The present invention relates to a method for operating a hydroprocessing process, in which a sulfided catalyst comprising a base metal hydroprocesses a feed in a position upstream a reduced catalyst comprising noble metal catalyst hydroprocesses the feed further, without intermediate purification.

In hydroprocessing specific catalysts for specific chemical transformations are often used, and the physical process conditions, such as temperature and pressure, are optimized for the specific processes. When the catalysts being used comprise noble metals, it is, however, also necessary to consider the chemical process conditions, such as the composition of the reaction atmosphere as a process condition, in the sense that the presence of sulfur must be very low, for the noble metal catalyst to function. In this situation the choice of a noble metal catalyst has implications in that it requires additional equipment for purification.

Now, according to the present invention it is realized, that it is possible to operate a process with a base metal catalyst requiring the presence of sulfur in the same reactor as a noble metal catalyst accepting only a very low level of sulfur, if the catalysts are activated prior to loading in the reactor. To provide prolonged operation the operating conditions may beneficially be selected such that the catalysts are protected against inhibition by adsorption of chemical poisons reacting with the active sites of the catalyst surface and physical deactivation by carbon deposits blocking the access to the active sites of the catalyst.

Specifically the present disclosure may with benefit be implemented as a single stage process for hydrocracking and hydrodearomatisation or as a single stage process for hydrotreatment and hydroisomerisation, but other processes may also be implemented according to the present disclosure.

Base metals shall in the context of the present application be construed as the group 6 metals such as W or Mo and the non-noble group 8,9 or 10 metal such as Ni or Co, e.g. called NiMo or CoW. A base metal catalyst often comprise one of the metal W and Mo as well as one of Ni and Co, but base metal catalysts may, however, also be active even if comprising only a single of these metals. Base metal catalysts have som activity in their elemental form, but the highest activity is in the sulfide form and therefore, base metal catalysts are often called sulfided catalysts.

Noble metals shall in the context of the present application be construed as Pd, Pt, Ru, Rh, Os and Ir. A noble metal catalyst often comprise a single noble metal, but may also comprise two or more in combination. Noble metal catalysts are produced by precipita- tion of salts, such as Pt(NOs)2, but ionic Pt ⁺ or PtO is not catalytically active. Therefore the Pt ⁺ must be reduced to be catalytically active, and noble metal catalysts may also be called reduced catalysts.

A catalyst support shall in the context of the present application be construed as a sta- ble material having a high porosity and a high surface area, typically silica, alumina or silica-alumina. Catalyst supports may also have a chemical function, and can contain zeolite.

Hydrodesulfurization (HDS) shall in the context of the present application be construed as a process in which organic sulfur in hydrocarbons is released as hydrogensulfide by reaction with hydrogen. Other heteroatoms may be released in the same process stage.

A hydrocracking (HDC) process shall in the context of the present application be con- strued as a process in which hydrocarbons are reacted in the presence of hydrogen to form smaller hydrocarbons. Materials catalytically active in HDC may also catalyze ring opening of aromatics, in which a bond of an aromatic ring is broken without decomposition of the molecule. A hydrodearomatisation (HDA) process shall in the context of the present application be construed as a process saturating aromatic compounds by reaction with hydrogen. A hydroisomerisation (HI) process shall in the context of the present application be construed as a process in which linear hydrocarbons - especially paraffins - are reacted in the presence of hydrogen to form branched hydrocarbons. Hydrodeoxygenation (HDO) shall in the context of the present application be construed as processes in which oxygenates are converted to hydrocarbons in the presence of hydrogen.

Hydrotreatment (HDT) shall in the context of the present application be construed as the group of processes in which hydrocarbons are converted to other hydrocarbons of similar size in the presence of hydrogen. HDT includes a.o. HDS, HDA and HI.

Hydroprocessing (HDP) shall in the context of the present application be construed as the group of processes in which hydrocarbons are converted to other hydrocarbons in the presence of hydrogen. HDP includes HDT and HDC.

The processes hydrotreatment and hydroprocessing processes above shall not be limited to processes proceeding by specific mechanisms, but only to processes performing the specified conversion in the presence of hydrogen, by one or more steps.

Boiling in the diesel range, shall in the context of the present application be understood as a hydrocarbon mixture of which at least 80% boils in the range 150-400°C.

