Hydroconversion of heavy oils and hydrotreatment of distillates using a descending liquid bubble column slurry catalyst reactor

Abstract

Descending liquid bubble column slurry reactor, equipped with hydrogen distribution means with high density of orifices, is used to hydroconvert, at high capacity, heavy oils The reactor allows heavy distillates to be hydrogenated to saturation so that they can be used as feedstocks for steam cracking.

Description

[0001] “HYDROCONVERSION OF HEAVY OILS AND HYDROTREATMENT OF DISTILLATES USING A DESCENDING LIQUID BUBBLE COLUMN SLURRY CATALYST REACTOR”

[0002] Field of application of the invention

[0003] Processing of crude oil to obtain, in place of the usual energy carriers (fuels and fuel oil), hydrocarbons suitable for use as raw materials for steam crackers. Review of the prior art

[0004] As a result of the energy transition towards a low-carbon energy system, refining will have to produce less and less fuel oil and, due to the decarbonisation of transport, less and less fuels. In other words, less and less hydrocarbons to be used as energy carriers will be required.

[0005] Among the technologies currently used on a commercial scale to convert crude oil into products, the technology using bubble column reactors with a dispersed (slurry) catalyst, called “slurry hydroconversion”, does not involve the production of fuel oil. From a process that uses this technology, a set of products is obtained including LPG, naphtha, atmospheric gas oil and vacuum gas oil. Of these products, however, only the lightest, as they can be used as raw materials to fuel steam crackers, can be placed on a market other than that of energy carriers.

[0006] Due to the presence of aromatic structures and unsaturated hydrocarbons, the heavy distillates present in the product slate, those having a boiling point higher than that of naphtha, hereinafter referred to as “extra-naphtha distillates”, cannot be conveniently used to feed steam crackers.

[0007] For the same reason, even extra-naphtha distillates obtained directly from crude oil cannot be used as feedstocks for steam crackers, except in rare cases of the use of a highly paraffinic crude oil.

[0008] However, it is good to note that even an extra-naphtha distillate can be used to feed steam crackers if it is previously subjected to a hydrogenation treatment aimed at its dearomatization and elimination of unsaturated structures.

[0009] Applied to extra-naphta distillates, a hydrogenating treatment would allow you to divert towards petrochemical uses, at least in part, the surplus in distillates that the energy transition, in particular following the decarbonization of the transport, will entail. A hydrogenating treatment of the extra-naphtha distillates, obtained from the distillation of crude oil and deriving from the hydroconversion of a vacuum residue, it would also allow the use of crude oil exclusively to fuel chemical productions, without these productions having to involve, unlike what currently occurs, production of fuels when pursuing the decarbonisation of transport.

[0010] Chemical production could then continue in its growth trend without being in conflict with the energy transition.

[0011] According to this patent application, the aforementioned hydrogenation of extra- nafta distillates is implemented in a slurry catalyst bubble column reactor.

[0012] To bring the extra-naphtha distillates to complete saturation, the reactor used must have a high hydrogenation capacity and must be able to operate with high hourly flow rates for an economically sustainable hydrogenation process.

[0013] Once the conditions have been defined for operating at high hydrogenation rate, these can also be used in the slurry hydroconversion process of heavy oils to improve a not entirely satisfactory unit capacity, this being limited to about 0.1 m3 / h of heavy oil (vacuum residue) converted per cubic meter of reactor volume.

[0014] In the following considerations, relating to slurry catalyst bubbles column reactors, we will make use of the following definitions: • unit hourly flow rate = hourly charge flow rate per reactor section unit (m3 / h per m2, or m / h);

[0015] • hydroconversion unit capacity = hydroconversion capacity per unit volume of reactor (m3 / h per m3, or IT1); from which:

[0016] • unit hourly flow rate (m / h) = reactor height (m) x unit hydroconversion capacity (h-1);

[0017] • reactor capacity (m3 / h) = reactor section (m2) x unit hourly flow rate (m / h).

[0018] In a bubble column reactor, once the pressure is fixed, the contact surface area between gas bubbles and liquid is the physical quantity that determines the diffusion of hydrogen from the gaseous phase to the reaction liquid, therefore the hydrogenation rate. In order for a bubble column reactor to operate with a high hydrogenation rate it will have to create a bubbling gas-liquid mixture with a high gas-liquid unit surface (gas-liquid contact surface area in m2per m3of reactor volume).

