Background of the Invention
Field of the Invention
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The present invention relates to a process according to the preamble of claim 1 for dearomatization of petroleum distillates. In general, such a process comprises the steps of feeding an aromatic hydrocarbon feedstock into a hydrotreating process unit, contacting the feedstock with hydrogen in the presence of a catalyst in order to hydrogenate the distillate and the aromatic compounds contained therein to produce a dearomatized product, and recovering the dearomatized product from the hydrotreating process unit.
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The present invention also relates to a dearomatization process according to the preamble of claim 18.
Description of Related Art
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Many petroleum distillate fractions contain aromatic compounds which are harmful for several uses. Typical aromatics of petroleum distillates are benzene and alkyl-substituted aromatic compounds, such as toluene, tetramethylbenzene, diethyl methylbenzene, pentylbenzene and indene.
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In order to provide non- or low-aromatic petroleum products, a number of dearomatization processes have been developed. Principally, they are based on conversion of the aromatic compounds to the corresponding saturated hydrocarbons by reacting said aromatic compounds with hydrogen in the presence of a suitable catalyst at increased pressure and temperature. After dearomatization, the hydrogenated petroleum distillates are usually stabilized by removal of the light, volatile hydrocarbon components.
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Catalytic hydrogenation of aromatic hydrocarbons, in particular benzene and toluene, is also carried out in connection with the preparation of cyclohexane and methylcyclohexane.
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A conventional method for hydrogenation of naphthenic hydrocarbons is disclosed in U.S. Patent Specification No. 3,597,489. The hydrocarbon feed which comprises pure benzene or toluene is hydrogenated in two serially arranged reactors in the presence of a nickel catalyst at a temperature between 140 and 200 °C. The effluent of the first reactor is cooled and the vapour and liquid phases are separated. The liquid phase of the effluent of the first reactor is partially recycled to the first reactor, whereas the vapour phase is conducted to the second reactor.
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A common feature of all prior art dearomatization/hydrogenation processes is that they are applicable only to the specific feedstock for which they have been designed. Thus, the catalyst and the operating conditions employed will differ so much that a dearomatization process developed for hydrogenation of, say, middle distillates cannot be used for converting light naphtha. Nor can a hydrogenation process of solvents be used for dearomatization of middle distillates.
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Furthermore, hydrogenation processes developed for pure feed components, as the above mentioned hydrogenation process of benzene or toluene, cannot be used for processing distillate fractions containing mixtures of aromatic hydrocarbons. This is partly because the reactivity of different aromatic hydrocarbons varies. To take mesitylene as an example of a substituted benzene compound contained in an aromatic feedstock, hydrogenation of 50 % of that compound will require more than 2 times the reaction time needed for hydrogenation of benzene. This means in practice that, under the reaction conditions suitable for complete hydrogenation of benzene, considerable amounts of the heavier aromatic compounds will remain unsaturated.
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Another significant problem relating to hydrogenation of petroleum distillates and other mixed hydrocarbon feedstocks is formed by the impurities contained therein. These impurities stem from various additives used during the refining process. Of particular importance are sulphur and halogen compounds, which poison hydrogenation catalysts and reduce the effective life thereof.
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As a result of the above, at a modern petroleum refinery producing a large variety of petroleum distillates to be used as light and heavy fuels and solvents as well as raw material for petrochemicals and plastics, there must be a number of different hydrogenation units.
Summary of the Invention
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It is an object of the present invention to eliminate the problems of the prior art and to provide an entirely novel process for dearomatization of petroleum distillates and similar hydrocarbon feedstocks containing mixtures of different aromatic hydrocarbons. In particular, it is an object of the invention to achieve a multipurpose hydrotreating process for catalytic dearomatization of petroleum distillates, according to which catalytic removal of aromatics from solvents and middle distillates, catalytic removal of aromatics and/or benzene from naphtha can be carried out by the same basic process configuration, preferably in the same process unit.
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According to one preferred embodiment the above objects are achieved by a process based on the concept of subjecting at least two, and preferably all three, of the following petroleum distillates a to c to hydrogenation in the same hydrotreating process unit:
- a) a solvent cut having a boiling point range of 60 to 270 °C,
- b) a middle distillate having a boiling point range of 180 to 320 °C, and
- c) a naphtha having a boiling point range of 20 to 190 °C.
