FI100535B - Process for dearomatizing oil distillates - Google Patents

Description

, 100535 METHOD OF OIL DISTILLATES PEAROMATIS01M1SEKS1

The present invention relates to a process 5 according to the preamble of claim 1 for dearomatizing oil distillates. Generally, according to such a process, the aromatic hydrocarbon feedstock is fed to a hydrogenation process unit, the feedstock is contacted with hydrogen in the presence of a catalyst to hydrogenate the feed and aromatics contained therein to provide a dearomatized product, and the dearomatized product is recovered from the hydrogenation unit.

10

The present invention also relates to a dearomatization process according to the preamble of claim 19.

Many oil distillates contain aromatic compounds that are harmful for a variety of uses. Typical aromas of oil distillates include benzene and alkyl-substituted aromatic compounds such as toluene, tetramethylbenzene, diethylmethylbenzene, pentylbenzene and indene.

. Several dearomatization processes have been developed to produce aroma-free and low-aroma petroleum products. They are mainly based on the conversion of aromatic compounds into the corresponding saturated hydrocarbons by reacting said compounds with hydrogen in the presence of a suitable catalyst at elevated pressure and temperature. . you. After dearomatization, hydrogenated petroleum distillates are generally stabilized by the removal of light • · *: *: ‘: volatile hydrocarbons.

25 »:, · ,: Catalytic hydrogenation of aromatic hydrocarbons, especially benzene and toluene • · ·:.; · Also carried out in connection with the preparation of cyclohexane and methylcyclohexane.

: A conventional process for the hydrogenation of naphthenic hydrocarbons is disclosed in U.S. Patent 3,597,489. The hydrocarbon feed, consisting of pure benzene or toluene, is hydrogenated in two reactors connected to Saija in the presence of a nickel catalyst at a temperature between 140 and 200 ° C. The effluent from the first reactor is cooled and the vapor and liquid phases are separated. The liquid phase of the first reactor effluent is partially recycled back to 2,100,535 in the first reactor, while the vapor phase is passed to the second reactor.

A common feature of all known dearomatization / hydrogenation processes is that they can only be applied to the specific feed for which they are designed. Thus, the catalyst and the operating conditions used differ so much that, for example, the dearomatization process developed for middle distillates cannot be used for the conversion of a light gasoline fraction. Similarly, the solvent hydrogenation process cannot be used to dearomatize middle distillates.

Furthermore, a hydrogenation process developed for pure feedstock components, such as the benzene or toluene hydrogenation process mentioned above, is not suitable for use in the treatment of distillate fractions containing mixtures of aromatic hydrocarbons. This is due in part to variability in the reactivity of different aromatic hydrocarbons. For example, of the substituted benzene compounds contained in the aromatic feed, hydrogenation of mesitylene requires more than 2 times as long as hydrogenation of benzene. In practice, this means that under the conditions under which benzene is completely hydrogenated, considerable amounts of heavier aromatic compounds remain unsaturated.

Another significant problem with the hydrogenation of petroleum distillates and other hydrocarbon feeds is the impurities they contain. These impurities are derived from the original crude oil or are formed during the process, for example during the reforming. Of particular importance are sulfur and halogen compounds, which poison the catalyst and shorten its effective life.

It follows from the above that a modern refinery producing several oil distillates for use as light and heavy fuel and as feedstocks for petrochemicals and plastics must have several different hydrogenation units.

• · * ♦ · · • · ·

The object of the present invention is to eliminate the problems which have arisen in the past and to provide a completely new process for the dearomatization of petroleum distillates and similar hydrocarbon feeds containing various aromatic hydrocarbon mixtures. In particular, it is an object of the invention to provide a multi-purpose hydrogenation process for the catalytic dearomatization of oil distillates, according to which the catalytic removal of aromatics from solvents and middle distillates and the removal of aromatics and / or benzene from petrol fractions can be carried out in the same basic process configuration, preferably in the same process unit.

According to a preferred embodiment, the above objects are achieved by a process, the basic concept of which is that at least two, and preferably all three, of the following oil distillates a to c are hydrogenated in the same hydrogenation unit:

(a) a solution fraction with a boiling range of 60 to 270 ° C

(b) middle distillate boiling in the range of approximately 180 ° C to 320 ° C; and (c) straight-run naphtha fraction, reformate, cracked light or heavy naphtha fraction boiling in the range of approximately 20 ° C to 190 ° C.

According to the invention, each feedstock is contacted with hydrogen in a ratio of 100-500 Nm3 of hydrogen / m3 of hydrocarbon contained in the feedstock. The hydrogenation of each feedstock is further carried out in a pressure range of 15-55 bar, a temperature range of 50-250 ° C and a volume rate, calculated as LHSV, of 0.5-101 / h. The temperature rise, ΔΤ, in the reaction zone during dearomatization is kept below 60 ° C.

The invention is further based on the fact that it is possible to process different oil fractions containing a large number of aromatics with different reactivity towards hydrogen in the presence of a catalyst if the sulfur content is reduced to less than 100 ppm, preferably less than 50 ppm and particularly preferably less than 5 ppm by weight. , before hydrogenation. Correspondingly, the concentrations of halides *: 1 1: should be less than 50 ppm, preferably less than 25 ppm and especially less than 2.5 ppm.

More specifically, the process according to the invention for dearomatizing oil distillates is characterized by what is stated in the characterizing part of claim 1.