[0019] Recently, it has been found that the gas-liquid unit surface is expanded when a gas-liquid mixture consisting of small diameter gas bubbles, packed to the limit of coalescence, is created within the reactor.

[0020] As described in PCT / IT2022 / 050288, in the hydroconversion of a heavy oil and in the hydrotreatment of distillates, small diameter packed gas bubbles are obtained by feeding hydrogen to the reactor base using gas distribution means with a high orifice density, at least 100 orifices per m2, and simultaneously reducing the surface velocity of the fed hydrogen in relation to the increase in the density of the orifices.

[0021] Unfortunately, having to lower the surface velocity of the hydrogen, and with this the extraction of products at the head of the reactor, means that, in order to increase the gas-liquid unit surface according to the aforementioned PCT, the hourly flow rate of the fed charge must be reduced thus lowering the capacity of the reactor.

[0022] Object of the invention

[0023] With reference to a bubble column reactor, using slurry catalyst, the aim of the present invention is to make the extraction of the reaction liquid and products from the reactor independent of the surface velocity of the hydrogen, with a flow rate beyond the current limits.

[0024] Summary of the invention

[0025] The reactors currently used for the slurry hydroconversion of heavy oils are of the upflow type. In this type of reactor, where the gas flow and the liquid flow are ascending and both the gas and the liquid exit at the head, the extraction of the reaction liquid and the products depends on the surface velocity of the hydrogen, which cannot be increased beyond the current limits.

[0026] According to the present invention, so that the extraction of reaction liquid and products can be carried out without flow limitations and is independent of the surface velocity with which the hydrogen is fed, the flow of liquid inside the reactor is diverted downwards to be extracted from the bottom.

[0027] The bubble column reactor, with slurry catalyst, used in the process object of the present invention, is provided at the base with gas distribution means which introduce hydrogen into the reaction liquid. The charge to be treated, consisting of a heavy oil, for example a residue from vacuum distillation of crude oil, or a distillate, is fed, not to the bottom of the reactor, as occurs in bubble column reactors of the upflow type, but rather above said gas distribution means.

[0028] At the bottom of the reactor, below the gas distribution means, instead of feeding the charge to be treated, liquid (reaction liquid) is extracted with a flow rate such as to keep the height of the bubbling liquid constant in an upper part of the column.

[0029] The extracted liquid is then sent directly, as it is gas-free, to a distillation section having a final stage under vacuum for the recovery of high-boiling conversion products.

[0030] The vacuum residue produced by the distillation section is recycled to the reactor by being introduced above the hydrogen distribution means.

[0031] In order for the liquid to be diverted downwards inside the bubble column, the vacuum residue from the distillation section is fed to the reactor with a flow rate greater than that with which the heavy oil or distillate is fed.

[0032] In a bubble column reactor, in which the flow of the liquid is descending, the extraction of the reaction liquid at the bottom, being independent of the surface velocity of the hydrogen which continues to exit at the head, is implemented, without limitations, at a predefined hourly flow rate.

[0033] The descending liquid bubble column slurry reactor, in order to expand the gasliquid contact surface and therefore operate with high hydrogenation capacity, employs gas distribution means with a high density of orifices, even if said means require to be fed with a reduced surface velocity of the hydrogen, but without the use of said means leading to a loss of capacity of the reactor. This is because the descending liquid bubble column slurry reactor allows freely predefined reaction liquid extraction flow rates.

[0034] The high hydrogenation capacity thus achieved is used to raise the productivity of the slurry reactor in the hydroconversion of heavy oils and to obtain from the hydrotreatment of distillates their complete hydrogenation so that they can be used as raw materials to feed steam crackers.

[0035] The hydrogenation capacity achieved is such that a single hydrotreatment reactor is capable of hydrogenating the entire quantity of hydrocarbons necessary to feed even a large-scale steam cracking unit.

[0036] Description of the figure

[0037] Further objects and advantages of the present invention will be clear from the following detailed description of examples of its implementation and from the attached drawing (figure 1 ), given purely for explanatory purposes, which schematically shows a dispersed catalyst bubble column reactor characterized by a liquid flow which, unlike an upflow reactor, is directed downwards, for an extraction of reaction liquid carried out on the bottom, without flow rate limitations and with a flow rate that is independent of the surface velocity of the hydrogen.

[0038] The scale and proportions of the various elements depicted do not necessarily correspond to the real ones.

[0039] Detailed description of some preferred embodiments of the invention

[0040] How to govern the extraction of conversion products, a critical issue that has always limited the performance of upflow reactors used in the hydroconversion of heavy oils, is the aspect that is addressed first.

