Scope of the invention

The present invention relates to hydroconversion systems for heavy oils using dispersed catalyst reactors (i.e. "slurry bubble column"), including molybdenum, in a single reaction stage with vacuum recycling. More specifically, the present invention relates to a system of the above type in which the conventional extraction of the reaction liquid via biphasic effluent (single carrier with which the conversion products are usually conveyed outside the reactor) is accompanied by a second extraction of reaction liquid, so as to remedy the low extraction capacity occurring with this type of reactors (as a result of the low surface velocity at which hydrogen can be introduced at the reactor base). The present invention further relates to the hydroconversion process steps outside the reactor, related to recycling the unconverted charge and the catalyst in the reaction, so as to counter the formation of coke and the deactivation of the catalyst which occur in the above steps.

The invention also relates to a hydroconversion method for heavy oils which can be implemented by the above system.

Description of the prior art

Heavy oils (such as crude oil, bitumen, oil from tar sands, shale oils and their residues of atmospheric distillation, vacuum distillation and thermal visbreaking) contain hydrocarbons in variable percentage having 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) are a heavy fraction of said oils which is not totally distillable. If subjected to evaporation, said hydrocarbons in fact produce a quantity of carbon residue (expressed as % CCR, namely Conradson Carbon Residue - ASTM D189) which is the 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 capable of reaching 50% of their weight. Such a carbon residue is reduced to 20% for hydrogen contents of around 10% by weight and zeroes when the hydrogen content in the hydrocarbon is around 12% by weight.

In order to transform heavy oils into lighter products which are more valuable to the market, heavy oils are subjected to a temperature treatment with hydrogen and suitable catalysts by means of which the above heavy fraction (also known as "carbonaceous fraction") is converted to into distillable hydrocarbons (i.e. practically free from the carbon residue). Said treatment is also known as "hydroconversion". The catalysts used in such treatments are generally defined "hydrogenation" catalysts or "hydroconversion catalysts". The hydroconversion treatment is aimed at obtaining products free from carbon residue, which can therefore be fed to subsequent hydrocracking and hydrotreating treatments by which said products achieve quality specifications required by the market, or may be used for other refining processes. Hydrocracking and hydrotreating technologies are well tested and available on the market. Therefore, further details are not provided.

Incidentally, if a fraction of the heavy oil consists of hydrocarbons having a boiling point no higher than 540 °C (and therefore constituting a fraction of the heavy oil already free from carbon residue), it may be advantageous to subject only the fraction of the heavy oil having a boiling point higher than 540 °C to hydroconversion. Said heavy fraction is obtainable, by way of example, by preliminarily subjecting the heavy oil to vacuum distillation. Incidentally, in the case of heavy oils including a large amount of carbon residue, the vacuum distillate boiling point may have to be lowered to 520 °C and more in order to be free from carbon residue.

The hydroconversion treatment can be carried out in cylindrical pressure containers (called "reactors") with distribution of hydrogen at the base thereof, where the heavy oil (i.e. the "charge") to be converted is also introduced. Hydrogen and the charge to be converted come into contact with each other in the presence of a hydrogenation catalyst (usually comprising molybdenum) dispersed in the reaction liquid ("slurry catalyst") or deposited on a solid support ("supported catalyst") structured into small cylinders or microspheres (2 - 3 mm in diameter) and consisting of silica and/or alumina or equivalent material. The catalyst deposited on a solid support may be referred to in the remainder of the present description also by the phrase "supported catalyst".

If the hydrogenation catalyst is dispersed in the reaction liquid, the reactor in which the hydroconversion treatment is carried out is referred to as "dispersed catalyst reactor" (or "slurry bubble column reactor"). If the hydrogenation catalyst is deposited on a solid support, the reactor in which the hydroconversion treatment is carried out is referred to as "ebullated catalytic bed reactor". Incidentally, the most frequently used reactors on a commercial scale are the ebullated catalytic bed reactors. Recently, also the dispersed catalyst reactor type has been brought to a commercial scale.

In dispersed catalyst reactors, the catalyst can be introduced into the reactor in various manners, such as using an oil-soluble precursor (i.e. a metal compound capable of generating the active species when it comes into contact with the charge). Hydrogen is usually introduced through a grid with nozzles placed at the base of the reactor. This allows the catalyst to remain evenly and stably dispersed in the reaction liquid from which it is separable, by way of example, by filtration, centrifugation or by decanter.

In ebullated catalytic bed reactors, hydrogen and the charge to be converted are typically introduced in the reactor through a perforated plate placed at the bottom thereof. The solid elements on which the catalyst (usually transition metals) is deposited are kept suspended in the reaction liquid by a circulation of the same, from bottom to top, obtained by means of a pump (ebullating pump) either internal or external to the reactor. Inside the reactor, in the upper reaction zone, a funnel with a descending pipe (downcomer) may be provided that collects the reaction liquid at the top of the reactor and conveys it down, in suction to the circulation pump so that it is returned to the upper reaction zone going up inside the reactor. By suitably adjusting the number of revolutions of the pump, the supported catalyst remains suspended in the reaction liquid and at the same time confined inside the reactor, so as to prevent leaks thereof as a result of an exit from the reactor with the reaction liquid.

In both types of reactors mentioned above, hydrogen is generally introduced at the base of the reactor. The introduction of hydrogen generates a set of bubbles, which going up in the reaction liquid, facilitates the mixing thereof, thereby ensuring high coefficients of mass and heat transfer both in the axial and in the radial direction of the reactor, even in the absence of stirrer or mechanical mixing systems.

By the effect of reaction conditions (temperature and pressure) and of the catalyst present, a gas-liquid effluent containing conversion products is produced at the head of the reactor (whatever the type of catalyst used: dispersed or supported). The solids simultaneously generated by the reaction (including the sulfides of the metals present in the charge, coke, insoluble asphaltene resins and solids due to the catalyst) are finely dispersed in the liquid phase. The effluent generated at the head of the reactor is also referred to as "biphasic flow" or "biphasic effluent". As the conversion products having the lowest boiling point (usually lower than 300 °C) at atmospheric pressure do not significantly accumulate in the reaction liquid, they are mainly found in the gaseous phase of the biphasic effluent. In contrast, products having the highest boiling point (usually between 300 and 540 °C) at atmospheric pressure mainly remain liquid constituting, together with the non-converted charge fraction, the reaction liquid exiting from the reactor as liquid component of the biphasic effluent.