• · · • · · · · 1 · • · »·. According to another embodiment, the object of the invention is achieved by a modified process based on reversing the steps of the oil feedstock to be fed to the new concept and its hydrogenation. Thus, the second embodiment is characterized in that in the process "" the aromatic oil-containing feedstock is fed to a hydrogenation unit, the feedstock is contacted with hydrogen in the presence of a catalyst to hydrogenate the feed and the aromatics contained therein to produce a dearomatized product, dearomatized from the process hydrogenation unit and the dearomatized product is divided into at least two of the following dearomatized oil distillates (a) to (c): 5 a) a solvent fraction with a boiling range of 60-270 ° C and containing less than 2% by volume of aromatics, b) a middle distillate with a boiling range is 180-320 ° C and contains less than 20% by volume of aromatics, and c) a petrol fraction having a boiling range of 20-190 ° C and containing less than 10% by volume of aromatics.

More specifically, the second embodiment is characterized by what is set forth in the characterizing part of claim 19.

The general concept of the process according to the invention is suitable, as such or with minor modifications, for the preparation of a wide variety of dearomatized and / or benzene-free petroleum distillate products with a boiling point range of 20-300 ° C.

Hydrogenation of aromatic solvents and middle distillates, straight run gasoline fraction, ..: 20 reformate, cracked light and heavy gasoline fraction aromatics and / or benzene; the hydrogenation can all be performed in the same multi-purpose process unit.

The dearomatization process of the above-mentioned first embodiment is suitable for the production of different aroma-free and / or benzene-free products in the same process unit and is particularly advantageous for smaller refineries. Particular advantages are obtained from the fact that only one process unit is needed instead of two or three units, which leads to savings in investment costs. The process also makes it possible to produce a number of different "·. · Products periodically, and even when annual production volumes are small.

The dearomatization process according to the second embodiment mentioned above, in which the oil feedstock is first fed to a hydrogenation reactor, after which the treated material is divided into dearomatized oil products (solvents, gasoline fraction and diesel) in a distillation column, is preferred because many different dearomatics are preferred. in the unit at the same time. Particularly advantageous in this embodiment is that since only one unit needs to be built for the refinery, large savings in investment costs are achieved; many different products can be produced at a constant production rate.5 and there is no need for storage of intermediates.

Figure 1 is a simplified diagram of a dearomatization process comprising two hydrogenation reactors and a product stripper according to the present invention;

Figure 2 shows a process diagram of a modified process particularly suitable for the hydrogenation of light feedstocks, and

Figure 3 shows a diagram of an alternative dearomatization process consisting of a hydrogenation reactor and a distillation section for fractionating the dearomatous feed into suitable oil distillate fractions.

In the context of the present invention, the term "aromatics" means all aromatic compounds present in the feed oil-based hydrocarbon feedstock. Of particular interest to this invention are benzene and the following typical straight-run or cracked distillates of oil (primarily from the FCC; fluid catalytic cracking unit), alkyl-substituted aromatics such as toluene, tetramethylbenzene, diethylmethylbenzene, pentylbenzene, pentylbenzene.

•; * ·: The terms “hydrogenation” and “dearomatization” are used interchangeably to describe the process in which the unsaturated bonds of aromatic compounds are saturated by reacting with hydrogen (H2).

25 • · · '; 'Solvents' are aromatic hydrocarbon feedstocks with a typical boiling point * \ '(hereinafter also abbreviated as kp) in the range of 60 to 270 ° C, a specific gravity of 0,680 to 8,825:' (kg / dm3), and containing 20 to 30% before hydrogenation. vol .-% aromatics.

. j .: 30 “Middle distillates” shall be fed to an aromatic feedstock commonly used, for example: ... as a fuel for diesel engines, with a boiling range of 180 to 320 ° C. It has a specific gravity of about 0.820 and contains about 20-30% by volume of aromatics before hydrogenation.

6 100535

The term “gasoline fraction” is used from a feedable aromatic feedstock that can be used as fuel for automotive engines. It consists of one or more fractions, such as a straight-run gasoline fraction, a reformate, a cracked light gasoline fraction and / or a cracked heavy gasoline fraction. Gasoline fraction kp. the range is 20-190 ° C, the specific gravity is 0.600-0.830 and the aromatic content is 0-30% by volume (before hydrogenation).

By means of the present invention, it is possible to reduce the aromatic content of the solvent to 0.1-2% by volume, 1-20% by volume of middle distillate products and 0-5% by volume of gasoline fraction.

As previously mentioned, each feedstock is hydrogenated under the following conditions: hydrogen / hydrocarbon ratio: 100-500 Nm3 hydrogen / m3 feed hydrocarbons, pressure: 15-55 bar, temperature: 50-250 ° C, volume flow (LHSV): 0.5 -101 / h 15

Reaction conditions of each starting material from the above ranges. Thus, the temperature range (feed temperature) for the hydrogenation of the solvent starting material, for example, is 130-160 ° C, while the temperature range suitable for the hydrogenation of light solvents, hexane feed and reformate is. . below 140 ° C, typically 50-135 ° C. For a total feed, e.g. a feed consisting of hydrocarbons with a wider boiling range of 20 and containing middle distillates, a temperature range of 130-200 ° C is suitable.

«·: The hydrogen-carbon hydrogen ratio varies depending on the amount and quality of aromatics present. For example, in the case of benzene, with a benzene content of 0.1 to 10% by weight, the feed is hydrogenated with a hydrogen-benzene ratio of 3.2: 1 to 6: 1 to produce a hexane fraction containing less than 0.1% by weight of benzene. .