[0041] Product extraction In the hydroconversion of heavy oils and in the hydrotreatment of distillates in bubble column upflow reactors, with the use of a dispersed catalyst, the oil to be treated is fed to the bottom of the reactor, while the hydrogen is fed, at a higher height, by means of gas distribution means placed in a lower portion of the column. Both the gas, in the form of bubbles, and the liquid are ascending and both, gas and liquid, exit at the head of the reactor.

[0042] In the gas-liquid mixture exiting the reactor head, there is a gaseous phase which includes unreacted hydrogen and reaction products in the gaseous state and in the vapor state. In said gas-liquid mixture, the liquid phase is made up of reaction liquid including the high-boiling conversion products to be extracted and the residual oil. The reaction liquid contains, in dispersed form, the catalyst and insoluble material generated by the reaction. The gas-liquid mixture is treated in a phase separator, at the head of which the gaseous phase containing hydrogen and the conversion products volatile to the reaction conditions are obtained.

[0043] The remaining products are recovered from the reaction liquid, collected at the bottom of a separator, in a subsequent distillation section with a final vacuum stage. The residue from the vacuum distillation, containing the catalyst and insoluble residue, is recycled in the reaction.

[0044] In the specific case of the hydroconversion of a heavy oil, for example a vacuum residue (540°C+), carried out in 50 meter high commercial reactors, due to insufficient extraction, the conversion products accumulate in the reaction liquid reaching a concentration of 50%, and above, thus limiting the concentration of residue in the reaction liquid to no more than 50%. Since the reaction volume is occupied for half or more by already converted products, the hydroconversion process of the residue is consequently slowed down.

[0045] At the head of an upflow reactor, the extraction of the products is favored by a higher surface velocity of the hydrogen which, however, cannot be increased beyond certain limits as it would lower the liquid filling of the reactor. In a bubble column slurry reactor, hydrogen is thus fed with a surface velocity that typically does not exceed 2.5 cm / s. This limit in the superficial velocity of hydrogen means that in commercial reactors used in the hydroconversion of heavy oils, despite having reached heights of 50 meters or higher, the flow of hydrogen (m3 / h per m2of reactor section, i.e. m / h , therefore cm / s) must remain the same, even when the hourly charge rate supplied and the flow of products to be extracted increase.

[0046] As the height of the reactor increases, extracting more products at the same superficial velocity of the hydrogen means that, as a result of a decreasing ratio between hydrogen flow rate and charge flow rate, there must be a natural accumulation of these products in the reactor. Given that the accumulation of reaction products reduces the unit capacity to hydroconvert heavy oils (m3 / h per m3), the production of an upflow reactor can never, for this reason, be proportional to its height.

[0047] In the hydroconversion of heavy oils, as in the hydrotreatment of distillates, to ensure an extraction of the products which can instead increase with the height of the reactor and, more generally, can increase with the capacity of the reactor itself, for example following a higher hydrogenation capacity, the reaction liquid must be able to be extracted from the reactor regardless of the dragging action that the hydrogen performs at the head of an upflow reactor.

[0048] In order to be able to extract the reaction liquid from a bubble column reactor without the aid of the gaseous flow exiting the reactor head, the reaction liquid, usually ascending, is diverted downwards to be extracted at the bottom, in the manner described below.

[0049] With reference to figure 1 , the bubble column reactor 1 is equipped at the base with gas distribution means 2. Said means are fed with hydrogen, or gas including hydrogen, via line 3. The charge to be treated, consisting of a heavy oil, for example a residue from the vacuum distillation of a crude oil, or a distillate, is fed, above said gas distribution means, via line 4. Below the means of gas distribution, instead of feeding the charge to be treated, as happens in upflow reactors, liquid (reaction liquid) is extracted through line 5 with a flow rate such as to keep the height 6 of the bubbling liquid constant in an upper portion of the column .

[0050] The reaction liquid extracted at the bottom of the reactor, being gas-free, is sent directly to a distillation section for the recovery of the high-boiling conversion products. After passing through a flash tank, not shown in figure 1 , the reaction liquid is sent to an atmospheric distillation 8 and a vacuum distillation 10, for the extraction of the atmospheric gas oil in 9 and the vacuum gas oil in 11. The residue 12 of the vacuum distillation 10 is recycled to the reactor above the gas distribution means 2.

[0051] By purging 13, insoluble residues generated by the reaction are removed.