The biphasic effluent is conveyed to a phase separator, which is a vertical cylindrical vessel which can operate at the same pressure as the reactor and which allows to separate the gaseous phase of the biphasic effluent from the liquid phase (with the solids generated in the reaction finely dispersed in the latter) of the same. More precisely, the gaseous phase including the low boiling conversion products and the residual hydrogen exits from the head of the separator. The low boiling conversion products (also called "volatile") are recovered by cooling and condensation in one or more steps while the hydrogen is conveyed to the purification section to be then reused. At the bottom of the separator, due to density, collects the reaction liquid comprising the non-converted charge fraction, the high-boiling conversion products (also known as "high-boiling"), the catalyst and the solids generated in the reaction. Hyd reconversion systems employing ebullated catalytic bed reactors usually have multiple reaction stages. In particular, the reaction liquid collected at the bottom of a first phase separator which treats the biphasic flow from the first reactor is fed to a second reactor, in series to the first one, which produces further conversion products which are collected as gaseous phase at the head of a second phase separator, at the bottom of which the corresponding reaction liquid along with generated solids collects. A third reactor, in series with the first two, may be provided. The hydroconversion systems of this type are characterized by an extraction of the conversion products only from the gaseous phase (at the head of the separator) since the liquid phase (which, along with the solids, collects at the bottom of the separator) is cascade feeding of the next reactor. The conversion degree (given by the sum of the weight flow rates of the conversion products extracted as gaseous phase of each reaction stage, divided by the charge flow rate fed to the first reactor) progresses in the different reaction stages in cascade, each of which, however, results in an increase in the insoluble n-pentane (carbonaceous) moiety present in the reaction liquid. Above a predetermined insoluble n-pentane threshold, the homogeneity and stability of the reaction liquid is lost due to the precipitation of asphaltene and other carbon material in the final reaction stages. In order to prevent such a precipitation, the conversion degree is limited to values which hardly reach 75%. As a result, a significant fraction of the fed charge leaves the plant as non-converted residue the market placement of which is increasingly difficult, also as a result of stricter rules aimed at protecting the environment. Unlike the hydroconversion systems employing ebullated catalytic bed reactor with supported catalyst, the hydroconversion systems employing dispersed catalyst reactors are single reaction stage with recycling since the extraction of conversion products is also carried out from the liquid phase which collects at the bottom of the separator. In particular, said liquid phase, after depressuriza- tion with consequent flash in one or more stages, is conveyed to an atmospheric distillation and vacuum distillation treatment with recycling of the vacuum distillation residue to the reactor. This treatment "outside the reactor" (also called "down stream") of the reaction liquid allows both to recover, at least partially, the catalyst which at the end of the above treatments is dispersed in the vacuum bottom liquid along with the solids generated in the reaction, and extract the high-boiling conversion products. By operating in a single reaction stage with recycling, the limits to the conversion degree due to the precipitation of asphal- tenes which are found in multistage hydroconversion systems, which generally use ebullated catalytic bed reactor with supported catalyst are eliminated, thereby allowing the achievement of 90% or higher conversion degrees.

Hydroconversion systems using dispersed catalyst reactors are however characterized by a biphasic effluent at the head of the reactor (the only carrier which conveys the conversion products outside the reactor) which has a significantly lower flow rate than the biphasic effluent flow rate at the head of supported catalyst reactors. This is a consequence of a lower surface velocity by which hydrogen can be introduced at the base of dispersed catalyst reactors. The above causes a poor extraction capacity of the conversion products which results in an accumulation thereof in the reaction liquid, particularly high-boiling ones. The accumulation of conversion products of maltene - thus low reactivity - origin in the reaction adversely affects the capacity of the reactor, as it subtracts relevant reaction volume fractions to the charge to be converted. The negative effect of accumulation of high-boiling liquids on the hydroconversion capacity is especially relevant in very high reactors, where the accumulation continues with the height of the reactor, thereby reducing the hydroconversion unit capacity of the system (defined as m ³ of charge converted in an hour per m ³ reaction volume). A so-reduced unit capacity limits the economic convenience to use very high reactors which can currently be achieved due to the most advanced construction techniques. In order to overcome these limits and achieve (and possibly exceed) the unit capacity of a supported catalyst systems, it is therefore necessary to provide dispersed catalyst reactors with a different and enhanced conversion product extraction system.

In addition to that, in single reaction stage and vacuum recycling hydroconver- sion systems, the process steps outside the reactor, necessary to return the non-converted reaction fraction and the catalyst back to the reaction, as known to date, are not adequate to prevent the formation of coke and the deactivation of the catalyst therewith, factors which contribute both to an increase in the catalyst consumption.

The hydroconversion systems using dispersed catalyst reactors (including molybdenum) in a single reaction stage and recycling, even if preferable, in principle, due to the highest degree of conversion attainable, still have a limited spreading at an industrial level as a result of systems and methods of using molybdenum which are inadequate to prevent the deactivation thereof and contain consumption thereof within economically acceptable limits. In such single reaction stage hydroconversion systems, the recycling of the vacuum distillation residual liquid to the reactor in fact returns the catalyst to the reaction but it is a catalyst which, as a result of the coking it undergoes (i.e. as a result of the deposition of coke on the surface of the catalyst), has a reduced catalytic activity to an extent dependent on the amount of coke produced. For fractional amounts of coke with respect to the catalyst, the degree of deactivation is of course limited. When the weight of the coke formed is comparable to the amount of catalyst present in the reaction, the catalyst deactivation results in a reduction of catalytic activity compensated with higher, sometimes significantly higher, catalyst amounts. If the coke formed is greatly in excess, catalyst recycled in reaction through the vacuum distillation residue is substantially free of catalytic activity (as described in US 5,294,329). The catalytic activity of a molybdenum-based catalyst, upon the first use (i.e. when operating without recycling), is such as to allow the achievement of the plateau to the catalytic activity at molybdenum concentrations in the reaction liquid within 500-1000 ppm (as described in US 4,226,742). In contrast, by operating with recycling, the achievement of the plateau requires much higher molybdenum concentrations depending, as mentioned above, on the degree of coking.