A means for gas/liquid separation, shall in the context of the present application be understood as any means receiving a feed and providing more feeds separated based on their boiling point. This may include many types of devices, including but not limited to a flash drum separator or a stripper in which gas is driven from a liquid phase by a stripping medium but also a fractionator, or the combination of devices e.g. a flash drum and a fractionator receiving a liquid phase stream from the flash drum.

A stage, shall in the context of the present application be understood as a defined section of a process in which the temperature and pressure are decoupled from other sections, often due to removal of gaseous product prior to the stage. Specifically a two stage process shall be defined as a process in which an intermediate or a sideproduct, is removed within the process limits, and a one stage process shall similar be defined as a process in which nothing is removed within the prosces limits. Sweet operation or sweet stage, shall in the context of the present application be understood as the operation, or a defined stage of a process plant, in which the concentration of sulfur is relatively low.

Sour operation or sour stage, shall in the context of the present application be understood as a the operation, or a defined stage of a process plant, in which the concentration of sulfur is not relatively low. When referring to group X metals, reference is made to groups of the periodic system, reference is made to the lUPAC Periodic Table of the Elements dated 1 May 2013.

In a broad form the present disclosure relates to a process for reducing the amount of aromatics in a raw feed stream comprising hydrocarbons, more than 200 ppmw sulfur or 1000 ppmw sulfur as either hydrocarbon heteroatoms or as other sulfide compounds as well as at least 10% by weight di-aromatics or poly-aromatics and at least 30% by weight aromatics in total said process comprising the steps of

hydrotreating said raw feed stream in the presence of hydrogen and a material catalytically active in hydrotreatment with a severity resulting in a conversion of sulfur hydro- carbon heteroatoms to hydrogen sulfide of at least 50% providing a pre-treated stream, separating said pre-treated stream at least into a second stage feed stream and a stream rich in hydrogen sulfide,

directing said second stage feed stream to contact a material catalytically active in hydrocracking and ring opening, and to contact a material catalytically active in saturation of aromatics, wherein the material catalytically active in hydrocracking and ring opening is positioned upstream, downstream or mixed with said material catalytically active in saturation of aromatics, and withdrawing a dearomatized stream,

wherein said the amount of aromatics of said dearomatized stream is less than 50%, 70%, 90% or 95% of the amount of aromatics in said raw feed stream, with the associ- ated benefit of said process of providing efficient dearomatization with low yield loss.

In a further embodiment said material catalytically active in hydrocracking and ring opening first material comprises a base metal and is provided in presulfided form and said and said material catalytically active in saturation of aromatics comprises a noble metal and is provided in prereduced form with the associated benefit of avoiding a con- flict between conditions for in-situ sulfidation and in-situ reduction, and thus providing the possibility to operate a catalyst comprising a base metal and a catalyst comprising a noble metal in the same reactor, with the further benefit of a low catalyst cost from the use of base metal, while employing the preferred high performance of a noble met- al catalyst.

An alternative to the present disclosure relates to a process for conversion of a feed stream comprising hydrocarbons and/or oxygenates and from 1 ppmw sulfur or 10 ppmw sulfur to 100 ppmw sulfur, 200 ppmw sulfur or 1000 ppmw sulfur as either hy- drocarbon heteroatoms or as other sulfide compounds, said process comprising the steps of

a. directing said feed stream to contact a first material catalytically active in a first hy- droprocessing reaction under a first set of conditions favouring said first hydropro- cessing reaction, said first material comprising a sulfided base metal, providing a first hydroprocessed stream

b. optionally cooling said first hydroprocessed stream by addition of low temperature hydrogen,

c. directing all of said first hydroprocessed stream to contact a second material catalytically active in a second hydroprocessing reaction under a second set of condi- tions favouring said second hydroprocessing reaction, said second material comprising a noble metal, providing a second hydroprocessed stream

wherein said first material is provided in presulfided form and said second material is provided in prereduced form, and wherein the first material in sulfide form and the second material in reduced form are both stable at the first set of process conditions as well as at the second set of process conditions, with the associated benefit of avoiding a conflict between conditions for in-situ sulfidation and in-situ reduction, and thus providing the possibility to operate a catalyst comprising a base metal and a catalyst comprising a noble metal in the same reactor, with the further benefit of a low catalyst cost from the use of base metal, while employing the preffered high performance of a noble metal catalyst.

