatever the type of catalyst employed is. In particular, in the case of a slurry catalyst, the limited hydrogen diffusion rate does not allow a catalytic activity thereof proportional to the concentration to which it may be used, strongly limiting the potential thereof.

What is described above refers to the chemical aspect of hydroconversion. Once the heavy oil is converted into conversion products (distillable hydrocarbons, exempt from carbon residue), said products should evaporate from the reactor at the same rate at which they are generated. If this does not occur, the high-boiling conversion products (fraction having a boiling point indicatively from 300°C to 540°C) accumulate in the reaction liquid, unproductively occupying part of the reactor, to the detriment of productivity. This is what occurs, in practice, as a result of the limited value of the superficial velocity (gas flow rate in cm³/s, referred to the reactor pressure and temperature conditions, divided by the reaction section in cm², therefore cm/s) at which the hydrogen may be introduced into the column to operate under a bubbling regime. The maximum superficial velocity threshold is about 5 cm/s (“Process Biochemistry” 40 (2005) 2263-2283,“Bubble column reactors”, N. Kantarci et al., page 2269), rising to 5.5 cm/s, perfecting the gas distribution (see Figure 5 of the aforementioned US patent 5308476). Even when operating at such maximum superficial velocity values, the flow of hydrogen which may be introduced into the reaction is inadequate to produce the evaporation of the high-boiling conversion products which, therefore, accumulate in the reaction liquid to reach concentrations of 50% and beyond. The inadequacy of the vaporization is found to be more marked as the reactor height increases, to the extent of severely limiting the productivity thereof. Due to the limited amount of high-boiling conversion products that can be vaporized during reaction, the high-boiling conversion products are obtained from the reaction liquid, where they are concentrated, by means of distillation with an under vacuum final stage in a separate section. Alternatively, the reaction liquid may be sent to a next reactor in series with the first one.

The evaporation of the high-boiling conversion products from a column of bubbling liquid is therefore a further factor limiting the productivity of hydroconversion systems.

Hydroconversion systems employing a mixture of gas bubbles and reaction liquid under a bubbling regime are therefore conditioned by:

• a limited surface through which hydrogen may diffuse from the gas phase to the liquid phase, a factor limiting hydrogenation;

• insufficient values of the superficial velocity at which the gas may be introduced at the base of the reactor, a factor limiting the evaporation in reaction of the high-boiling conversion products.

The first point, related to the diffusibility of hydrogen in two-phase bubbling regime systems, used in the heavy oil hydroconversion, appears to have not been dealt with up to now.

The enhancement of the evaporation capacity in hydroconversion systems under a bubbling regime was instead addressed in US patent 8236170. The evaporation of high-boiling conversion products is obtained by the combined effect of temperature and a high superficial velocity of the hydrogen introduced at the base of the reactor. Since an increase in the superficial velocity of the gas in a bubble column reactor would provoke the phenomenon of the coalescence of the gas bubbles, jeopardizing the uniformity of the bubbling regime which is instead to be maintained, a liquid circulation reactor is employed to operate at a high superficial velocity of the liquid.

The high flow rate of liquid thus required makes, however, the gas-liquid separation step complex and expensive.

In patent application PCT/IT2015/000247, owned by the Applicant, a hydroconversion system is described employing a slurry reactor in which the introduction of gas at the head of the reactor improves the extraction of the high-boiling con- version products in the liquid state. The improved extraction of the high-boiling conversion products by evaporation during the reaction is not dealt with.

Objects of the invention

With reference to slurry reactor hydroconversion systems, the present invention overcomes both the hydrogenation rate limitation and the evaporation capacity limitation, at least partially replacing the bubbling liquid of the reactor by a mixture of gas and reaction liquid in the foam state, under a foaming fluid dynamic regime or foaming regime. Incidentally, a two-phase gas-liquid system in the foam state is characterized by a unit volumetric fraction of liquid not higher than 0.5 (cm³ of liquid per cm³ of foam).

Therefore, it is the object of the present invention to indicate a method and system for the hydroconversion of heavy oils with improved hydrogenation rate and improved evaporation capacity of the conversion products using a slurry foam column reactor.