In order to try to solve this problem, US patent 5,294,329, following the assump- tion that keeping the catalyst in a reducing atmosphere (that is, in hydrogen pressure) is sufficient to prevent the deactivation thereof, describes a hydroconversion method in which a part of the catalyst is recycled directly from the bottom of the phase separator (in hydrogen pressure), after concentration in a settler. Such a recycling method is not, however, entirely adequate to preserve the catalyst from coking (and therefore prevent the deactivation thereof), especially the portion of catalyst which remains in the reaction liquid which is subjected to the treatments (in the absence of hydrogen) needed for recovering high-boiling conversion products. In addition to the above, the recycling of the catalyst in the reaction, through the vacuum distillation residue, takes place only to a partial extent, since a part of the catalyst is removed by the bleeding necessary to stabilize the accumulation of solids generated in the reaction: metal sulfides contained in the charge and mainly coke. Therefore, the amount of catalyst lost by the hydroconversion system as a result of bleeding increases correspondingly to the formation of coke. In order to limit the formation of coke and thus improve the recovery of molybdenum (economically relevant factor), US patent 8,617,386, with reference to a generic hydroconversion system, describes the use of supported molybdenum which is separated from the vacuum bottom residue, subjected to treatment, and then re-fed into the reaction. While molybdenum consumptions are improved, they still remain high.

US patent 7,578,928, among other aspects, deals with the problem of coke formation downstream of the reactor, specifically in the phase separator, with the aim to oppose fouling and coke deposition in the separator. The actions indicated are, however, totally inadequate to prevent coking and the deactivation of the catalyst dispersed in the reaction liquid.

In the light of the above, how to contain the formation of coke and effectively preserve the catalyst from the coking phenomenon, particularly during treatment "outside reactor", is still an unsolved problem. Objects of the invention

It is a primary object of the present invention to overcome the drawbacks of low unit capacity by providing a hydroconversion system and method for heavy oils by means of a dispersed catalyst reactor, including molybdenum, in a single re- action stage with vacuum recycling, wherein the extraction the reaction liquid from the reactor is enhanced and adaptable, so as to overcome the limits of unit capacity and dependency of the unit capacity on the height of the reactor, these drawbacks being intrinsic in traditional hydroconversion systems.

It is a second object of the present invention to solve, at least partially, the prob- lems of coking and deactivation of the catalyst described above, associated with the treatment of the reaction liquid outside of the reactor, by providing a hydro- conversion system and method for heavy oils by a dispersed catalyst reactor in a single reaction stage with recycling which allow:

• to preserve, almost totally, the recycled catalyst from the coking phenomenon and subsequent deactivation;

• to limit the formation of coke to make the conversion degree close to completeness;

• to limit the consumption of dispersed catalyst including molybdenum.

Summary of the invention

The present invention relates to a hydroconversion system for heavy oils in a single reaction stage, comprising:

• a dispersed catalyst reactor;

• a first line for feeding the reactor with heavy oil or preferably, with a fraction of heavy oil having a boiling point, at atmospheric pressure, higher than 540 °C;

• a second line for feeding the reactor with a first gas including hydrogen;

• a third line for feeding the reactor with a hydroconversion catalyst,

the reactor containing a mixture called "reaction liquid" and including:

- a liquid phase comprising:

> a non-converted fraction of said heavy oil;

> hydroconversion products in the liquid state;

- a gaseous phase comprising: > an amount of said first unreacted gas;

> hydroconversion products in the gaseous state and in the vapor state;

- solids generated in the reaction which are dispersed in the liquid phase;

- said hydroconversion catalyst dispersed in the liquid phase,

the reaction liquid originating an effluent called "biphasic" in case of exit of the same from the head of the reactor;

• a separator suitable for separating the liquid phase, together with the solids and the catalyst dispersed therein, from the gaseous phase of the biphasic effluent;

· a fourth line for withdrawing the biphasic effluent from the upper part of the reactor and introducing the same into the separator;

• a first depressurization, preferably at atmospheric value, and subsequent distillation stage (also known as atmospheric flash-distillation stage), said first stage including at least one depressurization tank and a distillation column; · a fifth line for withdrawing, from the separator, said liquid phase, along with the solids and the catalyst dispersed therein, and introducing the same in said first stage;

• a second vacuum distillation concentration stage, said second stage including a vacuum column;

· a sixth line for withdrawing from said first stage a first residue of the same and for introducing said first residue in said second stage;

• a seventh line for withdrawing from said second stage a second residue of the same and for introducing said second residue in the reactor;

• an eighth line for withdrawing, at least partially, from said seventh line, a flow in which the solids generated in the reaction are dispersed,

wherein, according to the invention, the hydroconversion system further comprises:

• means for degassing the reaction liquid at least partially accommodated within the reactor at the upper part of the same;

· a ninth line for withdrawing the degassed reaction liquid (comprising the liquid phase along with the solids and the catalyst dispersed therein) from the degassing means and introducing the same in the separator at a lower part of the same.

Incidentally, the above atmospheric flash-distillation and vacuum distillation concentration stages are known. Therefore, further details are not provided. As will be explained hereinafter in the present description, due to the concurrent withdrawal of both the biphasic effluent and the degassed reaction liquid (the rate whereof can be modulated depending on the height of the reactor) from the reactor and introduction in the separator, the hydroconversion system of the invention has a higher unit capacity than the known systems (where only the bi- phasic effluent is withdrawn and introduced in the separator), a unit capacity which is also kept as the height of the reactor increases.

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

According to one aspect of the invention, the degassing means comprise a de- gasser-conveyor at least partially shaped as a conical cylinder including a cylindrical wall above a conical wall, the outer diameter of said cylindrical wall preferably being at least 0.71 fold, and even more preferably at least 0.82 fold the inner diameter of the reactor.

According to another aspect of the invention, the hydroconversion system of the invention further comprises:

• means for regulating the flow of the degassed reaction liquid along ninth line. As will be better illustrated hereinafter in the present description, the regulation means preferably comprise:

- a circulation pump placed at the ninth line;

- a differential pressure gauge for detecting a pressure difference between the outlet of a descending pipe of the degasser-conveyor and the reactor head.