• · · • · · •

It is advantageous to reduce the sulfur content of the oil and hydrogen feed before hydrogenation. The sulfur content of all feedstocks fed to the dearomatization unit should be less than 100: 30 ppm by weight, preferably less than 50 ppm by weight and particularly preferably less than 10 ppm by weight or even less than 5 ppm by weight. Correspondingly, the preferred maximum concentration limits for halogen compounds (chlorides, fluorides, bromides and the like elemental halogens) in the oil and hydrogen feed are 50% of the sulfur content limits, e.g. 50.25, 5 and 2 ppm by weight.

In some cases, depending on the origin of the crude oil, the sulfur concentrations meet the requirements of 5 above without further processing steps. However, it is often necessary to remove sulfur and halogen compounds from the feedstock.

Methods known per se can be used to reduce sulfur and halogen compounds. Catalytic desulfurization is the most commonly used, but inert gas stripping, sodium hydroxide washing, adsorption on an adsorbent such as clay or molecular sieves, or solvent treatment are also possible. Halogen compounds can be removed by a similar method.

According to the invention, the aromatics of the hydrocarbon feed are hydrogenated in the temperature range from 50 to 250 ° C in the presence of an active catalyst. Suitable catalysts are heterogeneous catalysts consisting of a Group VIII metal on a solid support. Preferred metals are nickel, iron, platinum and palladium, and preferred carriers include silica, alumina, magnesium oxide, zirconia, titanium oxide and natural and synthetic aluminum and magnesium silicates. It is particularly preferred to carry out the hydrogenation on a nickel catalyst, for example a catalyst consisting of nickel and an inorganic metal oxide support. . tin, in the presence. Such a catalyst may contain 0.1 to 70% by weight, preferably 1 to 50% by weight, of nickel on an alumina or silica support.

t! :: • · · J .:: According to a preferred embodiment, the dearomatization is carried out in the presence of a nickel catalyst 25 suitable for the hydrogenation of edible oils. In the context of the present invention: it has surprisingly been found that such catalysts, which are known to be highly selective for fatty acids containing conjugated or non-conjugated bonds, are quite capable of catalyzing a wide range of oil distillates. hydrogenation of hydrocarbons. In addition, these highly active catalysts have been shown to be reasonably resistant to poisoning by sulfur and halogen compounds contained in the oil feed. A third particularly significant advantage associated with the present applications is that by using catalysts for the food industry, the dearomatization temperature can be kept below 250 to 100535 ° C, preferably in the range of 50 to 240 ° C for all oil feeds. At such low temperatures, the formation of carbon scale is greatly reduced. In particular, as the following examples show, good results are obtained with HTC 400 (supplied by Crossfield) type catalysts consisting of nickel on an alumina support and having a typical nickel content of 10-30% by weight.

For the properties of nickel catalysts used in the food industry, reference is made to Heterogeneous Catalysis; Principles and Applications, Bond, G.C., pp. 75-78, Oxford University Press, 1974.

10

Hydrogenation is an exothermic reaction. Therefore, the temperature tends to rise in the hydrogenation reactor. To avoid excessive temperatures and excessive temperature differences above the catalyst beds, which have a devastating effect on the life of the catalyst, the temperature is kept below 60 ° C during the reaction.

15

There are several different ways to create adequate temperature control. Thus, one possibility is to dilute the feed with an outflow reflux to reduce the aromatic content of the feed. In this embodiment, the effluent from the reaction zone is cooled and partially recycled back to the reaction zone. The ratio of recycled outflow to fresh feed is 0.1-5, preferably 0.2-3. and particularly preferably 0.5-3. By circulating the liquid, the hydrogenation temperature rise in the reactor can be limited to less than ·: ··· 60 ° C. Typically, the temperature rise through the reactor is about 20-60 ° C. A maximum aromatic concentration of about 16% by volume or a maximum benzene concentration of 5% by weight (or even 4.3% by weight) can be achieved in the dilution feed.

25 :: The cooled effluent or liquid reaction product is recycled from the gas separation tank or • · · • # * ’· [from the bottom of the distillation column to the reactor inlet.

· T

Based on the above, a particularly preferred embodiment of the invention consists of: a combination of 30 steps: reducing the sulfur content of the feed to less than 5 ppm, hydrogenation in the presence of nickel catalyst at 50-250 ° C and 15-55 9 100535 bar, cooling the effluent from the reaction zone and recycle to maintain a temperature rise of no more than 60 ° C.

5 Another suitable way to limit the rise in reaction temperature is with hydrogen cooling. In this embodiment, hydrogen is separated from the reactor effluent and fed to the reactor between the catalyst beds. Hydrogen cooling can also be used to lower the temperature between two reactors connected in series, and fresh hydrogen can also be used for hydrogen cooling.

10

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

Hydrogen gas recycling represents a preferred embodiment. In this case, the hydrogen-containing gas is separated from the liquid product in a product separator and recycled to the reactor feed.

In order to separate the impurities from the recycled gas, a part is led to the fuel gas 20 or to the hydrogen network. The ratio of recycled hydrogen to fresh hydrogen is about 0.1-10. In general, fresh hydrogen is somewhat more hydrogen-containing: fresh hydrogen contains, e.g., 89-92 mol.% H2 and ....: recycled gas, e.g., 85-90 mol.% H2.

«1 <» # ♦ · • · · i *: ': The process according to the invention is carried out in a hydrogenation unit consisting of at least 25 reactors, preferably two reactors arranged in series. It is also possible to arrange the reactors in parallel. With respect to the basic concept of the invention, satisfactory results • * · *. * * is already achieved with one reactor. However, due to the potential poisoning of the catalyst, it is preferable to use two or more reactors to ensure continuous operation. In this case, when the catalyst in the first hydrogenation reactor is deactivated, it is possible to use: 30 second reactors, and the production continues without interruption. The cost of the reaction equipment, combined with the necessary instrumentation, is negligible compared to the cost of production losses. In addition, the use of two reactors reduces the amount of catalyst required, as only the reactor in operation needs to be filled with catalyst. If one large reactor is used, most of the catalyst is extra in the early stages of operation and therefore incurs additional costs. By using two reactors connected in series, it is generally possible to limit the aromatic content of the products more strictly than with one reactor. The reactors can be any suitable rectifiers for contacting the liquid and the gas, with bed bed reactors being particularly preferred.

