Description of Related Art Hydrogenation is utilized in many refining, petrochemical and chemical applications. In petroleum refining aromatics removal is important because there are specified limits for benzene content in gasoline and for aromatics content in diesel fuels. And due to environmental and healthy concerns those limits will be reduced even more in the future.

Nowadays, a process for selective aromatics removal from naphtha based on reactive distillation is also available (International Patent Application No. WO 96/27580, CD- Tech). The prior art process is based on the combination of a hydrogenation zone and a fractionation unit, wherein the hydrogenation zone is placed in the fractionation unit. It offers many advantages compared to conventional configurations. Lower operating pressures and fewer equipment items lowers investment costs compared to conventional aromatics removal processes.

Another known process is described in EP 0 781 830 A1 (Institut Français du Pétrole). It comprises a distillation column combined with a hydrogenation reactor. The process

configuration resembles that disclosed in US Patent No. 5,177,283 (Ward et al.) for hydrocarbon conversion processes.

The dearomatization process of EP 0 781 830 Al comprises two alternatives; viz. one in which all of the catalyst is placed in a side reactor and another one in which a part of the catalyst is placed within the column. In both embodiments, the sidedraw from the distillation column is conducted to the reactor. Hydrogen is fed into the sidedraw and the dearomatized effluent of the hydrogenation reactor is recirculated to the distillation column. In comparison to the conventional prior art catalytic distillation reactor taught in WO 96/27580, the hydrogenation can be optimized and, if desired, carried out at conditions different from the operation pressure and temperature of the distillation column.

The processes according to EP 0 781 830 A1 and WO 96/27580 are limited to the treatment of light hydrocarbon fractions and in particular to benzene removal therefrom.

The present invention deals with the dearomatization of middle distillates containing larger and more complicated molecules. This process is more complicated also because the cycle length of the catalyst is effected by the sulfur compounds in the feed; hydrogenation catalysts are, viz., very sensitive to sulfur and other impurities. After sulfur removal, middle distillate feed contains typically 1-300, occasionally up to 500, ppm-wt polyaromatic sulfur compounds like benzothiophenes and dibenzothiophenes which reduce catalyst life time in a aromatics hydrogenation reactor. A shorter cycle length causes extra operating costs.

Further, as a problem related to the process of EP 0 781 830 Al it can be mentioned that any excess hydrogen used for hydrogenation is recovered with the overhead stream of the distillation column and separated from the liquid recycle of the reflux drum. The presence of gaseous hydrogen in the distillation column impairs the operation of the distillation. No means are either provided for controlling the operation of the hydrogenation zone and in particular the heat generation in the hydrogenation zone. A too steep temperature rise of, say, more than 60 °C will adversely affect the activity of the catalyst of the zone.

Summarv of the Invention It is an object of the present invention to eliminate the problems related to the prior art of catalytic distillation reactors used for dearomatization of hydrocarbons and to provide a novel process for hydrogenation of hydrocarbon feedstocks, in particular middle distillates, containing sulphur-compounds.

This and other objects, together with the advantages thereof over known processes, which shall become apparent from specification which follows, are accomplished by the invention as hereinafter described and claimed.

The present invention is based on hydrogenation of middle distillate fractions in a system comprising a distillation column and one or more side-reactors. The sulphurous aromatic compounds contained in middle distillates are mainly high-boiling and they primarily gather in the bottom of a column during distillation of middle destillates. A side-cut taken from a point intermediate to the top and bottom of the distillation column, in particular intermediate to the top and the feedpoint of the hydrocarbon feedstock, will contain less sulfur than the feed. Thus, according to the present invention, hydrocarbons are fed into a distillation column or a similar fractionation zone providing fractionation of the hydrocarbon feedstock to produce a different hydrocarbon composition on each stage of the fractionation zone. A low-sulfur side-cut stream is withdrawn from the fractionation zone at a suitable stage and conducted to a reaction zone. The dearomatized reactor effluent is recycled to the distillation column. A dearomatized hydrocarbon product with sulfur compounds is recovered from the fractionation zone.