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According to the invention, each feedstock is contacted with hydrogen in an amount of 100 to 500 Nm3 hydrogen/m3 hydrocarbons of the feedstock. The hydrogenation of each feedstock is further carried out at a pressure in the range of 15 to 55 bar, a temperature of 50 to 250 °C, and at a space velocity, calculated as LHSV, of 0.5 to 10 l/h. The temperature rise, ΔT, in the reaction zone is kept below 60 °C during the dearomatization.
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The invention is further based on the finding that it is possible to process different petroleum distillates containing a large variety of aromatics with varying reactivity towards hydrogen in the presence of a catalyst if the sulphur concentration is reduced to less than 100 ppm, preferably to less than 50 ppm and in particular less than 5 ppm (by weight) before hydrogenation. Simultaneously, the concentration of halogenides should be kept below 50 ppm, preferably below 25 ppm and in particular below 2.5 ppm.
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In particular, the present process for dearomatization of petroleum distillates is characterized by what is stated in the characterizing part of claim 1.
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According to another embodiment, the object of the invention is achieved by a modified process, which is based on the novel concept of reversing the steps of fractionation of the petroleum feedstock and hydrogenation thereof. Thus, this second embodiment provides a process, which comprises
- feeding a petroleum feedstock containing aromatic compounds into a hydrotreating process unit,
- contacting the feedstock with hydrogen in the presence of a catalyst in order to hydrogenate the distillate and the aromatic compounds contained therein to produce a dearomatized product,
- recovering the dearomatized product from the hydrotreating process unit, and
- splitting the dearomatized product into at least two of the following dearomatized petroleum distillates a to c:
- a) a solvent cut having a boiling point range of 60 to 270 °C and containing less than 2 vol.-% aromatics,
- b) a middle distillate having a boiling point range of 180 to 320 °C and containing less than 20 vol.-% aromatics, and
- c) a naphtha having a boiling point range of 20 to 190 °C and containing less than 5 vol.-% aromatics.
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In particular, the process according to the second embodiment is characterized by what is stated in the characterizing part of claim 18.
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The general process concept according to the invention is applicable as such or subject to minor modifications to the preparation of a wide range of different dearomatized and/or benzene-free petroleum products having boiling points from about 20 to 300 °C. Hydrogenation of aromatic solvents and middle distillates, hydrogenation of aromatics and/or benzene of straight-run naphtha, reformate, cracked light and heavy naphtha can all be carried out in the same process unit of the multipurpose process.
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The dearomatization process of the first embodiment mentioned above is applicable to the preparation of different dearomatized and/or benzene free petroleum products in the same process unit and is especially advantageous for smaller refineries. Particular advantages are derived from the fact that one process unit only is needed instead of two or three units, which leads to savings in investment costs. The process also makes it possible to produce many different products periodically and also when annual production rates are low.
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The dearomatization process according to the second embodiment above, in which the petroleum feedstock is first fed into a hydrogenation reactor and then the treated material is split into dearomatized petroleum products (solvents, naphtha and diesel) in a distillation column, is advantageous because many different dearomatization products can be prepared in the same unit at the same time. Particular advantages of that embodiment are that since one unit only needs to be built at the refinery, there is a great saving in investment costs; many different products can be produced at a continuous production rate, and no need for storages of intermediate products.
Bried Description of the Drawings
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- Figure 1 depicts a simplified scheme of a dearomatization process according to the present invention, comprising two hydrogenation reactors and a product stripper;
- Figure 2 depicts a modified process scheme particularly suitable for hydrogenation of light feedstocks; and
- Figure 3 depicts a scheme of an alternative dearomatization process, comprising a hydrogenation reactor and a distillation section for fractionation of the dearomatized feed into suitable petroleum distillate cuts.
Detailed Description of the Invention
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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 benzene and the following typical aromatics of petroleum straight-run or cracked (mainly FCCU) distillates: alkyl-substitued aromatics, such as toluene, tetramethylbenzene, diethylmethylbenzene, pentylbenzene and indene (C9H8).