The first reactor effluent is cooled in a heat exchanger to the first reactor feed temperature before it is fed to the second reactor or before a portion thereof is recycled to the reactor inlet.

The liquid product obtained after hydrogenation is fed to the product stabilization section, where the hydrocarbons and hydrogen are separated in a distillation column or steam stripper.

15 Embodiments of the process are discussed in more detail below with reference to the accompanying drawings:

From the embodiment shown in Fig. 1, it can be seen that the process configuration used for dearo- maturing the solvent and middle distillate consists of two 20 bed effluent reactors 1,2 and a distillation column 3 for stabilizing the dearomatized product.

The oil distillate feed is led to the first bed bed reactor 1 by means of a pump 5. Before the i · * ....: reactor, the liquid stream is connected to the fresh hydrogen and the recycled hydrogen feed and heated in the heat exchanger 4 to the feed temperature.

• »« 1 · · f i «• 25 The temperature rise in the reactor is controlled by a liquid circuit with a cooled liquid product; · .: recycled from the gas separator 8 or from the bottom of the distillation column 3 to the reactor inlet.

(teaspoon • * * *. * * The liquid recirculation rate is 0.5-3 times the fresh feed rate.

»

The illustrated process configuration uses two reactors 1,2 in Saija. This embodiment makes it possible to achieve temperature control of the reactors by liquid recirculation, which is more accurate than with other ice systems. However, as mentioned above, the two reactors can also be in parallel, allowing the catalyst to be changed without shutting down the entire 11,100,535 units.

The effluent temperature of the reactor 1 is cooled to 130 ° C in a heat exchanger 4. The liquid to be recycled is separated from the liquid effluent in a gas / liquid separator 6 and 5 is introduced into the reactor 1 for flow.

The second reactor 2 operates in the same way as the first. The hydrogen-containing gas is separated from the liquid outflow in the gas separator 8 and recycled to the inlet of the reactor 1.

10

Reference numeral 9 denotes a compressor in the gas line. To purify the contaminants, a portion of the recycle gas is introduced into the fuel gas or hydrogen network, as shown in Figure 1.

Generally, the liquid effluent from the gas separator 8 is not recycled to either reactor.

The liquid effluent from the second reactor is fed to the product stabilization section 3. In the stabilization section, the process conditions are sufficient to produce products substantially free of hydrogen and light hydrocarbons. In the case of a solvent / middle distillate feed, the excess of the distillation column-20 nin 3 is divided after the condenser 10 in the transfer tank 11 into non-condensable gases used as fuel gases and light liquid hydrocarbons which

«· I

. ,,,: can be used as a gasoline fraction. The dehydrated solvent or middle distillate fraction is removed as a base product.

M4 «I · | · · Ϊ * 4 Figure 2 shows a similar but simplified process configuration, which is particularly suitable for hydrogenation of gasoline fraction feeds. The process consists of a single hydrogenation reactor and stabilization sections 13, 19 and 20 as described above.

• tl * <i t i This embodiment is operated in the same way as above: the oil feed is fed to the hydrogenation reac-; 30 to the market 12 by means of a pump 15. The liquid stream is combined with fresh hydrogen and recycled hydrogen feed and then heated in a heat exchanger 14 to the feed temperature. The reactor effluent is cooled in a heat exchanger 16, and the gas phase is separated from the liquid phase in a gas 12 100535 separation apparatus 17. A portion of both the cooled liquid flow and the hydrogen gas is recycled to the reactor 12 inlet.

The dearomatized liquid reaction product is passed to a distillation column 13, where hydrogen gas and 5 light hydrocarbons are separated from a dearomatized product consisting, for example, of cyclohexane or methylcyclohexane or a gasoline fraction.

Figure 3 shows a process diagram of an alternative embodiment in which the oil feed is first dearomatized in a hydrogenation section similar to that illustrated in Figure 1 and then distilled into different fractions in the distillation section.

Numbers 21 and 22 refer to two hydrogenation reactors connected to Saija. Numeral 23 indicates a distillation column used to distill the feed from the dearomatization section.

As in the embodiment of Figure 1, the oil feedstock is fed to the first dearomatization-15 reactor by means of a pump 25. Prior to the reactor, the liquid stream is combined with the hydrogen feeds and heated in a heat exchanger 24 to a suitable feed temperature. The temperature rise in the reactor is controlled by the circulation of liquid from the product separator 26 to the inlet of the reactor 21. The next reactor 22 is operated in the same manner as the first. The hydrogen-containing gas separated from the liquid outflow in the gas separator 28 is recycled! 20 to the inflow of reactor 21 through a compressor 29 in the gas line.

»»

The liquid effluent from the gas separator 29 is fed to the distillation column 23 by the heat exchanger 9 fj /. 30 through. The distillation section (numbers 23-40) consists of a distillation column 23, side rippers 32-34

Hl · '· and transfer tank 41. The distillation column is equipped with a bottom boiler (heater) 35 »25 and an intermediate boiler 31. Side strippers 32-34 consist of gas / liquid separators in which: s |: the condensed liquid stream is separated from the gas, which is then recycled to the distillation tank - • · V: yes. Each liquid stream is cooled in heat exchangers 37-39 to form distillation fractions A-C having different boiling point ranges. The bottom product of the distillation column is: the heaviest fractions, which are cooled in a separate heat exchanger 36. The over-. \ 30 The flow is cooled in an air cooler 40. In the bypass tank 41, the light hydrocarbons and water are separated.