According to another aspect of the invention, the hydroconversion catalyst comprises molybdenum,

the hydroconversion system further comprising:

• a tenth line wherein said first and seventh line flow, the heavy oil and said second residue, the latter deprived of the flow in which the solids generated in the reaction are dispersed, being able to be introduced in the reactor through said tenth line;

means for heat exchange between:

- the degassed reaction liquid exiting the reactor through said ninth line and

- the heavy oil and said second residue in input into the reactor through said tenth line, and/or a cooling fluid;

an eleventh line for introducing a second gas including hydrogen in the separator at the lower part of the same;

first means for distributing said second gas in a cross section of the separator, the first distribution means being connected to said eleventh line and being accommodated, at least partially, within the separator below a zone in which the degassed reaction liquid can be introduced in the separator;

the heat exchange means preferably comprise a heat exchanger between said ninth line and said tenth line. By the effect of heat exchange, the heavy oil (along with the second residue deprived of the solids generated in the reaction) is heated before being introduced in the reactor, and the liquid phase exiting the degasser-extractor (along with the solids and the catalyst dispersed therein) is cooled before being introduced in the separator. This causers a cooling of the liquid phase at the bottom of the separator;

a twelfth line for introducing a third gas including hydrogen in the depressuri- zation tank (or in the depressurization tanks) at a lower part of the same; second means for distributing said third gas in a cross section of the depressurization tank (or of the depressurization tanks), the second distribution means being connected to said twelfth line;

a thirteenth line for introducing a fourth gas including hydrogen in the distillation column at a lower part of the same;

third means for distributing said fourth gas in a transverse section of the distillation column, the third distribution means being connected to said thirteenth line;

a fourteenth line for introducing a fifth gas including hydrogen in said tenth line, said fifth gas being able to be introduced from said fourteenth line in said tenth line upstream of said heat exchange means;

• means for dispersing said fifth gas in said tenth line.

According to another aspect of the invention, the hydroconversion system further comprises:

· a fifteenth line for feeding the vacuum column with a compound capable of releasing hydrogen when it comes into contact with molybdenite (preferably a mono-cycloalkane).

As will be illustrated hereinafter in the present description, in the system of the invention the continuous introduction and the even distribution of hydrogen in the cooled reaction liquid at the bottom of the separator, in the reaction liquid in the various process steps outside the reactor and in the liquid to be converted (heavy oil and vacuum recycling), preventing the molybdenum dihydrosulfide (present in the reaction) from turning back to molybdenite (which generates coke), prevents the coking phenomenon. The undesired presence of non- hydrogenated molybdenite in the liquid to be converted is countered by the introduction and distribution of hydrogen also in the surge tanks which may be present in the various process steps, particularly outside the reactor. The solids to be disposed of are consequently considerably reduced, as well as the progressive deactivation of the catalyst.

The invention also relates to a hydroconversion method for heavy oils in a single reaction stage (which can be implemented, by way of example, by the system of the invention) comprising the following steps:

a) feeding a dispersed catalyst reactor with:

- heavy oil or preferably a fraction of heavy oil having a boiling point, at at- mospheric pressure, higher than 540 °C;

- a first gas including hydrogen;

- a hydroconversion catalyst,

a mixture called "reaction liquid" being generated inside the reactor and including:

- a liquid phase comprising:

a non-converted fraction of said heavy oil;

> hydroconversion products in the liquid state; - a gaseous phase comprising:

> an amount of said first unreacted gas;

> hyd reconversion products in the gaseous state and in the vapor state;

- solids generated in the reaction which are dispersed in the liquid phase; - said hydroconversion catalyst dispersed in the liquid phase,

the reaction liquid originating an effluent called "biphasic" in case of exit of the same from the head of the reactor;

b) withdrawing the biphasic effluent from the upper part of the reactor and separating, by means of a separator, the liquid phase along with the solids and the catalyst dispersed therein from the gaseous phase of the biphasic effluent;

c) subjecting the liquid phase, along with the solids and the catalyst dispersed therein, to depressurization, preferably at atmospheric value, and subsequent distillation by at least one depressurization tank and a distillation column, re- spectively;

d) subjecting a first residue of the distillation column to vacuum distillation concentration by means of a vacuum column;

e) introducing a second residue of the vacuum distillation concentration in the reactor,

solids generated in the reaction being withdrawn from said second residue before the same is introduced into the reactor,

where, according to the invention:

• the reactor operates at a temperature of between 380 °C and 430 °C and at a pressure of between 10 MPa and 30 Pa;

in addition to that:

• in step b), an amount of reaction liquid being degassed inside the reactor, and the degassed reaction liquid being introduced in the separator, in step c) also said degassed reaction liquid collected on the bottom of the separator together with said liquid phase being subjected to depressurization;

· 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, wherein Vs is the space velocity with which, in step a), said heavy oil is introduced into the reactor and H is the height of the reactor. As will be specified in the detailed description of the invention, said coefficient 0.5 instead of being equal to 0.5 can preferably be equal to 1.0 and even more preferably equal to 2.0.

By "unit flow rate of said second residue" it is meant the ratio of the rate at which said second residue is introduced into the reactor, expressed in m ³ per hour, to the cross sectional area of the reactor expressed in m ² . The unit flow rate is therefore expressed in meters per hour.

By "spatial velocity of said heavy oil" it is meant the ratio of the rate at which said heavy oil is introduced into the reactor, expressed in m ³ per hour, to the reactor volume expressed in m ³ . The spatial velocity is therefore expressed in hours ¹ .

The height H of the reactor is expressed in m.

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

According to one aspect of the invention of method, in step b), the degassing of the reaction liquid taking place by means of a degasser-conveyor shaped, at least partially, as a conical cylinder including a cylindrical wall above a conical wall, the degasser-conveyor being connected to a descending pipe for introducing the degassed reaction liquid in the separator, the introduction of the de- gassed reaction liquid in the separator being regulated so that the level of the degassed reaction liquid in the descending pipe is kept within the pipe itself, preferably in the lower half of the latter.

According to another aspect of the invention, the hydroconversion catalyst comprises molybdenum and, in addition to that, in step b):

· the degassed reaction liquid is cooled, prior to an introduction of the same into the separator, so that the temperature of the liquid phase at the bottom of the separator, at the same pressure at which the reactor operates, is lower by at least 20 °C than the temperature at which the reactor operates and is of between 340 °C and 410 °C,

and

• a second gas including hydrogen is introduced into the separator at a lower part of the same at a surface velocity of between 0.01 cm/s and 1 cm/s. By "surface velocity of the second gas" it is meant the ratio of the flow rate at which, in step b), the second gas is introduced into the separator, measured at the same temperature and same pressure as the latter, expressed in cm ³ per second, to the cross sectional area of the separator expressed in cm ² . The surface velocity of the second gas is therefore expressed in cm per second;

in step c), a third gas including hydrogen being introduced in the depressuriza- tion tank (or in the depressurization tanks) at a lower part of the same, said third gas being also distributed in a cross section of the depressurization tank, in step c), a fourth gas including hydrogen being introduced in the distillation column at a lower part of the same, said fourth gas being also distributed in a cross section of the distillation column,

in step e), a fifth gas including hydrogen being introduced and dispersed in the heavy oil and in the second residue before the heavy oil and the second residue are fed to the reactor.