In a further embodiment the first material comprises one or more metals taken from the group comprising group 6 elements, Ni and Mo, with the associated benefit of these metals having a moderate cost and a high ability to tolerate organic sulfur and hydrogen sulfide, and to convert organic sulfur to hydrogen sulfide.

In a further embodiment the second material comprises one or more metals taken from the group comprising Ru, Rh, Pd, Os, Ir and Pt, with the associated benefit of these metals providing a highly specific catalytic process.

In a further embodiment the second material comprises Pt and Pd, with the associated benefit of this combination of noble metals having an higher tolerance towards sulfur compared to individual noble metals.

In a further embodiment the maximum temperature of the first catalytically active material is 250°C-350°C, with the associated benefit of providing a set of conditions in which the first hydroprocessing reaction is active, while both of the catalytically active materi- als are maintained in their active form.

In a further embodiment the maximum temperature of the second catalytically active material is 250°C-350°C, with the associated benefit of providing a set of conditions in which the second hydroprocessing reaction is active, while both of the catalytically ac- tive materials are maintained in their active form.

In a further embodiment the difference between the outlet temperature of the first catalytically active material and the inlet of the second catalytically active material is less than 50°C, with the associated benefit of avoiding operational, material and mechanical instability from operating with a high temperature gradient in a reactor.

In a further embodiment the first hydroprocessing reaction is hydrocracking, with the associated benefit of providing a first hydroprocessed stream which is having a favorable range of molecular weight or a reduced amount of aromatics.

In a further embodiment the first hydroprocessing reaction is hydrotreatment, with the associated benefit of providing a first hydroprocessed stream having a reduced amount of heteroatoms such as nitrogen, sulfur and oxygen. In a further embodiment the second hydroprocessing reaction is hydrodearomatization, with the associated benefit of providing a first hydroprocessed stream which is having a high cetane number and a low aromatics content. In a further embodiment the second hydroprocessing reaction is isomerization, with the associated benefit of providing a first hydroprocessed stream which is having high branching and thus favorable cold flow properties.

In a further embodiment the feed stream comprises at least 10% by weight di- aromatics or poly-aromatics and at least 30% by weight aromatics in total, and the first hydroprocessing reaction is hydrocracking and the second hydroprocessing reaction is saturation of aromatics, with the associated benefit of such a process being able to provide a product rich in high quality diesel from a feedstock with poor diesel qualities. A further embodiment of the present disclosure relates to a process for hydrocracking and aromatics saturation of a raw feed stream comprising hydrocarbons and more than 200 ppmw sulfur or 1000 ppmw sulfur as either hydrocarbon heteroatoms or as other sulfide compounds comprising the steps of

i. hydrotreating said raw feed stream in the presence of hydrogen and a base metal catalyst with a severity resulting in a sulfur content of hydrocarbon heteroatoms below 200 ppmw or 1000 ppmw providing a pre-treated stream comprising said feed stream of claim 1 and a further amount of hydrogen sulfide, ii. separating said pre-treated stream at least into said feed stream and a stream rich in hydrogen sulfide.

iii. directing said feed stream of claim 1 to undergo the process described above, with the associated benefit of such a process being able to convert e.g. a sulfur laden LCO/gas oil feed into a product rich in high quality diesel.

In a further embodiment the feed stream comprises at least 1 % or at least 10% linear oxygenates, and in which the first hydroprocessing reaction is hydrodeoxygenation and the second hydroprocessing reaction is hydroisomerization, with the associated benefit of converting a feed rich in oxygenates, such as a biological feedstock or a Fischer Tropsch product into a diesel fuel with low yield loss. A further aspect of the present disclosure relates to a process plant for conversion of a stream of heavy aromatic hydrocarbon mixture into a hydrocarbon mixture rich in middle distillate comprising a first stage reactor unit containing a hydrotreatment catalyst, said first stage reactor unit having an inlet and an outlet, a means for gas/liquid separa- tion having an inlet and a gas outlet and a liquid outlet, and a second stage reactor unit comprising one or several reactors and containing at least a presulfided material catalytically active in hydrocracking and a prereduced material catalytically active in hydrodearomatization, said second stage reactor unit having one or more inlets and a single outlet,