Summary of the invention

The invention relates to a slurry reactor fed with gas at high superficial velocity and concomitantly fed with reaction liquid collected from the bottom of a phase separator, so as to generate, above the gas introduction, a mixture of gas and reaction liquid in the foam state, under a foaming regime. The gas fed at high superficial velocity increases the evaporation of the high-boiling conversion products during reaction. The foam state of the mixture of gas and reaction liquid involves a gas-liquid interface with a high specific surface which increases the hydrogen diffusion and the hydrogenation rate.

Another aspect of the invention relates to a method for operating the slurry foam column reactor of the present invention.

Further innovative features of the present invention are described in the dependent claims.

Brief description of the drawing

Further objects and advantages of the present invention will become apparent from the following detailed description of an embodiment thereof and from the accompanying drawing, given merely by way of explanation and not by way of limitation, which diagrammatically shows a hydroconversion system in an em- bodiment in which, in at least part of the slurry reactor, a mixture of gas and reaction liquid in the foam state, at a high superficial velocity of the gas, replaces the bubbling liquid. The scale and proportions of the various elements shown do not necessarily correspond to the real ones.

Detailed description of some preferred embodiments of the invention

A heavy oil, of mineral origin, due to the fact it contains heteroatoms with a surfactant effect, such as S, N and O, possesses itself surfactant properties. Heavy oils, such as for example crude oil, bitumen, tar sands oil, shale oils and the atmospheric distillation, vacuum distillation, thermal visbreaking residues thereof, ebullated bed residues and solvent deasphalting residues, by virtue of the foaming properties they possess (also possessed by the hydrocarbon mixtures containing them), allow, as described herein, to replace at least partially the bubbling liquid of the slurry reactors employed in the hydroconversion by mixtures of gas and reaction liquid in the foam state, under foaming fluid dynamic regime.

A conventional slurry reactor is fed at the base with the heavy oil to be converted, at a unit flow rate F (cm³/s divided by the reaction section in cm², therefore with superficial velocity F cm/s), with a precursor of the dispersed catalyst, preferably in the form of an oil-soluble compound of at least one transition metal (preferably comprising molybdenum), and with hydrogen, or a gas including hydrogen, at a superficial velocity uG (expressed in cm/s) within the bubbling regime threshold of 5.5 cm/s. The gas is preferably distributed by means of a nozzle grid, or an equivalent means, placed at the base of the reactor at a distance H from the gas exit mouth placed on or in the head vault of the reactor. The homogeneity of the reaction medium, material and thermal, is ensured by the gas-liquid fluid dynamics alone produced by the upward movement of the gas bubbles, without the need for employing mixing devices.

In the vicinity of the upper vault, at the surface delimiting the bubbling liquid, the gas bubbles detach and a two-phase flow in the foam state is simultaneously generated. At the mouth of the output line present on the vault or in the vault of the reactor, the foam transforms into a two-phase gaseous flow (reaction liquid, containing the catalyst and the solids generated by the reaction, dispersed in the gaseous phase) exiting from the head of the reactor. From the two-phase gaseous flow originating from the reactor, by means of the treatment in a phase separator, a gaseous phase is obtained at the head containing mainly low- boiling conversion products and a liquid phase is obtained at the bottom consisting of reaction liquid where the high-boiling conversion products are concentrated. In a hydroconversion system with a single reaction stage and recycle, to recover the high-boiling fraction of the conversion products, the reaction liquid at the separator bottom is collected, depressurized and subjected to distillation outside of the reactor, with an under vacuum final stage. The vacuum residue is sent back into the reaction where it is again enriched with high-boiling conversion products, giving rise to the flow of reaction liquid, which collects at the bottom of the separator, with an indicative flow rate of up to 1.5 F (below, for simplicity, such a flow rate will be indicated as 1.5 F). A fraction of vacuum residue is purged to remove the solids generated by the reaction.

In a slurry reactor, fed with heavy oil, the introduction of gas at high superficial velocity, due to the foaming properties of such oils, produces foam and the conveying of reaction liquid out of the reactor. The conveying entity grows progressively with the superficial velocity of the gas until zeroing the bubbling liquid which is replaced by foam.

In order to regularly operate a slurry reactor at a high superficial velocity of the gas, the liquid exit is compensated by introducing reaction liquid collected at the bottom of the phase separator into the reactor.