[0052] In order for the liquid to be diverted downwards inside the bubble column, the vacuum residue coming from the distillation section is fed to the reactor, via line 12, above the gas distribution means 2, with a flow rate greater than that of heavy oil, or distillate, fed into 4.

[0053] As the flow rate of said residue to the reactor increases, the flow of reaction liquid that is extracted in 5 increases. The method to be followed to feed the reactor a flow of residue higher than that of the feed from which it comes is described in example 1 reported below.

[0054] Used in a hydroconversion process of heavy oils, or hydrotreatment of distillates, the descending liquid bubble column slurry reactor, since the reaction liquid is extracted from it without flow limitations, allows you to choose the concentration of the products at which to extract them .

[0055] The concentration of the products can thus be lowered to eliminate the loss of unit capacity that occurs, as the height increases, when the reactor is of the upflow type. The same ratio between maltene and asphaltene hydrocarbons, present in the reaction liquid, can be lowered as necessary to avoid the precipitation of the asphaltene fraction which, in upflow reactors, occurs when the height of the reactor is excessive in relation to the hydrogen fed.

[0056] A low concentration of products in the reaction liquid, which allows for a higher residue concentration, is chosen to increase the cracking capacity, so that, when the hydrogenation capacity is also high, it is possible to increase the capacity of a slurry reactor to hydroconvert heavy oils.

[0057] Below it is shown how, and to what extent, a descending liquid bubble column slurry reactor allows you to reduce the percentage of converted products in the reaction liquid. The case described is that relating to the hydroconversion of a heavy oil consisting of vacuum residue (540°C+). Example 1. With reference to figure 1 , reactor 1 , 50 m high, is fed via line 3 with hydrogen at a surface velocity of 2.5 cm / s which generates a gas hold-up of 0.33. The catalyst is introduced through an oil-soluble compound of one or more transition metals. The feed (vacuum residue, density = 1 kg / dm3) is fed through line 4 with a flow rate of 5000 kg / h per m2of reactor section. This feed, being referred to a reactor height of 50 metres, corresponds to a unit hydroconversion capacity of 0.1 m3 / h per m3of reactor volume, a typical value for slurry hydroconversion, however not entirely satisfactory.

[0058] The reactor, above the gas distribution means, is fed vacuum residue from the distillation section, VR<DS), at an hourly flow rate of 10500 kg / h per m2.

[0059] When starting the reactor, to feed vacuum residue 12 with a flow rate greater than that of the feed fed in 4, proceed as follows. The surge tank that receives the residue VR<DS) from the vacuum column 10, not shown in figure 1 , is initially filled with straight run vacuum residue VR<SR), i.e. with the same charge to be convert. As startup progresses and the reactor reaches steady state, the straight run residue VR<SR) present in the surge tank is naturally and definitively replaced by VR(SD).

[0060] Above the bubbling liquid, whose pre-established height is positioned in an upper portion of the reactor, due to the foaminess of the heavy oils, the gaseous flow that rises towards the outlet at the head of the reactor produces a foam in which traces of reaction liquid are present. The gaseous flow is freed from the reaction liquid by passing through a cyclone separator, not shown in figure 1 , with the separated liquid returning to the reactor.

[0061] In the conditions currently used: temperature 400-450°C and pressure 100-200 atmospheres, the hydrocarbons produced by the hydroconversion of heavy oils that exit the reactor head in the gaseous and vapor state are mainly made up of hydrocarbons with a boiling point lower than 300° C. Therefore the products that exit the reactor head through the gaseous phase constitute approximately 30% by weight of the feed introduced in 4, therefore in the present example they are equal to approximately 1500 kg / h per m2of reactor section. The remaining conversion products (high boiling), for a total flow rate of 3500 kg / h per m2, are recovered through the distillation section. These high-boiling products are part of the (3500+10500)=14000 kg / h (or 14 m3 / h) per m2of reactor section of reaction liquid extracted from the bottom of the reactor, through line 5, where these high-boiling are therefore present in a percentage of 25% (100x3500 / 14000).

[0062] In the same reaction liquid the residue to be converted is present in a percentage of 75% (100x10500 / 14000), therefore significantly increased compared to the value of 50%, or less, found in the upflow reactors currently used. This brings a proportionate positive impact on the conversion capacity of the reactor.