in which the stream of heavy aromatic hydrocarbon mixture is in fluid communication with an inlet of the first stage reactor,

the outlet of the first stage reactor unit is in fluid communication with the inlet of the means for gas/liquid separation,

the outlet for liquid of the means for gas/liquid separation is in fluid communication with the inlet of the second reactor unit,

a stream of hydrogen is optionally in fluid communication with a further inlet of the second reactor,

the outlet of the second reactor unit provides the hydrocarbon mixture rich in middle distillate,

with the associated benefit that such a process plant allows for cost effective moderate purification of the hydrotreated heavy hydrocarbon mixture, providing an appropriate decoupling of hydrotreatment and hydrocracking and sufficient desulfurization of the stream directed to the material catalytically active in hydrodearomatization. A further aspect of the present disclosure relates to a process plant for conversion of a feed stream comprising oxygenates in to a hydrocarbon mixture rich in branched middle distillate comprising a reactor unit containing at least a presulfided material catalytically active in hydrodeoxygenation and a further prereduced material catalytically active in hydroisomerization, said reactor unit having one or more inlets and a single outlet, in which the feed stream comprising oxygenates is in fluid communication with the inlet of the reactor unit and

the outlet of the reactor unit is provides a hydrocarbon mixture rich in branched middle distillate, with the associated benefit that such a process plant allows for cost effective high yield conversion of oxygenates, e.g. of biological origin to diesel fuel or diesel component with a good flow property. In catalyzed hydroprocessing the catalyst comprises a metal oxide support and an active metal. The function of the support is mainly to provide a high dispersion of the active metal, which provides a hydrogenation function by chemisorption of hydrogen. In some situations the catalyzed process also requires an acidic function, which typically is provided by the support.

The hydrogenation function may be provided by either of the main categories of catalyt- ically active metals; noble metals and base metals.

The most common noble metals are in this respect mainly Pd and Pt but also Ru, Rh, Os and Ir may be used. The noble metals may be used alone or in combinations. It has been observed that combinations of noble metals may have specific properties and the combination may also be chosen to reduce the cost of the catalytically active material. When the catalyst is produced the noble metals are distributed on the catalyst surface by precipitation of salts, such as Pt(NOs)2, but ionic Pt ⁺ is not catalytically active in hy- droprocessing. Therefore the Pt ⁺ must be reduced to be catalytically active. This reduction is performed at elevated temperature in the presence of hydrogen, which provide conditions where the thermodynamically stable form of Pt is the elemental form and where the conversion from ionic to elemental Pt is kinetically fast. It must be realized that Pt(NC>3)2 is merely used as an example, other anions than NO3 ^2" are commonly used and the metal may of course be any of the catalytically active noble metals. The activity of noble-metal catalysts is very high, so the metal content can be very low, such as 0.05% or even 0.01 % and up to 1 %, 5% or 10%. A further benefit of noble-metal catalysts for HDA is that a noble metal catalyst typically operate at moderate temperatures, at which the equilibrium of the HDA reaction favours saturation, which again will favor a reduced pressure, all of which are beneficial for the overall cost of equipment and operation. For other processes the noble metal catalysts may have other benefits; e.g. for HI the noble metal catalyst will be more specific towards isomerization instead of cracking, and thus typically using a noble hydroisomerisation catalyst will reduce the yield loss. The name "noble metal" indicates a high level of inertness. . Noble metal catalysts are, however, highly sensitive to the presence of poisons such as sulfur, nitrogen, carbon monoxide and carbon dioxide in the reaction environment. The poisons are attracted to the surface of the metal, and will provide a high extent of deactivation of the catalytic activity. For this reason, noble metal catalysts have been operated under so-called sweet conditions, i.e. in the substantial absence of a.o. sulfur. Research has been carried out in which the catalytically active materials have been doped with addition elements, to provide moderate robustness of noble metal catalysts to the presence of sulfur, but only with limited success. One example of increased robustness is the ability to operate a noble metal catalyst comprising Pd and Pt at moderate sulfur levels up to 150 ppm as demonstrated a.o. in US 6,462,244.