13 100535

If the dearomatized feed of the distillation section contains a hydrocarbon fraction with a boiling range of about 20-300 ° C, the following products can be recovered: middle distillates as bottom products, three solvent products as by-products A-C, and gasoline fraction and fuel gas as overflow.

The embodiment of Figure 3 is described in more detail in Example 9. Example 1 10 Several dearomatized solvents and diesel were produced in the same process unit with two bed bed reactors in series. The process configuration corresponds to that shown in Figure 1. The test runs were performed at a capacity of 20 t / h.

To prepare the solvents, a feed containing 22% by volume of aromatics was hydrogenated in a dearomatization unit. The solvent feed density was 780 kg / m3 at 15 ° C and the distillation range was ASTM D86 distillation, v / v%.

IBP 150 ° C

: 20 10 161 ° C

30-169 ° C

«·. !!!: 50 173 ° c • ·

70-179 ° C

• · ·

90-190 ° C

25 95 193 ° C

: FBP 197 ° C

• · · tn • · · • · ·:: ’: The fresh hydrogen feed used, which had a molecular weight of 4.8, contained. 30 hydrogen 90.9 mol% methane 3.2 mol% ethane 3.0 mol% 14,100535 propane 1.6 mol% butane 0.6 mol% pentane 0.7 mol%

5 Solvent feed containing less than 5 ppm (w / w) sulfur was heated to 130 ° C

hydrogen with a mixture of fresh hydrogen and recycled hydrogen. The heated biphasic stream was fed to the first reactor of the dearomatization unit, where the aromatics of the hydrocarbon feed were hydrogenated in the presence of an active HTC-400 type nickel catalyst. The reactor pressure was 40 bar. The temperature rise in the reactor was 30 ° C. The temperature rise in reactor 10 was controlled by liquid recirculation consisting of recirculating the cooled liquid effluent from the gas / liquid separator to the first reactor. The outlet of the first reactor was cooled to 130 ° C in a heat exchanger. The liquid stream was separated in a high pressure separator and recycled back to the first reactor inlet.

After the separator, the hydrocarbon stream and the hydrogen-containing gas were mixed and fed to the second reactor of the unit, where the hydrogenation reactions were completed in the presence of an active HTC-400 type nickel catalyst. The reactor pressure was 36 bar and the temperature rise in the reactor was 8 ° C.

The reactor outlet was cooled to 40 ° C in a heat exchanger. A hydrogen-rich gas containing 87 mol% of hydrogen was separated from the liquid effluent in a gas separator and recycled to the first reactor feed. Fluid flow was directed to stabilization, where light com. . ·. the components were separated from the dearomatized product. The stabilization section consisted of a distillation column, • · ·. ’: *; boiler, condenser and transfer tank. In the stabilization column, hydrogen and light hydrocarbons were distilled from the dearomatized solvent product at a pressure of 200 kPa. Distillation: the feed temperature was 160 ° C. There was thus obtained a solvent product having a distillation range of 150 to 200 ° C, a specific gravity of 0.780 and an aromatic content of 0.5% by volume or less.

Example 2: 30 ': A second test run was performed in the same unit with a heavier solvent feed. with a specific gravity of 0.810. The feed used contained aromatics 25% by volume and sulfur 5 ppm (w / w).

15 100535

The distillation range was: ASTM D86 distillation, v / v%

5 IBP 198 ° C

10,209 ° C

30 212 ° C

50 215 ° C

70-217 ° C

10 90 230 ° C

95 239 ° C

FBP 248 ° C

The solvent feed and a mixture of hydrogen and recycled hydrogen were heated to 130 ° C. The heated biphasic mixture was fed to the first reactor containing an HTC-400 type nickel catalyst. The reactor pressure was 40 bar. The temperature rise in the reactor was 32 ° C and was controlled by liquid recirculation.

The effluent from the first reactor was cooled to 132 ° C in a heat exchanger, and the liquid stream was separated in a gas separator and recycled back to the inlet to the first reactor.

• · · • ·. The hydrocarbons and the hydrogen-containing gas were mixed and fed to a second reactor filled with a nickel catalyst of the HTC-400 type. The reactor pressure was 36 bar and the temperature rise in the 25 reactors was 9 ° C.

• · · • · · • · :: The effluent of the second reactor was cooled to 40 ° C in a heat exchanger. The hydrogen-containing gas containing 87 mol% hydrogen was separated from the liquid stream in the gas separator and recycled back to the first reactor inlet stream. The liquid stream was fed. . ·· to 30 stabilization sections where the light components were removed from the dearomatized solvent product at a pressure of 200 kPa. The distillation feed temperature was 200 ° C. The aromatic content of the product was 0.5% by volume or less.

Example 3 16 100535

The third test run was performed in the same unit with a feed containing middle distillates. The feed used contained 24% v / v aromatics and 5 ppm sulfur (w / w). The distillation range of the feed with a specific gravity of 5 0.820 was as follows: TBP distillation, vol.%

IBP 129 ° C

10 5 178 ° C

10 191 ° C

30 212 ° C

50-230 ° C

70 250 ° C

15 90 270 ° C

95 279 ° C

FBP 320 ° C

The middle distillate feed was heated to 170 ° C with a mixture of hydrogen containing 20 fresh hydrogen and recycled hydrogen. The heated two-phase stream was fed to the first reactor containing an HTC-400-type nickel catalyst. The reactor pressure was 40 bar. The temperature rise in the reactor controlled by the liquid circulation was 34 ° C.