A column reactor, operating at a temperature from 380°C to 440°C and at a pressure from 10 MPa to 30 Mpa and fed at the base with heavy oil with a unit flow rate F (cm³/s divided by the reaction section in cm², therefore cm/s) and with an oil-soluble precursor of a dispersed catalyst, as described above, is fed with reaction liquid originating from the bottom of the phase separator with a unit flow rate Q (cm³/s divided by the reaction section in cm², therefore cm/s) and, simultaneously, by means of the grid at the base of the reactor, at a distance H from the exit mouth on or in the head vault of the reactor itself, with gas including hydrogen introduced at a superficial velocity uG (cm/s). The unit flow rate Q of the liquid collected from the phase separator may take values from 0.5 to 10 times F. Such a flow Q sums to the liquid flow with a unit flow rate 1.5 F generated by the reactor, giving rise to a liquid flow exiting from the vault of the reactor with a unit flow rate L, equal to Q + 1.5 F. The value of the unit flow rate L, being expressed in cm³/s divided by the reaction section in cm², therefore in cm/s, also provides the value of the superficial velocity uL with which the liquid rises in the reactor to exit from the vault. The superficial velocity of the gas and of the liquid refer to the output from the cylindrical area, at the base of the upper vault of the reactor, i.e., at the inner section of the reactor which separates the cylindrical part from the head vault. The superficial velocity uG of the gas is at least 10 times the superficial velocity uL of the liquid, i.e. uG / uL > 10 (or, equivalently, uL / uG <0.1 ), a condition which implies the presence of a mixture of gas and reaction liquid in the foam state, under foaming fluid dynamic regime, in the space above the grid, replacing the bubbling liquid.

The liquid fraction present in the foam above the grid remains permanently sus- pended if the momentum it receives from the gas (proportional to the square of the superficial velocity of the latter and to the pressure at which the reactor operates, i.e., to the density of the gas itself, therefore to the mass thereof) is capable of counteracting the natural fallout thereof due to gravity (drainage). The column reactor of the present invention is fed with reaction liquid originating from the bottom of the phase separator at a unit flow rate Q generating a flow L of liquid exiting the vault of the reactor, equal to Q + 1.5 F. The introduction of gas at a superficial velocity uG, such that it results in uG > 10 uL, maintains the mixture of gas and reaction liquid above the grid in the foam state.

The gas introduced into the reactor transports the liquid present in the foam towards the exit mouth at the vault of the reactor. Due to the acceleration at which the gas is subjected to in the vicinity of the exit mouth, suitably calibrated, the disruption of the foam occurs with the consequent fallout of the preponderant part of the liquid in the underlying foam, while the remaining part of liquid is reduced in drops which remain dispersed in the two-phase gaseous flow which leaves the reactor and which is then treated at the phase separator. The reaction liquid originating from the bottom of the separator, which instead rises towards the vault, counterposes such a liquid which falls out from the vault of the reactor. Such counterposed liquid flows combine and mix at the“Plateau borders” of the polyhedral cells enclosing the gas present in the foam, inducing a mixing effect which produces a material and thermal uniformity equivalent to that produced by the gas bubbles in a bubbling column. A column in which the mixture of gas and reaction liquid is in the foam state, under foaming fluid dy- namic regime, may therefore replace, from the aspect of the material and heat transfer capacity, a bubbling column in a reactor employed in the hydroconversion of heavy oils.

With respect to the bubbling liquid which is replaced, the mixture of gas and reaction liquid in the foam state, under a foaming regime, is characterized by a gas-liquid interface with a significantly larger specific surface, being foam, and allows to introduce gas at high superficial velocity so as to increase the evaporation of the high-boiling conversion products in the reaction. The superficial velocity uG at which the gas is introduced is thus greater than the maximum gas velocity threshold under a bubbling regime of 5.5 cm/s, being capable of rising up to 15 cm/s, and preferably rising up to 30 cm/s and above, so as to increase the evaporation of high-boiling conversion products in the reaction.