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"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).
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"Solvents" is an aromatic hydrocarbon feedstock having typically a boiling point (in the following also abbreviated: b.p.) range of 60 to 270 °C, a specific gravity of 0.680 to 0.825 (kg/dm3) and containing before hydrogenation 20 to 30 vol-% aromatics.
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"Middle distillates" is an aromatic feedstock commonly used, for instance, as a fuel for diesel engines and having a b.p. range of 180 to 320 °C. Its specific gravity is about 0.820 and it contains some 20 to 30 vol.-% before hydrogenation.
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The term "naphtha" is used to designate an aromatic feedstock which can be used as fuel of automotive engines. It is, for instance, composed of one or more of the following distillates: straight-run naphtha, reformate, cracked light naphtha and/or cracked heavy naphtha. The b.p. range of naphtha is 20 to 190 °C, sp. gr. 0.600 to 0.830, and aromatics content 0 to 30 vol.-% (before hydrogenation).
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By means of the present invention, it is possible to reduce the aromatics content of solvent products to a value of 0.1 to 2 vol.-%, of middle distillate products 1 to 20 vol.-% and of naphtha components to 0 to 5 vol.-%.
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As earlier mentioned, each feedstock is hydrogenated at the following conditions:
- hydrogen-to-hydrocarbons -ratio: 100 to 500 Nm3 hydrogen/m3 hydrocarbons of the feedstock,
- pressure: 15 to 55 bar,
- temperature: 50 to 250 °C,
- space velocity (LHSV): of 0.5 to 10 l/h
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The particular reaction conditions for each feedstock are selected within the above ranges. Thus, the temperature range (feed temperature) employed for hydrogenation of a solvent feedstock is, for example, 130 to 160 °C, whereas for hydrogenation of light solvents, hexane feed and reformate a suitable temperature range (feed temperature) is below 140, typically 50 to 135 °C. For an integral feed, i.e. a feed comprising hydrocarbons of a broader boiling range and including middle distillates, a temperature of 130 to 200 °C is suitable.
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The hydrogen-to-hydrocarbons -ratio will depend on the amount and type of aromatics present. To take benzene as an example, at a benzene content of 0,1 to 10 wt-%, the feedstock is hydrogenated at a hydrogen-to-benzene -ratio of 3.2:1 to 6:1 to produce a hydrogenated hexane feedstock containing less than 0.1 wt-% benzene.
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It is preferred to reduce the sulphur content of the petroleum and hydrogen feedstock before hydrogenation. Each feedstock fed to a dearomatization unit should have a sulphur content of less than 100 ppm wt, preferably less than 50 ppm wt and in particular less than 10 ppm wt or even less than 5 ppm wt. The corresponding preferred maximum concentration limits of the halogen compounds (chlorides, fluorides, bromides and the corresponding elemental halogens) in the petroleum and hydrogen feedstock are 50 % of the sulphur limits, i.e. 50, 25, 5 and 2 ppm wt.
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In some cases, depending on the origin of the crude oil, the sulphur concentrations meet the above requirements without any additional processing steps. Often, however, the feedstock has to be subjected to desulphurization and dehalogenidization.
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A known process for reducing the sulphur and halogenide contents can be employed. Catalytic hydrodesulphurization is most commonly used, also stripping with inert gas, washing with sodium hydroxide, adsorption on an adsorbent such as clay or molecular sieves, or solvent treatment are possible. Halogen compounds can be removed by similar processes.
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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 250 °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. It is particularly preferable to carry out hydrogenation 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.
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According to a preferred embodiment, the dearomatization is carried out in the presence of a nickel catalyst which is suitable for hydrogenation of edible oils. In connection with the present invention it has surprisingly turned out that these kinds of catalysts, which are known to be highly selective in relation to fatty acids with conjugated or non-conjugated double bonds, are quite capable of catalyzing hydrogenation of a large variety of aromatics contained in petroleum distillates. Furthermore, these highly active catalysts have turned out not to be particularly sensitive to catalyst poisoning by sulphur and halogen compounds appearing in petroleum feedstocks. A third advantage of particular importance to the present applications is that by using food industry catalysts the dearomatization temperature can be kept below 250 °C, preferably in the range from about 50 to 240 °C for all petroleum feedstocks. At such low temperatures, coke formation is greatly reduced. In particular, as the examples below show, good results are obtained with catalysts of the type HTC 400 (supplied by Crossfield) which comprise nickel on alumina, the nickel content typically being about 10 to 30 wt-%.