According to one aspect of the invention of method, in step b), the degassed reaction liquid is cooled before introducing the same in the separator, so that the temperature of the liquid phase at the bottom of the separator, at the same pressure at which the reactor operates, decreases so as to be between 340 °C and 370 °C.

According to one aspect of the invention of method, in step d), a compound capable of releasing hydrogen when it comes into contact with molybdenite (preferably a mono-cycloalkane) is introduced in a liquid present in the vacuum column.

According to another aspect of the invention, the accumulation factor is not less than 30.

By "accumulation factor" it is meant the ratio of the flow rate at which, in step a), the heavy oil is introduced in the reactor, expressed in m ³ per hour, to the flow rate at which, in step e), there is a withdrawal from said second residue, ex- pressed in m ³ per hour. The accumulation factor therefore is dimensionless.

According to another aspect of the invention, the accumulation factor preferably is not lower than 40, and even more preferably not lower than 50. Brief description of the drawings

Further objects and advantages of the present invention will become apparent from the following detailed description of an exemplary embodiment thereof and from the accompanying drawings, which are merely illustrative and non-limiting, in which:

- figure 1 schematically shows a hyd reconversion system according to the present invention;

- figure 2 schematically shows a variant of the system in figure 1.

Detailed description of some preferred embodiments of the invention

Hereinafter in the present description, a figure may also be illustrated with reference to elements not expressly indicated in that figure but in other figures. The scale and proportions of the various elements depicted do not necessarily correspond to the actual ones.

Figure 1 shows a hydroconversion system for heavy oils in a single reaction stage with recycling, comprising a dispersed catalyst reactor 4. Reactor 4 is fed at the base with the liquid to be converted through a line 26 in which a line 3 for feeding reactor 4 with heavy oil (or, preferably, with a fraction of heavy oil having a boiling point, at atmospheric pressure, higher than 540 °C), and a line 20 for recycling a "vacuum bottom residue" to reactor 4, i.e., as will be better illustrated hereinafter in the present description, the residue of a vacuum distillation concentration stage 25 for extracting high-boiling conversion products. Said liquid to be converted therefore comprises heavy oil and the vacuum bottom residue.

Reactor 4 is also fed with hydrogen, or with a gas including hydrogen, through a line 1. Hydrogen (or the gas including the same) is introduced at the base of reactor 4 preferably through a nozzle distributor, such as a grid, or a perforated plate. In the latter case, it may be premixed with the charge to be converted (not shown in the figure). Hydrogen is possibly introduced into reactor 4 also at a second inlet placed at a higher height.

Reactor 4 is further fed, at the bottom of the same, with a hydroconversion catalyst through a line 2, from which it quickly disperses in the reaction liquid.

The catalyst may be introduced into reactor 4 dispersed in water or using an oil- soluble precursor, i.e. a composite of one or more transition metals, among which molybdenum, capable of generating the catalytically active species when it comes into contact with the heavy oil or with the reaction liquid or the liquid to be converted. Iron and molybdenum based catalysts may also be used. Reactor 4 is fed with the catalyst to compensate for the amount thereof which is removed when discharging the solids.

In addition to the above catalyst including molybdenum, reactor 4 may be fed at the base with a second zeolitic catalyst which, advantageously, improves the removal of nitrogen present in the heavy oil. The zeolitic catalyst may be either fed to reactor 4 through line 2 or through a dedicated line (not shown in the figure).

Reactor 4 preferably operates at a temperature of between 380 °C and 430 °C and at a pressure of between 10 MPa and 30 MPa. In reaction conditions, the hydrogen distributed at the base of reactor 4 generates a set of bubbles which, going up through the reaction liquid (also referred to as "bubbling liquid" for that reason), causes the mixing thereof, thereby ensuring high coefficients of heat and mass transfer, both in the radial direction and in the axial direction. This condition of substantial thermal and material uniformity of the reaction liquid is described as "bubbly regime". The presence and the upward motion of hydrogen bubbles inside reactor 4 further have the fundamental function of ensuring that the reaction liquid is almost constantly and uniformly hydrogen saturated, so that the action of the catalyst is maximally focused on hydrogenation rather than dehydrogenation.

The reaction liquid which is generated inside reactor 4 includes a gaseous phase which comprises a residual amount of hydrogen (i.e., not reacted in reactor 4) and conversion products in the gaseous state and in the vapor state, and a liquid phase which comprises a non-converted fraction of heavy oil and conversion products in the liquid state, mainly with a high boiling point. Solids generated in the reaction (comprising coke, insoluble asphaltene resins and metal sulfides from the heavy oil) and the hydrogenation catalyst reusable in the reaction by recycling the vacuum bottom residue are also finely dispersed in the liquid phase. The reaction liquid exits from the head of reactor 4, through a line 23, thereby generating an effluent referred to as "biphasic". The latter, through line 23, is fed to an external gas-liquid separator 11 operating at the same pressure as reactor 4 and preferably placed at the base of the latter. Incidentally, separator 11 is suitable for separating said liquid phase (together with the solids and the catalyst dispersed therein) from the gaseous phase of the biphasic effluent. The latter is withdrawn at the head of separator 11 and conveyed to a cooling and condensation stage (not shown in the figure) for the recovery of light conversion products and residual hydrogen. The latter, through a dedicated line, is conveyed to a purification stage (also not shown in the figure) to be then recycled to reactor 4.