The additional extraction of high-boiling conversion products by evaporation, thus carried out, may reach a significant fraction of F, such as to produce the evaporation of the totality of the high-boiling conversion products generated by the hydroconversion process, directly in the reaction. In such a case, the treatment outside of the reactor, lacking the need to extract the high-boiling conversion products under vacuum, may be limited to what is needed in order to purge the solids generated by the reaction, being capable of taking a reduced size as a consequence.

The unit volumetric fraction of liquid dispersed in the two-phase gaseous flow emerging from the head of the reactor may be expressed as uL / (uG + uL). Since uL/uG must be < 0.1 , the volumetric fraction of liquid in the two-phase gaseous flow is, itself, lower than 0.1. Conversely, the unit volumetric fraction of liquid present in the mixture of gas and reaction liquid in the foam state above the grid may take values up to 0.5. The difference between the values of the liquid fraction present in the foam (inside the vault) and in the two-phase gaseous flow (in the vault output line), is related to the amount of liquid falling out from the vault at the disruption of the foam. In order to maximize the liquid fraction present in the mixture of gas and reaction liquid, in the foam state, under a foaming regime, and therefore the liquid filling degree of the column, the user should intervene on the calibration of the gas exit mouth at the head of the reactor and act on the flow rate Q.

In one embodiment, the unit volumetric fraction of liquid in the gas-reaction liquid mixture in the foam state, under a foaming regime, is greater than 0.2. The value of said volumetric fraction may be obtained from the instrumental detection of the density of the gas-reaction liquid mixture at the cylindrical portion of the reactor, once the density of the liquid and of the gas at reaction conditions are known.

In one embodiment, the unit volumetric fraction of liquid in the gas-reaction liquid mixture in the foam state, under a foaming regime, is greater than 0.4.

In one embodiment, a second grid, above the first one, is placed at a distance h (smaller than H) from the exit mouth at the head vault of the reactor. The column reactor is always fed at the base with heavy oil, at a unit flow rate F, and with a dispersed catalyst. The reactor is also fed with the reaction liquid originating from the bottom of the phase separator at a unit flow rate from 0.5 to 10 times F. Gas at a superficial velocity uGi is introduced at the lower grid, within the maximum bubbling regime velocity threshold, while gas at a superficial velocity UG2 is introduced at said second grid. The superficial velocity uG of the gas at the output from the cylindrical portion, at the base of the upper vault of the reactor, corresponds to uGi + UG2.

The feeding of a unit flow rate Q of reaction liquid collected from the phase separator generates a liquid flow L, in output from the reactor, equal to Q + 1.5 F, to which corresponds a superficial velocity uL of the liquid in the reactor. The superficial velocity UG2 with which the gas is introduced at the second grid is such that uG > 10 uL, so that above the second grid a mixture of gas and reaction liquid in the foam state, under foaming fluid dynamic regime, is present. The part of the reactor between first and second grids is instead under a bubbling regime. The superficial velocity uG is greater than the maximum bubbling re- gime velocity threshold of 5.5 cm/s, being capable of rising up to 15 cm/s, and preferably up to 30 cm/s and above, so as to increase the evaporation of high- boiling agents in the reaction.

The temperature of the bubbling liquid and the temperature of the mixture of gas and reaction liquid in the foam state, under a foaming regime, are set in the range of 380°C to 440°C, possibly independently.

In an alternative configuration, to further increase the evaporation occurring in the foaming area, the gas introduced into the second grid has a temperature such as to heat the mixture of gas and reaction liquid in the foam state above 440°C up to a maximum of 480°C. Thereby, in the foaming area above the second grid, due to the increase in temperature, in addition to a greater evaporation, thermal cracking is also produced, which may thus be regulated independently of the temperature of the bubbling area.

If the gas-liquid separation is carried out by means of a conventional separator, the reaction liquid at the separator bottom, as a result of the temperature employed in the reaction, of the absence of dissolved hydrogen (in a phase separator the diffusion of hydrogen from the above gas phase to the underlying liquid is negligible) and therefore of the presence of the catalyst in the form of sulfide at a higher valence state (molybdenite in case the catalyst includes molybdenum), is subjected to dehydrogenation which leads to the undesired formation of asphaltene resins and coke. To avoid this, the phase separator is preferably provided with a line for introducing a cooling hydrocarbon into the separated liquid, capable of lowering the temperature thereof to below 400°C, preferably below 380°C. The cooling hydrocarbon introduced into the separator bottom liquid will have a boiling point preferably from 50°C to 300°C, so that, having to then enter the reactor bottom, it does not accumulate in the reaction liquid. Such a hydrocarbon, recoverable from the gaseous flow condensates at the head of the separator, will form a hydrogen source, so that in the reaction liquid at the separator bottom the reduced form of molybdenum HS-Mo-SH may prevail to the detriment of molybdenite to prevent, therefore, the formation of coke.