The most common base metal catalysts are combinations in pairs of a group 6 metal such as W or Mo and a non-noble group 8,9 or 10 metal such as Ni or Co. Base metal catalysts may, however, also be active even if comprising only a single of these metals. Base metal catalysts are somewhat active in their elemental form, but the highest activity is in the sulfide form. This means that base metal catalysts beneficially are activated in a sulfide atmosphere at elevated temperatures. Activation is expedited by increased sulfide concentration often 3-10% and elevated temperatures often 350°C to 400°C. The activity of sulfided base metal catalysts is lower than that of noble-metal catalysts with a similar metal content, so base metal catalysts typically have a higher metal content compared to noble metal catalysts, such as 5-30%.

Contrary to the noble metal catalysts, base metal catalyst are not only able to operate in the presence of sulfur, but actually the presence moderate levels of sulfur is a requirement for maintaining the sulfided form, and thus for the stable operation of base metal catalysts. Therefore in the operation of sulfided catalysts, sulfide may have to be added to a sweet feed to ensure stable operation. It is known in the art of refining to use processes which combine the cost effective and robust base metal catalysts with the more specific noble metal catalyst. Such a process is carried out as a two-stage process in which the first stage is based on operation in a sour environment with a robust sulfided base metal catalyst, which typically catalyzes hydrodesulfurization which releases organic sulfur into hydrogen sulfide. After this stage a separation process removes the gaseous hydrogen sulfide, and other unde- sired gases, and the hydrocarbon which is sulfur free - or having a very low sulfur content - is transferred to a so-called sweet stage, in which a reduced noble metal catalyst operates. In such a process, the separation requires a drop in temperature and pres- sure, and the hydrocarbon stream must be heated and re-pressurized in the downstream process. This has an energy cost, as well as the cost of additional equipment.

When operating such a process, it is simple to activate the catalysts; in the sour stage a sulfidation material such as dimethyl disulfide (DMDS) may be added to an oil stream in a concentration of 3-10% and directed to contact the catalytically active material at a temperature of 340°C to 400°C, which will sulfide the base metal, - as well as noble metals if such were present. Similarly noble metal catalysts of the sweet stage must be activated by reduction in a substantially pure hydrogen atmosphere at an elevated temperature of 300°C to 320°C. When increasing the temperature further platinum may sinter, with a reduced active surface area as the result. These conditions would, however, also reduce the base metals from their sulfided state. Accordingly the activation of base metal catalyst and noble metal catalyst in the same reactor is not practically possible, and furthermore even the operation of such a reactor has been considered impossible due to the contradicting process requirements, where the base metal cata- lyst requires that sulfur is present during operation which is incompatible with the required operation of the noble metal at low sulfur conditions. Therefore the operation of base metal catalysts and noble metal catalyst have been carried out in separate stages, or at least in process configurations where sulfur and other poisons released in an upstream sour operation was carried away from the sweet operation e.g. by a gas flow which was counter-current to the liquid flow, since the risk and inconvenience related to the immediate incompatibility has not been realized to be outweighed by any benefits.

Now according to the present disclosure it has been realized that the use of pre- activated catalyst, in combination with a choice of moderate process conditions makes it possible to operate a process in which a sulfided base metal catalyst is loaded immediately upstream a reduced noble metal catalyst, e.g. in the same reactor, with surprising benefits. Methods for ex-situ pre-activation of catalytically active materials are well known to the skilled person, and may in principle be carried out by the methods described above or by other methods, e.g. by aqueous reaction with sulfides. In addition to the problems of poisoning of catalytically active materials comprising noble metal catalysts by sulfur, noble metal catalysts are also sensitive to other compounds, including CO, NH3 and organic nitrogen and a strategy for operation must take this into account.

Hydroprocessing depends on the socalled severity of process conditions and on the activity of the catalyst. The severity reflects the temperature, the pressure and the presence of mildly deactivating species. The catalyst activity is defined with respect to a specific process (HDS, HDA, HDC etc.) and is influenced by parameters such as the catalyst composition, the catalyst morphology and the chemical composition of the reaction mixture. To the skilled person the implementation of the present invention may therefore require an amount of experimental work, identifying the limits of operation allowable in the specific case. This is realized by the skilled person and will not in practice prohibit the implementation of the present disclosure.

When operating and designing a process according to the present disclosure, it is preferred that the activities of the catalysts are balanced such that operation of the beds of catalytically active material in the same reactor is within the same temperature range, to avoid thermal stress of materials as well as runaway processes due to heat flux from hot beds to beds designed to be cold. Typically this is calculated as the difference between the temperature at the outlet of one bed and the temperature at the inlet of the following bed, and the value is desired to be below 50°C or even 30°C.