More specifically, the invention is characterized by what is stated in the characterizing part of claim 1.

Considerable advantages are achieved by the present invention. Thus, overall aromatics conversion of 50 to 95 % can be achieved. Sulfur compounds and other catalyst poisons are fractionated to the bottom and therefore catalyst cycle length is longer.

In comparison to conventional reactive distillation, the operating conditions of the column and the reactor are independent of each other. The operating temperatures and the pressures of the column and the reactor can be selected freely and optimised for each application.

Further, no need for special tower packing character is required, hence commercial trickle bed-type catalysts are adequate. No need to shut-down the unit for change of the catalyst (in the case of two or more side-reactors). No additional product stripper is needed.

Operation, start-up and control of both the column and the reactor are easier than with reactive distillation.

The present invention also provides important advantages compared to the prior art of EP 0 781 830 A1 and US 5,177,283. Hydrogenation reactions are exothermic and temperature rise in reaction zone can be controlled by a liquid recycle. For that purpose, preferably reactor effluent is cooled and sent to a gas-liquid separator. Temperature rise in the reactor is then controlled by liquid recycle from that gas-liquid separator. The gas mixture contains hydrogen and light C,-C4 hydrocarbons which can be recycled to the reactor and/or fed to fuel gas/hydrogen network. Liquid recycle is separated from the saturated hydrocarbons which are returned to the fractionation zone above or below the feed point of fractionation.

Preferably hydrogen is removed before the effluent is recycled which diminishes the volume of non-condensables in the column. Reaction zone effluent contains no unsaturated hydrocarbons because conversion in the reaction zone is 99-100 %. There is no need to separate hydrocarbon feed components from the dearomatized effluent of the reaction zone, as taught in US Patent No. 5,177,283. The fractionation zone is used as a product stabilizer for the reaction zone effluent which contains hydrogen and small amount of light hydrocarbons. Preferably the dearomatized product is recovered from the fractionation zone as bottoms, which provides for stabilization of the product. It can, however, also be recovered as a sidedraw.

Next, the invention will be examined in more closely with the aid of a detailed description and with reference to the attached drawing.

Brief Description of the Drawing Figure 1 depicts in a schematic fashion the process configuration of a preferred embodiment of the invention; Detailed Description of the Invention For the purpose of the present invention the term"aromatics"designates any aromatic compounds present in a petroleum-based hydrocarbon feed. Of particular importance for the present invention are the following typical aromatics of petroleum distillates: alkyl- substitued aromatics, such as butylbenzene, pentylbenzene and indene (CgH8) "Hydrogenation"and"dearomatization"are used interchangeably to denote a process, wherein unsaturated bonds of aromatic compounds are saturated by reacting the compounds with hydrogen (H2).

"Middle distillates"is an aromatic feedstock commonly used, for instance, as a fuel for diesel engines, comprising mainly C9-C2o hydrocarbons and having a boiling point area of about 150-400 °C, in particular 200 to 320 °C. It contains some 10 to 70 vol.-% aromatics before hydrogenation. Typical middle distillates are represented by straight run or cracked gasoil, FCC gasoil and visbreaker gasoil. The middle distillate feedstock of the present hydrogenation process is preferably subjected to a separate conventional desulfurization process, gasoil hydrotreater, before being fed into the present process.

The"sulfur compounds"of the middle distillates are mainly sulfides, disulfides, thiophenes, benzothiophenes, dibenzothiophenes and alkyl dibenzothiophenes. During a conventional desulfurization process, the lighter sulphur compounds are most easily hydrogenated and the heavy compounds are the last to react. Therefore, a gasoil which has undergone desulfurization does not contain light sulfides, thiophenes or benzothiophenes.

The distribution of dibenzothiophenes in the gasoil depends on the effectiveness of the desulfurization process, i. e. on the total sulphur content of the gasoil. Typical residual

sulfur compounds are 4,6-dimethyldibenzothiophene (4,6-DMDBT), 4-methyldibenzo- thiophene and dibenzothiophene.