The liquid phase of the biphasic effluent is collected by density at the bottom of separator 11 together with the solids generated in the reaction. An amount of hydrogenation catalyst at a concentration close to that of reaction is also present in the liquid at the bottom of separator 11. The liquid at the bottom of separator 11 with the solids produced by the reaction and the hydrogenation catalyst suspended therein, is conveyed through a line 14 to one or more depressurization stages 15, preferably at atmospheric pressure, and subsequent distillation (also known as "atmospheric flash-distillation stage"). The liquid at the bottom of separator 11 is conveyed to stage 15 for the recovery of the more volatile conversion products. The residue of stage 15 (namely, the liquid at the base of a distillation column included in stage 15) is conveyed, through a line 24, to a vacuum distillation concentration stage 25 for extracting high-boiling conversion products. The residue of stage 25 (namely, the above "vacuum bottom residue") is recycled to reactor 4 through line 20 to allow the reuse of the hydrogenation catalyst. A flow is branched off line 20, through a line 19, which is used to remove the solids generated in the reaction and accumulated in the reaction liquid from the hydroconversion system.

Incidentally, stages 15 and 25 may serve multiple reactors 4 in parallel.

Compared to known hydroconversion systems for heavy oils using a dispersed catalyst reactor in a single reaction stage with recycling, in the hydroconversion system of invention reactor 4 includes, at an upper zone thereof, a degasser- conveyor 5 shaped, at least partially, as a conical cylinder including a cylindrical wall above a conical wall. The degasser-conveyor 5 degasses a certain amount of reaction liquid and conveys the degassed reaction liquid (comprising said liquid phase together with the solids and catalyst dispersed therein) in a descend- ing pipe 6 at least partially accommodated inside reactor 4. In particular, the reaction liquid from which the gas is released to go up towards the head of reactor 4 falls into the volume delimited by said cylindrical wall. The reaction liquid, having traveled down the above cylindrical wall (which in any case ensures the completion of degassing), is conveyed through the conical end part of the de- gasser-conveyor 5 towards the descending pipe 6. A line 8, communicating with the latter, withdraws the degassed reaction liquid and introduces it into separator 11 , preferably at a lower part of the same. The degassed reaction liquid mixes with the liquid phase of the biphasic effluent which goes down at the bottom of separator 11. Incidentally, under normal operating conditions, the liquid phase obtained in separator 1 from the biphasic effluent has a dispersed solid and catalyst content substantially similar to that of the degassed reaction liquid from the degasser-conveyor 5. High-boiling conversion products, which can be recovered from the concentration stage 25 by vacuum distillation, come both from the liquid phase of the biphasic effluent and from the degassed reaction liquid.

The level of reaction liquid inside reactor 4 is almost at the top edge of the degasser-conveyor 5, the latter serving as an overflow. The degassed reaction liquid collected by the degasser-conveyor 5 goes down and is dynamically positioned inside the descending pipe 6 at such a height that the hydrostatic pres- sure, with respect to the level of liquid in separator 11 , equalizes the load losses of line 8. If the hydrostatic pressure is insufficient to bring the degassed liquid reaction to separator 11 (such as in the case of a positioning of separator 11 at a height), a circulation pump 28 is inserted between the outlet of the descending pipe 6 and line 8, which extracts high flows of degassed reaction liquid to be in- traduced into separator 11. Pump 28 keeps the degassed liquid level within the height of the descending pipe 6, and preferably within the bottom half of the same. The number of revolutions of pump 28 is therefore adjusted according to the level of degassed liquid inside the descending pipe 6. Said level is detectable starting from the measurement of the pressure difference between the lower end of the descending pipe 6 and the head of reactor 4. This pressure difference is measurable, for example, by using a differential pressure gauge (not shown in the figure). In light of the above, pump 28 together with said differential pressure gauge, act as a means for regulating the flow of degassed reaction liquid along line 8 (i.e. as means for regulating the flow of degassed reaction liquid introduced at the base of separator 11 ), corresponding to the liquid generated by reactor 4 in excess with respect to the liquid exiting through the bi- phasic flow, when operating with a vacuum bottom recycling with an unit rate of at least 0.5 x Vs x H.

In a single reaction stage hyd reconversion system, as a result of recycling, the solids generated in the reaction and the catalyst accumulate in the reaction liquid. The continuous bleeding 19 allows stabilizing the level of accumulation of such solids. The ratio of the flow rate at which heavy oil is fed through line 3 to the flow rate of bleed 19 determines the "accumulation factor". For example, such a factor indicates how many times the concentration of catalyst fed to reactor 4, expressed with respect to the fresh charge fed 3, increases in bleed 19 and thus, in a related manner, in the reaction medium. High accumulation fac- tors possible with low flow rates of bleeding, are therefore employed. In particular, only operating with an accumulation factor of at least 30, preferably at least 40, and even more preferably at least 50, a sufficient concentration of catalyst in the reaction is obtained, while operating at low catalyst reintegration through line 2. Operating conditions with a low production of solids allow to increase the value of the accumulation factor.

The complete degassing of the reaction liquid inside reactor 4 is obtainable by operating in conditions of coalescence of the gas suitably induced by a suitable narrowing of the passage section at the cylindrical part of the degasser- conveyor 5. In particular, in order to produce an adequate degree of coales- cence, the surface velocity of the gaseous phase at the passage section delimited by the inner wall of reactor 4 and by the outer cylindrical wall of the degas- ser-conveyor 5 is preferably not lower than 6 cm/s and even more preferably not lower than 8 cm/s. In such conditions, the degassing is facilitated by the higher ascending velocity of the bubbles as a result of coalescence, which greatly increases their average size (switching from "bubbly regime" to "coalesced bubble flow regime"). The degassing capacity limits which are encoun- tered with conventional degassers operating drowned in the bubbling liquid are thus overcome.

More precisely, in order to obtain the most complete degassing of the reaction liquid, the outer diameter of the cylindrical wall of the degasser-conveyor 5 is preferably at least 0.71-fold, and even more preferably at least 0.82-fold, the in- ner diameter of reactor 4. By the effect of the narrowing of the passage section between the inner wall of reactor 4 and the cylindrical outer wall of the degasser-conveyor 5, the velocity of the gaseous phase increases to values which produce the gas coalescence, thereby facilitating the degassing even at high liquid flow rates.

Incidentally, "surface velocity of the gaseous phase at said passage section" means the ratio of the flow rate with which, in step a), the first gas is introduced at the base of the reactor, measured at the same temperature and same pressure of the latter, expressed in cm ³ per second, to the area of said passage section expressed in cm ² . The surface velocity of the gaseous phase is therefore expressed in cm per second.