The present disclosure is well suited for a process for production of diesel fuel from a feed being a mixture comprising at least 30% by weight aromatics wherein at least 10% of the mixture is polyaromatics such as a mixture of light cycle oil and gas oil. Such a feed mixture requires removal of aromatics, which typically has been carried out by hy- drocracking for ring opening of the aromatics. If the sulfur content of the feed is high, the feed may require a pretreatment reducing the sulfur content to an acceptable level, such as below 200 ppm, by a customary hydrodesulfurisation and separation.

Now, with increasing demands for low aromatic content and high cetane values, the hydrocracking may be insufficient for obtaining the desired product, and it may be considered implementing a process combining pretreatment, hydrocracking and saturation of aromatics by hydrodearomatization. When three reactions are required for conversion of a feed (here HDS, HDC and HDA) and these are carried out in sequence their operational conditions become coupled, unless the reactions are decoupled by being operated in separate stages, e.g. with intermediate cooling and depressunzation. Such intermediate cooling and a depressunzation may also be required to remove excessive sulfur prior to the reaction mixture contacting noble metal catalysts.

Operation of a hydrotreatment catalyst, such as a hydrodesulfurisation catalyst, is typically related to a deactivation over time. As a consequence the process severity (typi- cally the operational temperature) is increased gradually with time. If the hydrotreatment is not separated from the hydrocracking - e.g. if the gas/liquid separation is positioned immediately before hydrodearomatization - the increase in pretreatment temperature also defines the temperature in the hydrocracking section, which may result in a more rapid deactivation of the hydrocracking catalyst due to sintering and possibly also coking with a risk of runaway of the exothermal HDC reaction. Furthermore the increased temperature also results in increased cracking, and thus increased yield loss.

Now if, according to the present disclosure, the majority of sulfur is removed prior to hydrocracking, the operation of the hydrotreatment section and hydrocracking section becomes decoupled. This may reduce the yield loss and increase the lifetime of the catalyst since it becomes possible to conduct hydrocracking at less severe conditions.

In a similar manner a process for the combined hydrodeoxygenation and hydroisomeri- zation of oxygenates, such as triglycerides, fatty acids and fatty acid esters, originating from biological sources, or alternatively from Fischer Tropsch processes may also be implemented by a first bed of base metal hydrodeoxygenation catalyst and a second bed of noble metal hydroisomerization catalyst, with the associated benefit of a low yield loss in the hydroisomerization bed, due to the use of noble metal catalyst. Biological feeds of oxygenates typically have a very low sulfur content, but it may be neces- sary to carry out pretreatment removing nitrogen compounds.

Fig.1 shows a process for conversion of a LCO/gas oil feed mixture to diesel according to the present invention. Fig.2 shows a process for conversion of a LCO/gas oil feed mixture to diesel according to the prior art.

In Fig.1 according to an embodiment of the present invention, a LCO/gas oil feed mix- ture in combination with hydrogen 2 is directed to contact a hydrotreatment catalyst 4 in a in a pretreater unit 6, in order to provide a pretreated feed 8. A gas phase 12, including hydrogen sulfide, is removed in a means of gas/liquid separation 10 (such as an interstage stripper or a flash drum), and the pretreated hydrocarbon feed 14 is combined with a stream of unconverted oil 40 and directed as second stage feed 16 to contact a base metal hydrocracking catalyst 18 in a hydrocracking catalyst bed in a second stage reactor 22. The entire effluent of the hydrocracking catalyst bed is transferred to contact a noble metal hydrodearomatization catalyst 20 in a separate catalyst bed or possibly in a separate reactor. The base metal catalyst 18 and the noble metal catalyst 20 can be operated in the same reactor, because both have been activated ex-situ and because a moderate level of sulfur is present in the reaction mixture. The second stage product stream 24 is first separated in a gas phase 28 and a liquid phase product 30 in a gas/liquid separator, and then fractionated in a fractionator 32 into naphtha 34, jet fuel 36 and an unconverted oil (UCO) fraction 38 dominated by product boiling in the diesel range. A part of the UCO 40 is directed to contact the hydrocracking catalyst and hydrodearomatisation catalyst again. This process has the benefit that the pressure and temperature of the HDC step 18 and the HDA step 20 may be optimized independently of the HDS step 4, and thus an increased specificity and yield may be obtained in the product 42 which is an unconverted oil stream comprising diesel. As the hydrocracking catalyst active in hydrodearomatization as well as hydrocracking, if it is operated at moderate temperature it will catalyze dearomatization by ring opening while it will catalyze hydrocracking at elevated temperature. Therefore, the inlet temperature to reactor 22 will make it possible to control the extent of hydrocracking, and accordingly it will be uncomplicated to switch between active hydrocracking (at el- evated temperatures) and ring opening and moderate hydrocracking at intermediate temperatures.