[0063] Incidentally, in the descending liquid bubble column reactor, under the conditions described above, the average velocity of the bubbles relative to the reactor (surface velocity of hydrogen / gas hold-up) takes on a value of 2.5 / 0.33 = 7 .57 cm / s. The surface velocity of the descending liquid, which is extracted with a flow rate of 14,000 kg / h per m2, i.e. 14 m3 / h per m2, is equal to 14 / (1-0.33) = 20.9 m / h, i.e. 0.58 cm / s: the bubbles inside the reactor continue to rise, ensuring mixing of the reaction medium, despite the liquid flow being directed downwards.

[0064] Having found the way to govern the extraction of the conversion products in a bubble column reactor, we will now look at how to expand the gas-liquid contact surface so as to be able to increase its hydrogenation capacity.

[0065] Expansion of the gas-liquid unit surface

[0066] In the slurry hydroconversion of heavy oils, as in the slurry hydrotreatment of distillates, the bubble column reactor is currently equipped at the base with gas distribution means having approximately 50 orifices per m2Such a distribution of hydrogen, due to the low gas-liquid unit surface it generates, allows a unit hydrogenation capacity limited to about 0.0035 t / h per m3of reactor (3.5% of hydrogen incorporated in the products, from 0.1 m3 / h of converted feed per m3of reactor volume).

[0067] As described in PCT / IT2022 / 050288, the gas-liquid unit surface area, and with it the hydrogenation capacity, grows when a bubbling mixture consisting of small diameter, packed bubbles is generated inside the reactor. To generate such a favorable fluid dynamic condition, the reactor is equipped with gas distribution means with a high density of orifices, at least 100 per m2, fed with hydrogen at a surface velocity which must be reduced as the density of orifices increases. For example, the surface velocity of hydrogen must be halved to achieve a doubling of the gas-liquid unit surface. In general, the surface velocity, expressed in cm / s, with which the hydrogen is fed to the distribution means, must not exceed: 2.5 I (orifice density 150)1 / 3, the ratio (orifice density / 50) being dimensionless.

[0068] As previously anticipated, such a lowering of the surface velocity of hydrogen however reduces the extraction capacity of the conversion products and with this, consequently, the unit capacity of the reactor: a negative implication to be overcome.

[0069] Slurry reactor with high hydrogenation capacity

[0070] The descending liquid bubble column slurry reactor, since it implements an extraction of the reaction liquid, and therefore of products, independent of the surface velocity of the hydrogen, to expand the gas-liquid contact surface, and therefore operate with high hydrogenation capacity, can use gas distribution means with a high density of orifices, without the use of such means having to lead to reductions in capacity.

[0071] Below, by way of example and without limitation, the use of the slurry reactor with high hydrogenation capacity will be described to carry out the hydroconversion of heavy oils with high unit capacity. Subsequently, its use for hydrogenating heavy distillates to saturation for their chemical use will be described.

[0072] Slurry catalyst reactor with high hydroconversion capacity

[0073] In the following example, the descending liquid bubble column slurry reactor, equipped with high orifice density hydrogen distribution means, is used to operate with high concentration of residue in the reaction liquid and to simultaneously generate a bubbling mixture with high gas-liquid contact surface. Example 2. Again referring to figure 1 , the reactor 1 , with a height of 50 m, in this case is equipped at the base with gas distribution means 2 having 400 orifices per m2, fed via line 3 with hydrogen, or gas including hydrogen, at a surface velocity of 1.25 cm / s. With such a gas distribution, the unit gas-liquid surface area is doubled compared to that generated by conventional distribution means, powered with hydrogen at a surface velocity of 2.5 cm / s and having 50 orifices per m2The hydrogenation capacity of the reactor, from a usual value of 0.0035 t / h per m3, rises to 0.0061 t / h, i.e. it increases by 1.74 times. It does not increase twice as much as the gas-liquid unit surface due to the reduction in the partial pressure of hydrogen along the height of the reactor as a result of the increased consumption.

[0074] The charge to be converted, consisting of a 540°C+ vacuum residue, is fed via line 4, above the gas distribution means 2, with a flow rate of 7500 kg / h (or 7.5 m3 / h) per m2of reactor section. Through line 5, reaction liquid is extracted from the bottom of the reactor, with a flow rate such as to keep the level of the bubbling liquid positioned in an upper part of the reactor constant. The reaction liquid thus taken from the bottom of the reactor, via a flash drum not shown in the figure, is sent to the distillation section to extract atmospheric gas oil in 9 and vacuum gas oil in 11. The residue from the vacuum distillation 10 is recycled to the reactor, above to the means of distribution 2.