The flow rate of the degassed reaction liquid obtained through the degasser- conveyor 5, to be used for the extraction of high-boiling conversion products in addition to the liquid phase of the biphasic effluent, is a function of the flow rate with which the residue of vacuum distillation is recirculated to reactor 4 through line 20. Within reactor 4, the above residue 20 is again enriched with high- boiling conversion products, taking the average composition of the reaction liquid. After leaving reactor 4 through line 6, the residue, now as degassed reaction liquid containing high-boiling conversion products, after the transit in separator 11 , makes the high-boiling conversion products available for vacuum ex- traction of treatment 15 and 25. By feeding vacuum residue flow rates to the bottom of reactor 4 (measured in m ³ per hour per m ² of section of reactor 4) at least equal to 0.5 x Vs x H (where Vs is the space velocity with which the charge is fed, expressed as m ³ per hour divided by the volume of reactor 4 in m ³ , and H is the height in meters of reactor 4), preferably at least equal to Vs x H, and even more preferably at least equal to 2.0 x Vs x H, the unit capacity of conversion reaches and exceeds values typical of systems employing support- ed catalyst reactors and is maintained, should the height of reactor 4 increase. More particularly, if the dispersed catalyst is derived from an oil-soluble molybdenum compound and the charge to be converted is a fraction of a heavy oil with boiling point higher than 540 °C obtained as residue of vacuum distillation, assume that, by way of example, reactor 4 has a height equal to 30 meters and is fed, at line 3, with a flow rate of charge to be converted corresponding to 4,62 m ³ /h per m ² of the horizontal section of reactor 4, which corresponds to a spatial velocity Vs of 0.154 h ^~1 . Also assuming that the residue of vacuum distillation is fed to the bottom of reactor 4 at a flow rate of 8.1 m ³ /h per m ² of section of reactor 4, corresponding to 1.75 times Vs x H. By operating according to this hydro- conversion system, the value 0.154 h ¹ of spatial velocity with which the charge is fed is close, as a consequence of the high degree of conversion, to that of the hydroconversion unit capacity of reactor 4 expressed in m ³ /h of charge converted by m ³ of reaction volume. With the same feeding, the higher degree of conversion results in a higher productive output, of about one third. Such a value of the unit capacity is comparable or higher than that of supported catalyst reactors. If, with equal reaction conditions, a higher reactor is used, the use of a vacuum bottom flow rate which increases to 1.75 times Vs x H (with H now numerically larger) allows to keep Vs constant, i.e. to keep the unit capacity even though the height of reactor 4 increases. In contrast, if a higher reactor is used without the possibility of adapting the extraction of high-boiling conversion products (conventional dispersed catalyst reactors), the increase of H would result in a reduction of Vs, i.e. the higher the H, the lower the unit capacity.

The hydroconversion system of the invention therefore allows to overcom the limits of unit capacity and the dependence of the unit capacity on the height of the reactor, intrinsic in slurry phase hydroconversion systems employing conventional methods of product extraction.

Figure 2 shows a hydroconversion system which differs from the system shown in figure 1 in that the degassed reaction liquid is suitably cooled before being introduced into separator 11 through line 8. In particular, the degassed reaction liquid is cooled so that the temperature of the liquid phase which collects at the bottom of separator 11 is preferably lower than the temperature threshold of "thermal cracking". In particular, the degassed reaction liquid is cooled so that the temperature of the liquid phase which collects at the bottom of separator 1 is preferably lower by at least 20 °C than the temperature at which reactor 4 operates and is preferably in the range between 340 °C and 410 °C. Even more preferably, the degassed reaction liquid is cooled so that the temperature of the liquid phase which collects at the bottom of separator 11 is of between 340 °C and 370 °C.

The cooling of the degassed reaction liquid preferably takes place by means of a heat exchanger 9 installed at lines 8 and 26 to allow a heat exchange between the degassed reaction liquid and the heavy oil together with the vacuum distillation residue (deprived of bleed 19 in order to stabilize the accumulation of solids generated in the reaction). Advantageously, the liquid to be converted is thus preheated. Alternatively or in addition to the heavy oil and to the vacuum distillation residue, the heat exchanger 9 is adapted to allow a heat exchange between the degassed reaction liquid and a cooling fluid.

The increase of load losses along line 8, caused by the insertion of the heat exchanger 9, is compensated by the circulation pump 28 inserted between the outlet of the descending pipe 6 and the inlet of line 8. Pump 28 ensures adequate hydrostatic pressure in input to line 8, so as to enable the degasser- conveyor 5 to operate as overflow discharge of reactor 4, similarly to the case occurring in the hydroconversion system shown in figure 1.

The hydroconversion system shown in figure 2 involves the use of a catalyst comprising or consisting of molybdenum, and differs from the hydroconversion system shown in figure 1 , also in that it comprises a line 27 for the introduction of hydrogen (or a gas including hydrogen) in separator 11 at a lower part of the same, i.e. where the liquid phase of the biphasic effluent and the degassed reaction liquid are collected. Preferably, the introduction of hydrogen takes place by means of a nozzle distributor-grid connected to line 27 and placed inside separator 11 , at the bottom of the same. The nozzle distributor causes the hydrogen to be preferably distributed throughout the cross section of separator 11. The hydrogen or the gas including hydrogen is introduced into separator 11 at a surface velocity preferably comprised between 0.01 cm/s and 1 cm/s.

Since reactor 4 is fed with a compound of molybdenum, as a result of the reaction of the latter with the charge to be converted or with the reaction liquid, a catalytically active species is produced consisting of molybdenum disulfide or "molybdenite" (S=Mo=S) in the solid state finely dispersed. In the presence of hydrogen dissolved in the reaction liquid, molybdenite is hydrogenated to dihydrosulfide (HS-Mo-SH), thus going to constitute the actual hydrogenating species capable, at the same time, to carry out the function of radical scavenger capable of inhibiting radical reactions leading to the formation of coke:

R ^* + HS-Mo-SH→ RH + S=Mo-SH without originating new radicals.

In the lack or, even more, in the absence of hydrogen dissolved in the reaction liquid, molybdenite is neither able to hydrogenate nor to oppose the formation of coke. On the contrary, it causes dehydrogenation and can generate living radicals (R ^* ) which are coke precursors:

S=Mo=S + HR→ S=Mo-SH + R ^* .

Since dihydrosulfide and not molybdenite is the hydrogenating species, the factor which determines whether the reaction proceeds towards hydrogenation or dehydrogenation is the hydrogen concentration in the liquid containing the catalyst, whatever the point along the hydroconversion process at which such a liquid is located (inside the reactor or outside it).