Generally, in the present process for hydrogenation, a petroleum distillate feed containing C9-C20 hydrocarbons, like naftenes, parafines and aromatics, and polyaromatic sulfur compounds, like benzothiophenes and alkyl dibenzothiophenes, is fed to a fractionation zone. The fractionation zone includes a column, a reboiler, a condenser and an overhead drum. Typically the feed contains 20-30 vol-% aromatics and 1-100 ppm-wt sulfur.

From fractionation an overhead stream and a bottoms stream can be obtained having different hydrocarbon compositions. In particular the fractionation zone is operated so as to provide a different hydrocarbon composition in each stage of the zone. Thus, during fractionation heavier hydrocarbons including polyaromatic sulfur compounds are separated from the lighter components. The typical operating pressure range of the fractionation zone is 200-700 kPa. The temperature profile of the fractionation zone depends on the operating pressure and feed composition and should be optimized case by case.

The lighter fraction is taken from the fractionation zone as a side-cut stream and hydro- genated in a reaction zone. The hydrogenated stream from the reactor (s) is then recycled to the fractionation zone. Preferably the side-cut flow is taken from the fractionation zone above the fractionation feed point. The dearomatized reactor effluent can be recycled to the fractionation zone to a stage below the feed point (stage). The effluent from the reaction zone can be divided into two streams, the first of which is recycled to the reaction zone and the second of which is recycled to the fractionation zone (1). The first stream can be combined with the side draw fed to the reaction zone.

The side draw fed into the reaction zone contains less sulfur than the middle distillate feed.

Generally, the decrease in the sulfur content is at leat 50 %, in particular at least 90 %, so that the side draw generally contains a maximum of 50 ppm-wt sulfur.

The reaction zone includes heat exchangers, a trickle-bed type reactor and gas-liquid separator. The feed from the fractionation zone containing typically less than 50 ppm-wt, in particular 0-10 ppm-wt sulphur, and 10-30 vol-% aromatics, is cooled or heated to the

optimal reaction temperature before it is conducted into the reactor. The temperature at the reactor inlet depends on feedstock properties, on the hydrogenated aromatics and on the catalyst used. Aromatics are saturated in the presence of hydrogen and an active hydrogenation catalyst.

The hydrogen used for hydrogenation can be comprised of hydrogen gas of any suitable purity (typically 10-100 wt-%). The gas can contain 0 to about 90 wt-%, preferably only 10 to 70 wt-% of other volatile components, such as hydrocarbons, which remain inert during hydrogenation. Thus, hydrogen gas containing up to about 50 to 60 wt-% of methane and other light paraffinic hydrocarbons can be employed.

The catalyst will be discussed in greater detail below. However, in this context should be noted that using a nickel catalyst, hydrogenation can be carried out at the following conditions: -hydrogen-to-hydrocarbons-ratio: 100 to 500 Nm3 hydrogen/m3 hydrocarbons of the feedstock, -pressure: 35 to 55 bar, - (feed) temperature: 150 to 250 °C, -space velocity (LHSV): of 0.5 to 101/h With a noble metal catalyst the feed temperature typically lies in the range of 250 to 350 °C and the pressure is 35 to 60 bar.

To avoid too high temperatures and to avoid too great temperature differences over the catalyst beds, which will have a detrimental influence on the catalyst life, the temperature rise is kept below 60 °C during the reaction. According to the invention, the feed is diluted with the reaction zone effluent to lower the concentration of aromatics in the feed by cooling the effluent and partially recycling it to the reaction zone. The ratio of effluent recycle to fresh feed rate is 0.1 to 5, preferably about 0.2 to 3, and in particular about 0.5 to 3. By means of the liquid recycle the temperature rise in the reactor during hydrogenation can be restricted to less than 60 °C. Typically the temperature rise across the reactor is about 20 to 60 °C. In terms of feed dilution, a maximum aromatics concentration of about

16 vol.-%.