[0075] The vacuum residue of the distillation 10 is collected in a tank, not shown in the figure, used as a surge tank to feed said residue to the reactor. The residue is fed above the gas distribution means 2, via line 12, with a flow rate equal to 15750 kg / h per m2.

[0076] To feed residue with a flow rate greater than the feed fed, the same reactor start-up method previously described in example 1 is followed.

[0077] From the top of the reactor comes out the gaseous phase 7 containing the unreacted hydrogen and the volatile conversion products, the latter corresponding to approximately 30% by weight of the feed, as seen in example 1 , therefore equal to approximately 2250 kg / h per m2. The remaining high- boiling conversion products, which are extracted from the reactor as liquids, are recovered via the distillation section, at a flow rate of approximately 7500-2250 = 5250 kg / h per m2of reactor section. In the reaction liquid extracted at the bottom of the reactor there are high-boiling conversion products in a percentage equal to 100x5250 / (5250+15750)=25% and a residue to be converted in a percentage equal to 100x15750 / (5250+15750)=75%. The high concentration of residue in the reaction liquid and the greater availability of hydrogen, provided by a descending liquid bubble column slurry reactor equipped with gas distribution means with a high density of orifices, bring the hydroconversion capacity to 7.5 m3 / h per m2, for a reactor height of 50 meters. The unit capacity thus rises to 0.15 m3 / h per m3, from a value of 0.1 m3 / h per m3found in conventional upflow reactors. With a height of 50 meters and a diameter of just over 5 metres, a single reactor, which will therefore have a useful reaction volume of 1000 m3, converts 1 ,314,000 t / y of vacuum residue (540°C+) .

[0078] A greater availability of hydrogen (1.74 x) compared to the increase in conversion capacity (1.5 x), improves the degree of hydrogenation, denitrification and desulfurization of the conversion products. It also reduces, to the point of eliminating, the formation of insoluble residues to be disposed of and therefore the consumption of catalyst.

[0079] A greater availability of hydrogen also lowers the pour point of the vacuum distillation residue 12.

[0080] Due to the high hydrogenation capacity it achieves, the descending liquid bubble column slurry reactor, equipped with gas distribution means with a high density of orifices, can be used to convert residues even with a high asphaltene content, such as the final pitch produced from the only partial hydroconversion of heavy oils when implemented with ebullated catalytic bed reactors.

[0081] Slurry catalyst reactor with descending liquid for the hydrogenation of distillates to saturation

[0082] In the distillates obtained from a crude oil, as in the distillates obtained from the slurry hydroconversion of heavy oils, light hydrocarbons are present, currently used as steam cracker feedstocks. Together with these there are hydrocarbons with higher boiling points, in particular atmospheric gas oil and vacuum gas oil, rich in aromatic structures and unsaturated hydrocarbons, which are used to essentially produce only energy carriers. As an alternative to the production of energy carriers, distillates with a boiling point higher than that of naphtha (extranaphtha distillates), if subjected to a hydrogenation treatment to be freed from aromatic and other unsaturated hydrocarbons, can be used, even if of high molecular weight, as a raw material to feed steam crackers.

[0083] It has been found that the slurry reactor with high hydrogenation capacity, described above, when used for the hydrotreatment of distillates, and is operated at a temperature of 350°C or lower, is capable of bringing the distillates to total saturation, making them suitable for feeding steam crackers. The distillates can be those produced by the slurry hydroconversion of a heavy oil, but they can also be the extra-naphtha distillates of a first distillation of the crude oil.

[0084] Example 3. With reference to figure 1 , the descending liquid bubble column reactor 1 , equipped with gas distribution means 2 having 400 orifices per m2, fed with hydrogen with line 3 at a surface velocity of 1.25 cm Is, generates a gas-liquid mixture with a high gas-liquid unit surface which, compared to a conventional reactor, increases the hydrogenation capacity by 1.74 times: from 0.0035 to 0.0061 t / h per m3, as seen previously, in example 2.

[0085] The reactor always uses a dispersed catalyst based on oil-soluble transition metal compounds.

[0086] The feed to be hydrotreated, in this case an extra-naphtha distillate, is fed to reactor 1 via line 4.

[0087] The reaction liquid, which in this case contains hydrogenated hydrocarbons, is extracted from the bottom of the reactor, via line 5, with a flow rate such as to keep the level of bubbling liquid constant, indicated with 6 in figure 1 , in an upper part of the column.