• · • · ·

The effluent from the first reactor was cooled to 174 ° C in a heat exchanger, and the liquid to be recycled was separated in a gas separator and recycled back to the inlet to the first reactor. Hydrocarbons and hydrogen-containing gas were mixed and fed to another reactor containing an HTC-400-type nickel catalyst. The reactor pressure was 36: bar and the temperature rise in the reactor was 9 ° C.

. The reactor effluent was cooled to 40 ° C in a heat exchanger. The hydrogen-containing gas containing 87 1: mol% hydrogen was separated from the liquid stream in the gas separator and recycled back to the first reactor inlet. The liquid flow was fed to the stabilizing section, i; 17 100535 wherein light components such as hydrogen, C 1 -C 5 hydrocarbons and gasoline fraction components were removed from the dearomatized middle distillate fraction. The distillation temperature was 240 ° C. The aromatic content of the product was 5% by volume.

5 Example 4

A test run to hydrogenate benzene from the hexane feed was performed in a unit of the type shown in Figure 1 without recycle gas and with only one reactor. Benzene was hydrogenated to cyclohexane with an HTC-400 type nickel catalyst in a bed bed reactor. The sulfur-10 content of the feed was less than 5 ppm (weight).

The hexane feed used contained: isopentane 0.5 wt% 15 n-pentane 3.0 wt% cyclopentane 4.0 wt% methylcyclopentane 9.0 wt% cyclohexane 7.0 wt% benzene 8.6 wt.% 20 n-hexane 24.0 wt.% other C6 hydrocarbons 39.6 wt.% C7 hydrocarbons 4.3 wt.% • · · «· · •« · • 1 1 I · · • · · • · 1 • · · i: · · • · · 18 100535 Used fresh hydrogen, molar weight 5.24, contained: hydrogen 90.8 mol.% 5 methane 2.8 mol.% Ethane 2 .5 mol% propane 2.0 mol% isobutane 0.5 mol% n-butane 0.4 mol% 10 pentane 0.4 mol% hexane 0.6 mol%

The hexane feed was heated to 90 ° C with fresh hydrogen in a heat exchanger and fed to the reactor. The reactor pressure was 2100 kPa. The hydrogenation of benzene to cyclohexane is an exothermic reaction and the temperature rise in the reactor was 31 ° C. The reactor temperature rise was controlled by liquid recirculation, in which the hydrogenated liquid flow was recycled back to the reactor inlet flow.

The reactor effluent was cooled in a heat exchanger. The hydrogen gas was separated from the 20 liquid streams in a gas separator. A hydrogen-containing gas containing 87.8 mol% of hydrogen was fed into the hydrogen network. The liquid to be recycled in the gas separator was separated from the hydrogenated liquid stream and t 1 .... was fed to the reactor inlet.

•: • • i »• t · • · ·.: 1: 1: The hydrogenated product was fed to the product stabilization section, where light hydrocarbons were distilled off from 25 hexane products at a pressure of 600 kPa. The benzene content of the hexane product was 0.1% by weight.

• · · «· · • · i: · 1 • · · • ·

Example 5 19 100535

A second test run for the hydrogenation of benzene was carried out with a reformate feed having a b.p. range of 55-222 ° C, a specific gravity of 0.831 and a benzene content of 4.0% by weight.

5

The reformate feed was heated to 90 ° C in a heat exchanger and fed to a ladle reactor filled with HTC-400 type nickel catalyst. The reactor pressure was 2100 kPa. The temperature rise in the reactor was 19 ° C and was controlled by liquid recirculation.

The reactor effluent was cooled in a heat exchanger. The hydrogen gas was separated from the product in a gas separator. The hydrogen-containing gas was fed into the hydrogen network, and the liquid was recycled to the reactor inlet stream.

The hydrogenated liquid stream was fed to a stabilizing section where light hydrocarbons were distilled off from the 15 reformates at a pressure of 600 kPa. The benzene content of the product was 0.1% by weight or less.

Example 6

In the third benzene hydrogenation test run, light solvents containing 3.9 to 20% by weight of benzene were used. The feed range was 88-105 ° C and the specific gravity was 0.720.

, **. j The solvent feed was heated to 110 ° C with fresh hydrogen and fed to HTC-400- • ·. . ·. to a bed bed reactor containing a type of nickel catalyst. The reactor pressure was 2100 • «·. *: * kPa. The temperature rise in the reactor controlled by the liquid circuit was 17 ° C.

»25: The reactor effluent was cooled in a heat exchanger. Hydrogen gas was separated • «·

Vs the product in the gas separator. The hydrogen-containing gas was fed into the hydrogen network. The liquid to be recycled in the separator was separated from the hydrogenated liquid stream and fed to the reactor inlet.

The hydrogenated product was fed to a stabilizing section where light hydrocarbons were distilled from the solvent product at a pressure of 600 kPa. The benzene content of the product was 0.01% by weight.

30

Example 7 20 100535

A fourth benzene hydrogenation test run was performed with light solvents containing 5.6 wt% benzene. The feed used had a density of 0.680 and a distillation range of 65-75 ° C.

The solvent feed was heated to 110 ° C with fresh hydrogen in a heat exchanger. The heated mixture was fed to a casting bed reactor filled with HTC-400 type nickel catalyst. The reactor pressure was 2100 kPa. The temperature rise in the reactor was 28 ° C and was controlled by liquid circulation.