The cooled effluent or liquid reaction product is recycled from a gas separator vessel.

In addition to the above process steps, the reaction temperature rise can be restricted by hydrogen quenching. In that embodiment hydrogen is recirculated from the reactor effluent and fed into the reactor between the reaction beds. Hydrogen quenching can also be used to lower the temperature of the effluent feed between two serially arranged reactors, and it can be carried out by using make-up hydrogen.

Hydrogen-rich gas is separated from the liquid product in product separator and recycled to the reactor feed. To remove impurities in recycle gas some of the gas is bled to fuel gas or hydrogen net work. The ratio of recycled hydrogen to make-up hydrogen is about 0.1 to 10.

In general, the make-up hydrogen is somewhat richer in hydrogen: make-up contains, e. g., 89 to 92 mol-% H2 and recycle, e. g., 85 to 90 mol-% H2.

The process according to the present invention is carried out in a hydrotreating unit comprising at least one reactor, preferably two reactors arranged in series (in a cascade). It is also possible to arrange the reactors in parallel. As far as the basic concept of the invention is concerned, satisfactory results are obtained already with one hydrogenation reactor. However, because of the possibility of poisoining of the catalyst, two or more reactors are preferred to ensure continuous operation. Thus, when the catalyst in the first hydrogenation reactor is deactivated, the second reactor can be used, and the production continued. The cost of a reaction vessel together with necessary instruments is minor to the cost of production losses. Furthermore, the use of two reactors will reduce the amount of catalyst needed, because only the operative reactor needs to be filled with catalyst. If one big reactor is used, most of the catalyst is superfluous during the initial stage of the operation and therefore causes unnecessary costs. By using two reactors in a cascade it is in general possible to meet stricter limits for the aromatics contents of the products than with one reactor. The liquid reactor outlet from the first reactor can be cooled in a heat exchanger to the feed temperature of the first reactor, before it is fed into a second reactor or before a part thereof is recycled to the inlet of the reactor.

The reactors can be of any suitable type for contacting liquid with gas, trickle bed reactors being particularly preferred.

The saturated hydrocarbon stream contains some hydrogen and light (Cl-C4) hydrocarbons which are removed in the fractionation zone. The distillate of the fractionation zone contains saturated hydrocarbons. The bottom product of the fractionation zone contains hydrogenated hydrocarbons, polyaromatic sulphur compounds and small amounts of unsaturated heavier aromatics which were separated from the reaction zone feed.

Depending on the operating conditions the overall aromatics conversion of 50-95 % can be achieved. A significant advantage of this process compared to conventional ones (separated reactor and product stripper) is the possibility to control the aromatics content in the product. The control of the aromatics conversion enables the economical optimisation of the process.

It should be pointed out that the dearomatized product can be recovered also as a sidedraw.

According to the invention, the aromatics of a hydrocarbon feedstock are hydrogenated in the presence of a catalyst which is active in the temperature range of 50 to 300 °C. Suitable catalysts are heterogeneous catalyst comprising a metal of group VIII on a solid support.

Preferred metals are nickel, iron, platinum and palladium, and preferred supports are silica, alumina, magnesia, zirconia, titanium oxide, and natural and synthetic aluminium and magnesium silicates including natural and synthetic zeolites. According to a preferred embodiment, hydrogenation is carried out in the presence of a nickel catalyst, for instance a catalyst comprising elemental nickel on an inorganic metal oxide support. Such a catalyst can comprise 0.1 to 70 wt-%, preferably 1 to 50 wt-% nickel on an alumina or silica support. However, the selection of the catalyst depends on the amount of sulfur in the feed of the reactor. The afore-mentioned nickel catalyst is suitable for low-sulfur feeds. For feeds having high sulfur contents (up to 50 ppm or even more) Pt and/or Pd catalysts give longer operation cycles. The supports of these noble metal catalysts comprise alumina, silica or zeolite.