In Fig.2 according to the prior art, a LCO/gas oil feed mixture in combination with hydrogen 2 is directed to contact a hydrotreatment catalyst 4 in a in a pretreater unit 6, in order to provide a pretreated feed 8, A gas phase 12, including hydrogen sulfide, is re- moved in a means of gas/liquid separation 10 (such as an interstage stripper or a flash drum), and the pretreated hydrocarbon feed 14 is combined with a stream of unconverted oil 40 and directed as second stage feed 16 to contact a base metal hydrocracking catalyst 18 in a hydrocracking catalyst bed in a second stage reactor 22. The sec- ond stage product stream 24 is first separated in a gas phase 28 and a liquid phase product 30 in a gas/liquid separator, and then fractionated in a fractionator 32 into naphtha 34, jet fuel 36 and an unconverted oil (UCO) fraction 38 dominated by product boiling in the diesel range. A part of the UCO 40 is directed to contact the hydrocracking catalyst and hydrodearomatisation catalyst again. Compared to the process of Fig.1 , this process does not contain a specific HDA step, and thus the only reduction of aromatics will be due ring opening by the HDC catalyst. Accordingly, this process relative to the process of Fig.1 will result in a lower diesel quality (due to the higher aromatics content and the lower cetane number) as well as a higher yield loss (due to more severe process conditions, selected to increase the extent of ring opening.

In Fig.3 according to an embodiment of the prior art, a LCO/gas oil feed mixture in combination with hydrogen 2 is directed to contact a hydrotreatment catalyst 4 and a base metal hydrocracking catalyst 18 in a in a pretreater unit 6, in order to provide a pretreated feed 8. A gas phase 12, including hydrogen sulfide, is removed in a means of gas/liquid separation 10 (such as an interstage stripper or a flash drum), and the pretreated hydrocarbon feed 14 is combined with a stream of unconverted oil 40 and directed as second stage feed 16 to contact a noble metal hydrodearomatization catalyst 20 The second stage product stream 24 is first separated in a gas phase 28 and a liquid phase product 30 in a gas/liquid separator, and then fractionated in a fractionator 32 into naphtha 34, jet fuel 36 and an unconverted oil (UCO) fraction 38 dominated by product boiling in the diesel range. A part of the UCO 40 is directed to contact the hydrodearomatisation catalyst again.

In this configuration, the temperature of the hydrocracking catalyst is not independently controlled, as the effluent of hydrotreatment will be directed immediately to the hydrocracking catalyst. In addition, the presence of a high amount of sulfur will result in decreased catalyst activity, which therefore would require an elevated temperature - with the consequence of reduced ring opening selectivity. This will cause an elevated diesel yield loss at similar dearomatization levels compared to the configuration of Fig.1 .

Example 1

Example 1 discloses operation of HDS, HDC and HDA process in accordance with Fig.1 , i.e. in a configuration where HDC and HDA are operated in the same stage, independently of the HDS stage,

In Table 1 an example of operation according to the present process scheme (e.g. in accordance with Fig.1 ) is shown. The results were consistent over more than 2000 run hours. The example clearly shows the ability to convert LCO/gas oil into a high amount of quality diesel with good cetane properties, even in the presence of moderate levels of sulfur (25 ppm). Example 2

Example 2 discloses the influence of sulfur impurities in the feed to the second stage of a process such as the one shown in Example 1 .