The ratio between the hourly flow rate with which the liquid charge to be con- verted is fed and the volume of the reactor (m³/h divided by m³, i.e., h ¹, referred to as the liquid hourly space velocity Vs) is an independent variable which, in once through systems, may be placed in a wide range of values. In the present hydroconversion system, which is of the closed type with a single reaction stage and recycle, the liquid hourly space velocity at which the charge may be fed becomes necessarily an observable value, since it must coincide, with the exception of the purge, with the unit conversion capacity of the system, resulting in this being mainly dependent on the content of the carbon fraction present in the oil to be converted. Depending on the type of charge, Vs values from 0.1 to 0.25 h ¹ are observed, when the feeding consists of a heavy oil vacuum residue. Reference is now made to the attached figure to further illustrate the embodiment of the invention.

The Figure diagrammatically shows an embodiment of a system for the hydroconversion of heavy oils in which the bubbling liquid of the slurry reactor is replaced by a mixture of gas and reaction liquid in the foam state, under foaming fluid dynamic regime. The reactor 1 consists of a column at the bottom of which, at a unit flow rate F, heavy oil 2 and an oil-soluble precursor of a catalyst 3 are fed. At the base of the reactor, at a distance H from the output mouth of the upper vault, a nozzle grid 4, or an equivalent means, is present, ensuring a uniform distribution of hydrogen, or of gas containing hydrogen, at a surficial velocity uG. At the upper vault of the column an output line 5 is present, at which mouth, suitably calibrated, the disruption of the foam is produced with the consequent dispersion of the liquid in the two-phase gaseous flow which is then fed to a phase separator 6. At the head of the phase separator 6, by means of the line 7, the gaseous phase containing a fraction of conversion products, mainly with a low boiling point, exits. At the bottom of the phase separator 6, the liquid phase is collected, consisting of reaction liquid where high-boiling conversion products are dissolved and in which the solids generated by the reaction as well as the finely dispersed catalyst are dispersed. A separator bottom liquid flow, collected by means of line 8, is depressurized by means of a depressurization valve 17 and sent to a flash and distillation treatment, with final under vacuum stage, outside the reactor 9, which has the dual function of extracting high- boiling conversion products 10 and producing a concentrated purge 12 required to remove the solids generated by the reaction. The residue from the vacuum distillation 11 is returned in the reaction.

By means of line 13, equipped with a flow regulator (not shown in the Figure), reaction liquid is collected from the bottom of the phase separator 6, in addition to the flow 8 intended, as mentioned above, for the treatment outside of the reactor, to be fed at the bottom of the reactor 1 , at a unit flow rate Q (cm³/s per cm² of the reaction section, therefore with superficial velocity cm/s). The liquid flow Q of the line 13 may be naturally generated when the phase separator 6 is positioned at a height corresponding to that of the head of the reactor, as a result of the piezometric effect resulting from the greater density of the separator bottom liquid 6, being degassed, compared to the density of the gas-liquid system of the reactor 1. Conversely, a pump (not shown in the Figure) may be used.

The feeding 13 of the reaction liquid, originating from the phase separator 6, with a unit flow rate Q, together with the flows 11 , 2, and in a negligible manner 3, generates an overall flow of liquid in output from the head of the reactor 1 with a unit flow rate L (cm³/s per cm² of the reaction section, i.e., cm/s), equal to Q + 1.5 F, the value of which corresponds to the superficial velocity value uL (cm/s) of the liquid in the reactor 1.

In concomitance with and in relation to the feeding 13 of the reaction liquid originating from the bottom of the phase separator 6, gas including hydrogen is introduced by means of the grid 4 at a superficial velocity uG, such that uG > 10 uL, so as to obtain the replacement of the bubbling liquid by a mixture of gas and reaction liquid in the foam state, under foaming fluid dynamic regime, over the grid 4.