The reactor effluent was cooled in a heat exchanger. The hydrogen-containing gas was separated from the liquid stream in a gas separator. The liquid circuit was separated from the hydrogenated liquid effluent and fed to the reactor inflow.

15

The hydrogenated product was fed to a stabilization section where light hydrocarbons were distilled from the solvent product at a pressure of 600 kPa. The benzene content of the light solvent was 0.01% by weight.

Example 8 20

Toluene hydrogenation was studied by simulating a bed bed reactor model. Simulation Model # '”; presented an industrial-scale reactor filled with a hydrogenation catalyst. The diameter of the reactor was,, 1.5 m. The properties required of the catalyst for mass transfer and loss calculations are shown in • 1 · Table 1.

• i * 25:: ’: Table 1. Properties of catalyst particles: Particle diameter 1 mm

Particle length 3 mm

Bed pore size 0.40: 30 Bulk density_830 kg / m3_

The liquid feed to the simulated reactor was 100 t / h and consisted of toluene (5 wt%), methyl cyclo-21 100535 salohexane and ten inert hydrocarbon compounds. The gas feed contained 95 mol% hydrogen and a small amount of light hydrocarbons. The gas feed rate was 1.0 t / h. The liquid and gas supply temperature was 110 ° C. The operating pressure was 3100 kPa.

According to the results obtained with the reactor model, 85% conversion of toluene was achieved with a reactor 2 m high, and a reactor 4 m high was sufficient for 99% conversion of toluene.

The temperature rise was about 38 ° C. The calculated pressure drop in the reactor was 80 kPa.

Example 9 10

The following example relates to the dearomatization process shown in Figure 3, in which the oil feed is first fed to a bed bed reactor, after which the liquid stream is divided into dearomatized oil distillates; solvent, gasoline and middle distillate fractions. The unit's feed capacity was 50 t / h.

15 The characteristics of the oil feed were as follows: density at 15 ° C at 808 kg / m3 and distillation range (ASTM D86 distillation) at 145-320 ° C. The feed contained 20% v / v aromatics and 5 ppm (w / w) sulfur.

Used fresh hydrogen, mol. weight 4.1, contained: 20 hydrogen 93.0 mol% methane 3.1 mol% ethane 1.5 mol% I a t i ';'. propane 1.3 mol .-% 25 butane 0.8 mol .-% ϊ pentane 0.3 mol .-% 41 ·

• · I

• · · • The oil feed was heated to 150 ° C with a mixture of hydrogen and fresh hydrogen. The heated biphasic stream was fed to a reactor with aromatics,. 30 was hydrogenated from the hydrocarbon feed in the presence of an active HTC-400 type nickel catalyst.

; The reactor pressure was 40 bar. The temperature rise in the reactor was 34 ° C and was controlled by a liquid circuit in which the cooled liquid effluent was recycled from the separator to the reactor inlet.

The reactor effluent was cooled to 40 ° C in a heat exchanger. The hydrogen-containing gas containing 87 mol% hydrogen was separated from the liquid effluent in a gas separator and recycled back to the reactor inflow. The liquid stream was fed to a distillation section where the liquid stream was divided into oil fractions.

In the distillation section, the liquid feed was heated to 175 ° C in a heat exchanger and fed to a distillation column in which the liquid stream was distilled into dearomatized gasoline, solvent and middle distillates at a pressure of 160 kPa. The column had a diameter of 3 m and a height of 50 m. The column had 72 distillation bottoms and a cooled 2.3 MW pumping circuit from bottom No. 38. The feed was directed to bottom No. 68.

The column further contained three side effluents from bottoms Nos. 58, 38 and 18, from which three different solvent products A, B and C with different b.p. regions were removed. Each solvent was stabilized in 8 bottom steam strippers. The side rippers of solvents A and B had a diameter of 1.0 m and a height of 6.0 m. The side rappers of solvent C had a diameter of 1.1 m and a height of 6.0 m. The distillation products are specified in Table 2.

The bottom temperature of the column was 310 ° C and the bottom of the column was boiled with a 10 MW boiler. The bottom product of the column was middle distillate.

I *> «· t The top temperature of the column was 120 ° C and the condenser capacity of the top was 10 MW. The product of the upper · * · I ′ olir was unstabilized gasoline, which was fed to a stabilizing section where light hydrocarbons were removed from the dearomatized gasoline fraction. Alternatively, the reaction product can be passed first to the stabilization section and then to the distillation column to separate the oil distillates.

tn • € * 4 ψ P «23 100535

Table 2.

Product Speed, Aromatics Qty range Characteristic Flash point

ton / h vol .-% ° C weight point ° C

Solvent A 1.7 <0.5 213-270 0.823 91 5 Solvent B 11.0 <0.5 188-246 0.810 72

Solvent C__16 ^ 0__ £ 0 ^ __ 170-221 0.797__54_

Unstabilized benzene 9.4 <0.2 -184 0.777 rail fraction __ "middle distillate 11.8 <1.0 241-328 0.845 10 • · ·« · • · • · · • · · · * · · • «· • · · • · · • · · · «1» • · · ·

Claims (26)