[0088] From the reactor head, through line 7, a gaseous phase emerges including hydrogen and gaseous products of the hydrotreatment (NH3, H2S and volatile hydrocarbons), without liquid carryover.

[0089] The liquid extracted from the bottom of the reactor is sent to a distillation section to obtain hydrogenated distillates. The residual liquid from the distillation section is collected in a tank, not shown in figure 1 , to be fed to the reactor through line 12, above the gas distribution means 2, with a flow rate equal to twice that of the feed introduced with line 4, so as to return the catalyst to the reactor. Insoluble material in limited quantities is purged through line 13. The reintegration of the catalyst, to compensate for the purge 13, is of limited extent. The reactor is started in a manner similar to that described in example 1 .

[0090] Reactor 1 temperatures higher than 350°C, activating thermal cracking, would lead to additional consumption of hydrogen to saturate the free radicals generated. Such a consumption of hydrogen would thus be subtracted from the hydrogenation of aromatics and unsaturated hydrocarbons. Since the reduction of the length of the hydrocarbon chain is not necessary when using hydrogenated distillates to feed steam crackers, reactor 1 operates at a temperature that does not exceed 350°C. To shift the hydrogenation balance towards saturated hydrocarbon structures, which lead to a higher quality index of the distillates when used as raw materials for steam crackers, reactor 1 operates at temperatures lower than 350°C, for example 300°C.

[0091] By using a reactor 1 temperature of 350°C or lower, the incorporation of hydrogen into the distillate occurs substantially while maintaining the length of the hydrocarbon chain.

[0092] An extra-naphtha distillate, with a hydrocarbon chain length equal to n carbon atoms, for a value of n equal to or greater than 12 can contain hydrogen in a maximum percentage corresponding to that of a paraffin with n=12 carbon atoms: 100 X H2n+2 I Cn+H2n+2 = 1 5.3%.

[0093] Therefore, a hydrogenating hydrotreatment which, by maintaining the length of the hydrocarbon chain, is able to bring the hydrogen content of an extranaphtha distillate to a percentage of 15.3%, is a hydrogenating hydrotreatment which transforms an extra-naphtha distillate into a saturated hydrocarbon.

[0094] If an extra-naphtha distillate, which can be assumed to have a hydrogen content of 13%, is subjected to a hydrogenating treatment, the complete hydrogenation process will transfer to the distillate a quantity of hydrogen equal to a maximum of 2.3% of its weight.

[0095] In the hydrotreatment carried out using the descending liquid bubble column slurry reactor, equipped with hydrogen distribution means with high orifice density, described above, the hourly flow rate of extra-naphtha distillates hydrogenable at saturation is given by the ratio between the hydrogenation capacity of the reactor and the hydrogen incorporated by the hydrogenated products. With a hydrogenation capacity equal to 0.0061 t / h per m3and an incorporation of hydrogen into the distillate of 0.023 t per m3, a single reactor, having the previously described geometry, therefore having a volume of 1000 m3, is capable of produce at least (0.0061 / 0.023) x 1000 = 265 t / h, equal to 2,320,000 t / year, of saturated hydrocarbons, usable as a steam cracker feed. It is noted that by using hydrogenated extra-naphtha distillates to feed chemical steam crackers, the chemical production obtained will see the increasingly sought-after BTX (benzene, toluene, xylenes) prevail at the expense of olefins, which can also be obtained from other feedstocks.

[0096] It is also noted that since the % of hydrogen in an extra-naphtha distillate, even before the hydrogenation hydrotreatment, is largely higher than that of the BTX produced downstream of steam cracking, an integrated unit: hydrogenating hydrotreatment of extra-naphtha distillates + steam cracking, produces hydrogen (green).

[0097] Depending on the changes that the energy transition will impose, the extranaphtha refinery distillates, possibly obtained from the slurry hydroconversion of vacuum residue, currently used to obtain energy carriers, could be hydrogenated to be used to feed chemistry.

[0098] The hydrotreatment process with high hydrogenation capacity, described above, can be used in the hydrogenation of any hydrocarbon, whatever its origin, whether mineral, vegetable or synthetic. It can also be used to hydrogenate Fishers-Tropsch waxes.

[0099] Processing of crude oil into products that can be used entirely to feed the petrochemical industry

[0100] \Ne will now see how, through the complete hydroconversion of the vacuum residue and the hydrotreatment for the hydrogenation of extra-naphtha distillates, the products of crude oil refining can be totally used to feed chemical production, via steam cracking, without having to produce fuels or other energy carriers.