With reference to a predetermined volume of reaction liquid inside the reactor, a low hydrogen diffusion condition in the reaction liquid, through the liquid-gas bubble interface, with respect to the speed with which hydrogen is consumed from that volume of reaction liquid, causes the regression of molybdenum dihydrosulfide to molybdenite, with the consequent loss of the effect of dihydrosul- fide of inhibiting the formation of coke and the onset of dehydrogenation reactions catalyzed by molybdenite to generate coke precursor macroradicals. This condition occurs, in particular, by operating the reactor at high temperatures. Conversely, by limiting the temperature of reaction in such a way as to prevent the regression of molybdenum dihydrosulfide to molybdenite, the formation of coke during the reaction can be substantially reduced to zero. In known hydro- conversion systems, the regression of the molybdenum dihydrosulfide to mo- lybdenite is produced, at same temperature and hydrogen pressure, also as a consequence of a poor and sometimes completely lacking, hydrogen dispersion in the reaction liquid, as it usually happens at the bottom of the liquid-gas phase separator, where there is the formation of coke with possible fouling of the equipment, although the reaction liquid is at hydrogen pressure. As described above, in the hydroconversion system of invention, in order to prevent the formation of coke, the reaction liquid at bottom of separator 11 is cooled below the temperature threshold of "thermal cracking" and, in the meantime, hydrogen is dispersed in the liquid itself to prevent the regression of the molybdenum dihydrosulfide to molybdenite (the latter producing coke precursor living radicals even at temperatures below the above thermal cracking threshold).

In order to prevent the regression of molybdenum dihydrosulfide to molybdenite, the introduction and distribution of hydrogen, in addition to separator 11 , is also preferably carried out in other process steps external to reactor 4.

In particular, the hydroconversion system shown in figure 2 comprises a line 30 for introducing hydrogen or a gas including hydrogen in the flash bottom liquid (that is, in the liquid at the base of one or more depressurization tanks included in stage 15) and a line 31 for introducing hydrogen or a gas including hydrogen in the liquid at the base of the distillation column (which is also included in stage 15). Preferably, the introduction of hydrogen in stage 15, i.e. in the flash bottom liquid and in the liquid at the base of the distillation column, takes place by means of respective nozzle distribution grids connected to lines 30 and 31 and located within the depressurization tank (or depressurization tanks) and the distillation column, respectively, at the bottom of the same. The nozzle distribution grids ensure that the hydrogen is distributed uniformly, preferably in the entire cross section of the depressurization tank and of the distillation column.

The hydroconversion system shown in figure 2 comprises a line 32 for introducing hydrogen or a gas including hydrogen in the liquid to be converted, at line 26, upstream of the heat exchanger 9. Preferably, the introduction of hydrogen in line 26 takes place by means of a "static mixer" or other equivalent system for line dispersion of a gas in a liquid. The gas introduced into separator 11 through line 27 and/or of gas introduced in stage 15 through lines 30 and 31 and/or the gas introduced into line 26 through line 32 may consist of or comprise the same gas which is introduced in reactor 4 through line 1 , before a possible heating. Alternatively or in addition to this, said gas can consist of or comprise the hydrogen resulting from the above cooling and condensation stage. Alternatively or in addition to this, said gas can consist of or comprise the recycle hydrogen, i.e. the purified hydrogen resulting from the above purification stage.

The introduction of hydrogen can also be performed in the vacuum column included in stage 25. In this case, the hydroconversion system of the invention further comprises a line (not shown in the figure) for feeding the vacuum column with a compound capable of releasing hydrogen when it comes into contact with molybdenite, preferably a mono-cycle-alkane, such as for example cyclohex- ane. The mono-cycle-alkane is preferably fed dissolved in the hydrocarbon fraction extracted from the vacuum column itself.

Similarly to the description given with reference to the hydroconversion system shown in figure 1 , the continuous bleeding 19 allows to stabilize the level of ac- cumulation of the solids. By limiting the formation of solids (inorganic sulfides and coke) as described above, both inside and outside of reactor 4, within 3 kg per 1000 kg of processed charge, bleed 19 is greatly reduced. Also in the hydroconversion system shown in figure 2, high accumulation factors are therefore used, which are possible with low bleeding flow rates. In particular, only operating with an accumulation factor of at least 30, preferably at least 40, and even more preferably at least 50, a sufficient concentration of catalyst in the reaction is obtained, while operating at low catalyst reintegration through line 2. Operating conditions with a low production of solids allow to increase the value of the accumulation factor.

By operating under such conditions as to prevent the regression of molybdenum dihydrosulfide to molybdenite, in all the process steps, including any surge tanks, the formation of coke can be limited within 2 kg per 1000 kg of processed charge, preferably within 1 kg per 1000 kg of processed charge and even more preferably within 0.5 Kg per 1000 kg of processed charge, thus substantially zeroing the coking phenomenon of the catalyst and thus preventing the deactivation thereof. Accordingly, reduced molybdenum concentrations are required in the reaction medium in order to reach the plateau of the catalytic activity. Said reduced concentrations, in combination with the high accumulation factors which can be used, allow to reduce greatly the consumption (replenishing) of molybdenum:

consumption (replenishing) of molybdenum (ppm with respect to the processed charge) = (ppm molybdenum in the reaction medium) / (accumulation factor) With the hydroconversion system shown in figure 2:

• coke formation is substantially reduced to zero the coking of the catalyst therewith, which is thus preserved from deactivation phenomena;

• total solids (metal sulfides and coke) generated in the reaction are greatly re- duced (less than 3 kg per 1000 kg of charge fed, determined mainly by the content of metals in the processed charge). They can therefore be removed with limited bleeding flow rates, to the advantage of a higher degree of conversion which can substantially reach completeness;

• the accumulation factor can rise to values of 30 and beyond, which results in a drastic reduction of the amount of catalyst to be dosed. A dosage of 40 ppm molybdenum with respect to the charge fed causes a molybdenum concentration of 1200 ppm in the bleeding liquid, which corresponds to a concentration of molybdenum in the reaction medium of 1000 ppm. The possibility to operate with high accumulation factors while preserving the catalytic ac- tivity of molybdenum therefore entirely solves the economic problems related to the use of the latter as a catalyst for the hydroconversion of heavy oils. According to the description given as a preferred embodiment, it is apparent that some changes may be made by those skilled in the art without departing from the scope of the invention as defined by the following claims.