A process for the dearomatization of oil distillates, wherein A) an oil distillate feed containing aromatic compounds is fed to a hydrogenation unit, B) the feed is reacted with hydrogen in a reaction zone to hydrate the feed and the aromatic compounds in the feed to a dear flavored product, and C) the dearomatized product is recovered from the hydrogenation unit, characterized by D) atm instone two of the following oil distillates a) - c) hydrogenated in the same hydrogenation unit: a) a solvent fraction having a boiling point range of 60 - 270 ° C, b) an average distillate with a boiling point range of 180 - 320 ° C, and c) a primary distillation gasoline reformer, a reformed, a light gasoline fraction or a cracked heavy gasoline fraction, all of which exhibit boiling points in the range of 20 - 190 ° C, feed is reacted with hydrogen utilizing a ratio of 100 to 1 500 Nm3 of hydrogen / m3 of hydrocarbons in the feed, and the hydrogenation of each feed is carried out at a pressure of 15 - 55 bar, at a temperature of 50 - 250 ° C and with a volume flow of 0.5 - 10 1 / h, calculated as an LHSV. F) the temperature rise within the reaction range is limited to less than 60 25 ° C by cooling the outlet from the reaction zone and by recirculating a portion of the same to the reaction zone. :. ·. 2. Process according to claim 1, characterized in that all three oil distillates a - c,, φ · are hydrogenated in the same hydrogenation unit. . * 30 • * '·. 3. Process according to claim 1 or 2, characterized in that a solvent whose distillation range is between 60 and 270 ° C and which contains not more than 30% aromatic compounds is hydrogenated to produce a hydrated distillate containing less than 2%. % aromatic compounds. 4. A process according to any one of claims 1 to 3, characterized in that the average distillate fraction is hydrogenated, the distillation range of which is between 180 and 320 ° C and containing up to 30% aromatic compounds, to produce a hydrated distillate containing less than 20 vol. -% aromatic compounds. Process according to any one of claims 1-4, characterized in that a primary distillation gasoline, a reformate, a cracked light gasoline fraction or a cracked heavy gasoline fraction having a distillation range between 20 and 190 ° C is hydrated. up to 30% aromatic compounds, to prepare a gasoline fraction containing less than 5 vol.% aromatic compounds. 15 6. Process according to claim 5, characterized in that the benzene content of the feed is 0.1 - 10% by weight and the feed is hydrogenated with a hydrogen-benzene ratio of 3.2: 1 -6: 1 to obtain a hexane feed containing less than 0 , 1 wt.% Benzene. Method according to claim 1, characterized by the ratio between the circulated outlet and the feed rate being 0.1 - 5, preferably 0.2 - 3. Preferably. 8. A process according to claim 1, characterized in that the hydrocarbon inputs to the de-aromatizer have a sulfur content less than 100 ppm (weight), preferably less than 50 ppm (weight) and preferably less than 5 ppm (weight), and the halide content. Is less than 50 ppm, preferably less than 25 ppm and more preferably less than 2.5 ppm. A process according to any of the preceding claims, characterized in that the hydrogenation takes place in the presence of a nickel catalyst. . '30. Process according to any of the preceding claims, wherein the hydrogenation takes place in the presence of a heterogeneous nickel catalyst, characterized in that it contains elemental nickel on an inorganic metal oxide support. 11. Process according to claim 10, characterized in that the hydrogenation catalyst 5 contains 0.1-70% by weight, preferably 1-50% by weight nickel on an alumina or silica support. Process according to any of the preceding claims, characterized in that the hydrogenation is carried out in a hydrogenation unit having at least one drainage bed reactor. 10 Method according to claim 12, characterized in that two drainage bed reactors are connected in series. 14. A method according to claim 13, characterized in that the liquid flow from the first reactor is cooled in a heat exchanger and measured to the second reactor. 15. A process according to claim 14, characterized in that hydrogen is separated from the outlet from the last reactor and circulated to the inlet to the first reactor. itii; 16. Process according to any of the preceding claims, characterized in that the hydrogenated distillate is stabilized to remove light components from the flavored products. And (C) the dearomatized product is recovered from the hydrogenation unit, characterized by D) using hydrogen utilizing a ratio of 100 to 500 Nm3 of hydrogen / m3 of hydrocarbons in the feed, and the hydration of each feed is carried out at a pressure. of 15 - 55 bar, at a temperature of 50 - 250 ° C and with a volume flow of 0.5 - 10 1 / h, calculated as an LHSV value, and E) the temperature rise within the reaction range is limited to less than 60 ° C by cooling the outlet from the reaction zone and by recirculating a portion of the same to the reaction zone, and F) the product is divided into at least two of the following dearomatized oil fractions a - c: a) a solvent fraction having a boiling point range of 60-270 ° C and with less than 2% by volume of aromatics, b) an average distillate having a boiling range of 180 - 320 ° C and with less than 20% by volume of aromatics, and c) a gasoline fraction having d a boiling range of 20 - 190 ° C and with less than 5 vol.% of aromatic substances. 19. Process according to claim 18, characterized by the reduced sulfur content of the oil feed. The solution is lowered to a value below 100, especially below 50, and preferably below 5 ppm (weight), and the halide content less than 50, especially less than 25 and preferably. .1. preferably less than 2.5 ppm. The method according to claim 18 or 19, characterized in that the input is reversed. is added with hydrogen in the presence of a nickel catalyst. *. 21. A process according to any one of claims 18 to 20, characterized by the reduced feed comprising an oil fraction having a boiling point range between 20 and 320 ° C. · 30. ·. 22. A method according to any of claims 18-21, characterized in that the feed comprises heavy and light gas oils. 23. The use of a nickel catalyst suitable for the hydration of food oils in the food industry, in the hydration of oil inputs. 5 Use according to claim 23, characterized in that a nickel catalyst is used which contains 10-30% by weight nickel on an alumina or silica support. Use according to claim 23 or 24, characterized in that a nickel catalyst of type HTC 400 is used. · · ··· · · · · · · · · · · · · · · • · 1, * ·