In such an embodiment of the invention, in which the bubbling liquid is replaced by a mixture of gas and reaction liquid in the foam state, under foaming fluid dynamic regime, the gas, not being subject to the limitations of superficial velocity encountered when operating under bubbling regime, may be introduced at the grid 4 at a superficial velocity adapted to evaporate also the totality of the high-boiling conversion products generated by the conversion of the heavy oil 2. The unit volumetric fraction of liquid present in the two-phase gaseous flow 5, uL/(uG+uL), from which uL/uG may be obtained, is detectable, once the density of the reaction liquid and the density of the gas at temperature and pressure conditions of the reactor are known, from the measurement of the density of the two-phase gaseous flow 5 itself. The condition uG/uL > 10, i.e., uL/uG < 0.1 is therefore instrumentally verifiable by means of the value of the two-phase gaseous flow density read by a densimeter 15 along the line 5.

In an alternative embodiment, the positioning of a second grid 14, above a first grid 4, at a distance h from the exit mouth at the upper vault of the reactor 1 al- lows a second introduction of gas including hydrogen at a superficial velocity UG2. By means of said first grid 4, gas Gi is introduced at a superficial velocity within the maximum bubbling regime velocity threshold. Such an embodiment allows to operate the reactor simultaneously under two distinct fluid dynamic regimes: a bubbling regime between the grid 4 and the grid 14, and a foaming re- gime above the grid 14, where the mixture of gas and reaction liquid is in the foam state, under foaming fluid dynamic regime. The gas G2, not being subject to the superficial velocity limitations encountered when operating under a bubbling regime, may be introduced by means of the grid 14, with superficial velocities which are suitable for evaporating even the totality of the high-boiling con- version products generated. The gas G2 may also be introduced at a high temperature to raise the temperature of the mixture of gas and reaction liquid in the foam state, under a foaming regime, above the grid 14.

Preferably, to limit the dehydrogenation which occurs in the liquid at the bottom of the phase separator 6, resulting in dehydrogenation and coke formation, a cooling hydrocarbon with a boiling point ranging from 50°C to 300°C is introduced into the separator 6 by means of line 16.

Based on the description given for a preferred embodiment, it is obvious that changes may be introduced by those skilled in the art without thus departing from the scope of the invention as defined by the following claims.

Claims

C L A I M S

1. A method for the hydroconversion of heavy oils comprising the following steps:

a) introducing into a column reactor (1 ), at a lower portion thereof:

• heavy oil;

• a precursor of a hydroconversion catalyst of the dispersed type,

furthermore, introducing into said reactor (1 ), at least at said lower portion:

• gas including hydrogen,

generating in said reactor (1 ) a two-phase mixture comprising reaction liquid and a gaseous phase,

said reactor (1 ) comprising a cylindrical portion interposed between said lower portion and a head vault thereof;

said reactor (1 ) operating at a temperature from 380°C to 440°C, and at a pressure from 10 MPa to 30 MPa;

b) extracting from said reactor (1), at said head vault, a two-phase gaseous flow originating from said two-phase mixture;

c) introducing said two-phase gaseous flow into a phase separator (6) so as to separate it into a liquid phase and a gaseous phase;

d) subjecting a first fraction of said liquid phase obtained in said separator (6) to flash and distillation, with an under vacuum final distillation stage;

e) introducing into said reactor (1), at said lower portion, a residue of said under vacuum distillation, deprived of a fraction thereof to remove solids generated in said reactor (1 ),

said method being characterized in that it includes the following further step: f) introducing into said reactor (1 ) a second fraction of said liquid phase obtained in said separator (6),

in step a), said gas being introduced into said reactor (1 ) at a superficial velocity uG higher than the maximum superficial velocity threshold usable in a bubbling fluid dynamic regime, said superficial velocity uG referring to the inner section of said reactor (1), at the junction between said cylindrical portion and said head vault, said threshold being 5.5 cm/s,

said superficial velocity uG being, furthermore, greater than ten times the super- ficial velocity uL of said reaction liquid exiting from said reactor (1 ) at said head vault, said superficial velocity uL being referred to the inner section of said reactor (1 ), at the junction between said cylindrical portion and said head vault, so that said two-phase mixture is in the foam state, under foaming fluid dynamic regime, in at least part of said reactor (1 ).