Table 2 shows the effect of increased H2S on 2nd stage effluent properties. Experiments A and B were carried out with a commercial base metal HDC catalyst loaded in a first reactor, from which the entire effluent was directed to a second reactor loaded with a commercial noble metal HDA catalyst. From the results it is seen that in this setup with either 17 wppm S or 163 wpp S in the feed mixture, the yield loss (i.e. the fraction boiling below 221 °C) is the same and that the dearomatization is close to complete in both cases as well.

Example 3

Example 3 discloses in Table 3 the influence of carbon monoxide impurities in the feed to the second stage of a process such as the one shown in Example 1 . Experiments C and D were carried out in a setup similar to that of Example 2 - but with the catalysts loaded in the same reactor, hence the results are not directly comparable with those of Experiments A and B. Here the experiments shows that the presence of 140 ppm CO results in the same dearomatization, but a slightly increase in yield loss of 1 .5%. Example 4

Example 4 discloses in Table 4 the characteristics of two alternative pretreatments of the same LCO feed, corresponding to reactor 6 of Figure 1 and 3 respectively, using the same HDT and HDC catalysts as in Example 1 . Hydrotreatment followed by hy- drocracking in accordance with reactor 6 of Figure 3 (the column HDT+HDC) according to the prior art has a yield loss (i.e. the fraction boiling below 221 °C) of 34.7%, whereas hydrotreatment alone in accordance with reactor 6 of Figure 1 (the column HDT only) only has a yield loss of 10.3%. When examining the effect of the HDC step of example 1 , a similar Diesel yield loss is observed, but the dearomatization is much lower. To obtain a dearomatization effect of the HDC step of Example 4 corresponding to that reported in the column HDC eff. of Example 1 , would demand a significant further increase in yield loss. If the liquid fraction of the effluents of the two pre-treatments of Example 4 would be directed to a process corresponding to Example 1 , either directly to HDA (for the combined HDT+HDC effluent according to the prior art) or to HDC/HDA (for the HDT only effluent according to the present disclosure), it is clear that the high dearomatization activity of the base metal HDC catalyst in the semi-sweet environment of reactor 22 of Fig.1 is able to dearomatize more efficiently at lower or comparable yield loss than the similar catalyst in the sour environment of reactor 6 of Fig.3.

The dearomatization observed over the HDC catalyst is assumed to be due to partial cracking processes, in which aromatic rings are opened, without decomposing the molecules further. If the process severity is increased further, such ring opening may be increased, but so will decomposition of the molecules, resulting in increased yield loss.

The results of Example 4 in combination with Example 1 therefore confirm the benefit of upgrading LCO to diesel in accordance with the present disclosure, as the yield loss for similar dearomatization is lower. Table 1

FEED HDT eff 2nd-stg HDC eff HDA eff

FEED

Sulfur, wt ppm 4200 25 17 <1 <1

Nitrogen, wt ppm 1000 2 1 <1 <1

Aromatics, D6591

1 ring, wt% 21.7 52.8 32.7 8.7 0.6

2 rings, wt% 30 2.3 1.5 <0.1 <0.1

3+rings, wt% 5.9 0.2 0.1 <0.05 <0.05

Fraction<221°C,wt% 11.1 15.9 10.9 35.4 35.8

Yield loss 5.4% 27.5% 0.6%

Table 2.

CASE A B

Sulfur feed, wt ppm 17 163

Sulfur, wt ppm <1 <1

SG 60/60F 0.852 0.853

Aromatics, D6591

1 ring, wt% 0.57 0.55

2 rings, wt% <0.1 <0.1

3+rings, wt% <0.05 <0.05

Fraction<221°C,wt% 24.6 24.3

Table 3.

CASE C D

CO feed, wt ppm 0 140

Sulfur, wt ppm <1 <1

SG 60/60F 0.859 0.858

Aromatics, D6591

1 ring, wt% 0.56 0.55

2 rings, wt% <0.1 <0.1

3+rings, wt% <0.05 <0.05

Fraction<221°C,wt% 19.8 21 .3 Table 4:

FEED HDT only HDT+HDC

Sulfur, wt ppm 3626 <1 <1

Nitrogen, wt ppm 744 <1 <1

Aromatics, D6591

1 ring, wt% 29.1 55.7 30.5

2 rings, wt% 34.2 2.8 0.7

3+rings, wt% 1 1.5 0.3 0.1

Yield, wt%FF

<221 °C wt% 27.1 34.6 52.4

Yield Loss 10.3% 34.7%