[0101] It is assumed that the steam cracking unit of a petrochemical complex will be fed starting from heavy crude oil, having a distillable hydrocarbon content equal to 65%, the remaining 35% being therefore made up of vacuum residue. The distillable fraction includes LPG and naphtha, in quantities equal to 20% of the crude oil, fed directly to the steam crackers.

[0102] It will now be seen how, through a complete hydroconversion of the vacuum residue and the hydrotreatment of the extra-naphtha distillates, the hydrocarbons obtained from the crude oil can be entirely used to feed a petrochemical complex, via steam cracking, without having to produce fuels or other energy carriers.

[0103] It is assumed to feed the steam cracking unit of a petrochemical complex starting from a heavy crude oil having a content of distillable hydrocarbons equal to 65%, the remaining 35% being constituted by the residue of a vacuum distillation. The distillable fraction includes LPG and naphtha, in quantities equal to 20% of the crude oil, which are fed directly to the steam crackers. By slurry hydroconversion, the vacuum residue is totally converted into distillates. The extra-naphtha distillates, obtained from the hydroconversion of the vacuum residue and the extra-naphtha distillates obtained from the primary distillation of crude oil, are both subjected to a hydrogenating treatment to generate saturated hydrocarbons which feed the steam crackers. Crude oil can thus be entirely used to fuel a petrochemical complex. For the wider validity of the following example, it is assumed that the petrochemical unit is of a large scale, for a production of 3,000,000 t / y of chemical products, obtained using 3,000,000 t / y of heavy crude oil. The residue to be converted will therefore be equal to 1 ,050,000 t / y and can be hydroconverted using the slurry reactor described in example 2 which has adequate capacity.

[0104] The slurry hydroconversion will generate a production of 250,000 (approximately 1 / 4 of 1 ,050,000) t / y of light hydrocarbons (LPG and naphtha), and 800,000 t / y of extra-naphtha distillates to be hydrogenated. These distillates obtained from the hydroconversion of the residue, together with 1 ,350,000 (45% of 3,000,000) t / a of extra-naphtha distillates from primary distillation, for a total of 2,150,000 t / a, are hydrotreated for hydrogenation to saturation. The hydrogenation treatment is carried out using the reactor described in example 3, which has a capacity to bring 2,320,000 t / a of distillates to saturation. The crude oil processed in this way, with total conversion of the vacuum residue and hydrogenation of the extra-naphtha distillates, is monetized as naphtha and becomes entirely usable to fuel chemical production without these having to conflict with the decarbonisation of transport. On the basis of the description provided for some application examples, it is obvious that some changes can be introduced by the person skilled in the art without thereby departing from the scope of the invention as defined by the following claims.

Claims

AMENDED CLAIMS received by the International Bureau on 26 June 2024 (26.06.2024)1. Process for the hydroconversion of heavy oils, including crude oils, and for the hydrotreatment of distillates by means of a bubble column reactor using a dispersed catalyst, said reactor comprising internally gas distribution means, said means of distribution:• being fed with hydrogen or with gas including hydrogen,• being housed in a lower portion of said reactor, said process including:• an extraction, in a distillation section downstream of said reactor, of conversion products from the reaction liquid produced by said reactor and• a recycling of residue from said distillation section to said reactor, said process being characterized by the fact of:• feeding said heavy oil or said distillate to said reactor above said means of distribution,• feeding said residue of said distillation section to said reactor above said means of distribution,• feeding said residue of said distillation section to said reactor with a flow rate greater than that with which said heavy oil or said distillate is fed,• extracting, from the bottom of said reactor, below said means of distribution, reaction liquid with a flow rate such as to keep constant the level of the bubbling liquid in an upper portion of said reactor,• feeding said reaction liquid extracted from the bottom of said reactor to said distillation section.

2. Process according to claim 1 , characterized in that:• said distribution means comprise a plurality of orifices, the number of orifices present being not less than 100 per m2of horizontal section of said distribution means, said number of orifices per m2of horizontal section of the means of distribution being called “orifice density” andAMENDED SHEET (ARTICLE 19)• hydrogen or gas including hydrogen is fed to said distribution means at a surface velocity, expressed in cm / s, not exceeding2,51 (orifice density / 50)1 / 3, the ratio (orifice density I 50) being dimensionless.

3. Process according to claim 2, characterized in that the reactor operates at a temperature not higher than 350°C.AMENDED SHEET (ARTICLE 19)

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