2. A hydroconversion method according to claim 1 , characterized in that, in step a), said gas is introduced into said reactor (1):

• both at said lower portion, at a superficial velocity uGi not greater than said threshold,

• and at said cylindrical portion, at a superficial velocity uG2,

said superficial velocities uGi and 11G2 referring to the inner section of said reactor (1 ) at the junction between said cylindrical portion and said head vault, the sum of uGi with UG2 corresponding to said superficial velocity uG, so that said two-phase mixture is in the foam state, under foaming fluid dynamic regime, above the introduction of said gas at said cylindrical portion.

3. A hydroconversion method according to claim 1 or 2, characterized in that, in step a), said superficial velocity uG is greater than 15 cm/s.

4. A hydroconversion method according to claim 3, characterized in that, in step a), said superficial velocity uG is greater than 30 cm/s.

5. A hydroconversion method according to one of the preceding claims, characterized in that, in step f), said second liquid phase fraction is introduced into said reactor (1 ) at a unit flow rate Q such that the unit volumetric fraction of liquid phase in said two-phase mixture in the foam state, under foaming fluid dynamic regime, is not lower than 0.2.

6. A hydroconversion method according to claim 5, characterized in that, in step f), said unit flow rate Q at which said second liquid phase fraction is introduced into said reactor (1 ) is such that the unit volumetric fraction of liquid phase in said two-phase mixture in the foam state, under foaming fluid dynamic regime, is not lower than 0.4.

7. A hydroconversion method according to claim 2, characterized in that, in step a), said gas is introduced at said cylindrical portion at a temperature such as to heat said two-phase mixture in the foam state, under foaming fluid dynamic regime, up to a temperature from 440°C to 480°C.

8. A hydroconversion method according to one of the preceding claims, characterized in that, in step c), said liquid phase, in said separator (6), is cooled down by introducing into said separator (6) a hydrocarbon having a boiling point from 50°C to 350°C.

9. A hydroconversion method according to claim 8, characterized in that, in step c), said liquid phase, in said separator (6), is cooled down to a temperature not higher than 400°C.

10. A system for the hydroconversion of heavy oils, in a single reaction and recycling step, comprising:

• a column reactor (1 ) comprising a cylindrical portion interposed between a lower portion and a head vault thereof;

• first means (2, 3) for the introduction into said reactor (1 ), at said lower portion, of heavy oil and of a precursor of a hydroconversion catalyst of the dispersed type;

• second means (4) for the introduction into said reactor (1 ), at least at said lower portion, of gas including hydrogen,

a two-phase mixture comprising reaction liquid and a gaseous phase generating in said reactor (1);

• a first line (5) for the extraction, from said reactor (1), at said head vault, of a two-phase gaseous flow originating from said two-phase mixture,

and for the introduction of said two-phase gaseous flow into a phase separator (6) adapted to separate it into a liquid phase and a gaseous phase;

• a second line (8) for the withdrawing of said liquid phase from said separator (6), and for the introduction thereof into a flash and distillation subsystem (9) with an under vacuum final distillation stage;

• a third line (11) for the withdrawing, from said subsystem (9), of a residue of said under vacuum final distillation stage, and for the introduction thereof into said reactor (1), at said lower portion;

• means (12) for the withdrawing, from said third line (11), of a fraction of said residue to remove solids therefrom generated in said reactor (1 ),

said system being characterized in that it further comprises: • a fourth line (13) for the withdrawing, from said second line (8), of a fraction of said liquid phase, and for the introduction of said liquid phase fraction into said reactor (1 ).

11. A hydroconversion system according to claim 10, characterized in that it comprises third means (14) for the introduction of said gas into said reactor (1 ) at said cylindrical portion.

12. A hydroconversion system according to claim 10 or 11 , characterized in that it comprises fourth means (16) for the introduction, into said separator (6), of a hydrocarbon having a boiling point from 50°C to 350°C so as to cool said liquid phase, in said separator (6), down to a temperature not higher than 400°C.