The process according to the present invention solves the problem of removing heavier

aromatics from sulfuric petroleum distillates. The sulfur compound in the feed are heavier polyaromatic sulphur compounds, like benzothiophenes, dibenzothiophenes and alkyl dibenzothiophenes. The hydrogenated heavier aromatics are alkyl benzenes, like butylbenzene, pentylbenzene etc. The process for hydrogenation of hydrocarbon feed in the distillation column with a side-cut stream to a reaction zone is especially advantageous for removing aromatics from from sulfuric petroleum distillates. With this configuration an overall conversion of up to 95 % can be achieved.

It should be pointed out that the sulfur specification for gasoils and similar middle distillates used as fuel for diesel engines is 500 ppm in many countries (e. g. U. S. A., Europe and Japan). However, the catalyst used for hydrogenation of aromates requires an essentially sulfur-free middle distillate feed (less than 50 ppm, preferably less than 10 ppm and in particular less than 5 ppm). As already briefly discussed above, the present process makes it possible to treat sulfurous feeds containing 50 to 100 ppm sulfur without any impact on the effective service life of the dearomatization catalyst. At the same time, also the costs of operation are reduced because a higher desulfurization degree would require higher temperature in the desulfurization reactor and higher operational temperature increase energy and catalyst consumption in the desulfurization section.

Turning now to the drawing, Figure 1 shows an embodiment particularly suited to the dearomatization of hydrocarbon fractions. The hydrocarbon feed containing aromatics is fed to a fractionation zone 1-5 at a feed point 2. The feed has preferably been subjected to desulfurization in a reactor not shown in the figure. The fractionation zone includes a column 1, a reboiler 3, a condenser 4 and an overhead drum 5. In the fractionation zone heavier hydrocarbons and sulfur compounds are separated from the lighter components.

The lighter fraction is taken from the fractionation zone as a side-cut stream 6 and hydrogenated in a reaction zone (7-10). The typical operating pressure range of the fractionation zone is 200-700 (or even higher, up to 1100) kPa. The temperature profile of the fractionation zone depends on the operating pressure and feed composition.

The reaction zone includes heat exchangers 8,9, a trickle-bed type reactor 7 and gas-liquid separator 10. The feed from the fractionation zone is cooled or heated in the heat exchanger

8 to the optimal reaction temperature before it is subjected into the reactor 7. The temperature at reactor inlet depends on the feedstock properties and hydrogenated aromatics and the catalyst used. Aromatics are saturated in the presence of hydrogen and an active hydrogenation catalyst. The H2 gas feed to the reaction zone contains 90-95 mol-% hydrogen and small amount of light hydrocarbons. The pressure in reaction zone is typically 3500-5500 kPa. Space velocity range is 0.5-10 1/h. Reactor effluent is cooled and sent into the gas-liquid separator 10. Temperature rise in the reactor is controlled by a liquid recycle (line 11) from a gas-liquid separator 10. The gas mixture obtained from the separator contains hydrogen and light C,-C4 hydrocarbons which can be recycled to the reactor and/or fed to fuel gas/hydrogen network. Liquid recycle is separated from the saturated hydrocarbons which are returned to the fractionation zone above or below feed point 12.

The following non-limiting calculated examples illustrate the invention.

Example 1 A middle distillate feed containing 50 ppm-wt polyaromatic sulphur compounds including dibenzothiophene, 4-methyldibenzothiophene and 4,6-dimethyldibenzothiophene was fed to a fractionation zone. The aromatic content of the feed was 23.5 wt-% C, o-C,3 mono-and diaromatics. The contents of the feed and the products are illustrated in Table 1. The fractionation zone included a column, a reboiler, a condenser and an overhead drum. In the fractionation zone heavier hydrocarbons including polyaromatic sulphur compounds were separated from the lighter components. The lighter fraction was taken from the fractionation zone as a side-cut stream and hydrogenated in a reaction zone. The operating pressure of the fractionation zone was 250 kPa. The temperature at the top of the column was 153 °C and 263 °C in the bottom of the column A side-cut flow which was taken from the fractionation zone above the fractionation feed point was subjected into reaction zone. The reaction zone included heat exchangers, a trickle-bed type reactor and gas-liquid separator. The feed from fractionation zone contained less than 1 ppm-wt sulfur, was cooled to the optimal reaction temperature before

it was subjected into the reactor. Aromatics were saturated in the presence of hydrogen and an active hydrogenation catalyst. The H2 gas feed to the reaction zone contained 92 mol-% hydrogen and small amount of C1-C4-hydrocarbons. The pressure in reaction zone was 4,000 kPa and the temperature at reactor inlet was 180 °C. Temperature rise in the reactor was controlled by a liquid recycle from a gas-liquid separator. The aromatics conversion of 99-100 % was achieved in the reactor zone.

Reactor effluent was cooled to a temperature of 40 °C and sent into the gas-liquid separator. Gas mixture containing hydrogen and light C1-C4 hydrocarbons was recycled to the reactor. Liquid recycle was separated from the saturated hydrocarbons which were returned to the fractionation zone below the feed point of the fractionation.

The saturated hydrocarbons from reaction zone contained some hydrogen and C1-C4- hydrocarbons which were removed in the fractionation zone. The liquid distillate of the fractionation zone contained light C1-C9 hydrocarbons. The bottom product of the fractionation zone contained saturated hydrocarbons, polyaromatic sulphur compounds and 4.8 wt-% unsaturated heavier aromatics. An overall aromatics conversion of 80 % was achieved.

Table 1 Component Feed, kg/h Bottoms, kg/h Light HC's, kg/h Height HC's, kg/h H, C-C6 hydrocarbons 1 482 812 409135560C7-C9hydrocarbons 100081757C10parafines 500471974C10naphthenes 45002610C10aromatics 4000378012C11parafines 2000108305C11naphthenes 89991664C11aromatics C12732617499 C12 naphthenes 3156 C, 2 aromatics 4000 888 Dibenzothiophene C, 3 parafmes 21000 20740 59993705C13aromatics Cl3 naphthenes 2323 C14 parafines 12000 11933 1400013970C15parafines 1099910992C16parafines 30002999C17parafines C, 8 parafmes 500 500 Dimethyl-dibenzothiophene 1 Methyl-dibenzothiophene 3 TOTAL 100000 99286

Example 2 A middle distillate feed (C10-C20 hydrocarbons) containing heavier aromatics having up to 16 carbon atoms and 100 ppm-wt polyaromatic sulfur compounds was fed to the distillation equipment illustrated in Example 1. The aromatic content of the feed was 26 wt-%. The contents of the feed and the products given in Table 2. The temperature at the top of the column was 157 °C and 277 °C in the bottom of the column.

A side-cut flow which was taken from the fractionation zone above the fractionation feed point contained sulfur compounds less than 1 ppm-wt. An aromatics conversion of 99 to 100 % was achieved in the reactor zone.

The bottoms product of the fractionation zone contained saturated hydrocarbons, polyaromatic sulfur compounds and 11.6 wt-% unsaturated heavier aromatics. An overall aromatics conversion of 57 % was achieved.

Table 2 kg/hBottoms,ComponentFeed, kg/h LightHC's, kg/h HC's, kg/h H23 C,-C6 hydrocarbons 1 373 752 4091293118C7-C9hydrocarbons mes 2000 817 133 1 30007479160C10naphthenes C, o aromatics 5000 32 14 4000382315C11parafines 400098437C11naphthenes 59991253C11aromatics 500049031C12parafines C122343 C, 2 aromatics 3000 698 Dibenzothiophene 5 5 C, 3 parafines 9999 9893 C, 3 aromatics 5999 3795 2238C13naphthenes 59995971C14parafines 99999982C15parafines C, s naphthenes79 C, 5 aromatics 2000 1921 C, 6 parafmes 6999 6999 Dimethyl-dibenzothiophene I I Methyl-dibenzothiophene 3 3 53C16naphthenes 50004947C16aromatics Cl, parafines 11999 11999 C20parafines 9999 9999 991852000TOTAL100000