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As regards the properties of nickel catalysts used in the food industry reference is made to the publication Heterogeneous catalysis; principles and applications, Bond, G.C., pp. 75-78, Oxford University Press, 1974.
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Hydrogenation is an exothermic reaction. Therefore, the temperature tends to rise in the hydrogenation reactor. To avoid too high temperatures and to avoid to 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.
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There are several ways of achieving proper temperature control. Thus, one possibility is to dilute the feed by effluent recycle to lower the concentration of aromatics in the feed. In this embodiment, the effluent of the reaction zone is cooled and partially recycled 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.-% or maximum benzene concentration of 5 wt.-% (or even 4.3 wt.-%) is strived at.
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The cooled effluent or liquid reaction product is recycled from a gas separator vessel or distillation column bottom to the reactor inlet.
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Based on the above, a particularly preferred embodiment of the invention comprises a combination of the following steps:
- reducing the sulphur content of the feedstock to less than 5 ppm wt,
- carrying out the hydrogenation in the presence of a nickel catalyst at a temperature of 50 to 250 °C and at a pressure in the range of 15 to 55 bar, and
- cooling the effluent of the reaction zone and recycling a part of it to the reaction zone in order to keep the temperature rise in the reaction zone at a maximum of 60 °C.
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Another suitable way of restricting the rise of the reaction temperature comprises 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.
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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.
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Recycling of the hydrogenation gas represents a preferred embodiment. Thus, 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.
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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. 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.
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The reactors can be of any suitable type for contacting liquid with gas, trickle bed reactors being particularly preferred.
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The liquid reactor outlet from the first reactor is 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.
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After hydrogenation, the liquid product obtained is fed to a product stabilization section, wherein light (paraffinic) hydrocarbons and hydrogen are removed in a distillation column or a steam stripper.
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In the following, the process embodiments will be examined in more detail with reference to the attached drawings:
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Turning first to the embodiment of Figure 1, it can be noted that the process configuration used for dearomatization of solvent and middle distillates comprises two trickle- bed hydrogenation reactors 1, 2 in cascade and a distillation column 3 for stabilization of the dearomatized product. The petroleum distillate feedstock is conducted to the first trickle-bed reactor 1 with the aid of a pump 5. As depicted in the Figure, the liquid stream is heated in a heat exchanger 4 to feed temperature before it is combined with recycle hydrogen feeds and makeup hydrogen. Alternatively, the feed stream can be combined with makeup hydrogen and recycle hydrogen before it is heated in the heat exchanger 4. In the latter case, the makeup hydrogen feed line is preferably connected to the feed conduit before heat exchanger 4.
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The temperature rise across the reactor is controlled by liquid recycle in which cooled liquid product is recycled from a gas separator vessel 6 or distillation column 3 bottom (the latter recycle is not depicted in the figure) to the reactor inlet. Liquid recycle rate is 0.5 to 3 times the fresh feed rate.
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In the process configuration shown, two reactors 1, 2 in series are used. This embodiment makes it possible to obtain temperature control of the reactors by liquid recycle which is more accurate than with other systems. However, as mentioned above, the two reactors can also be arranged in parallel so that catalyst can be changed without shut down the whole unit.
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Reactor 1 temperature outlet is cooled down to 130 °C in heat exchanger 4. Liquid recycle is separated from the liquid effluent in a gas/liquid separator 6 and fed to reactor 1 inlet.
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The second reactor 2 is operated in the same way as the first. Hydrogen-rich gas is separated from the liquid effluent in a gas separator 8 and recycled to the reactor 1 inlet. Reference numeral 9 indicates a compressor in the gas line. To remove impurities in recycle gas some of the gas is bled to fuel gas or hydrogen net, as shown in Figure 1.
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There is generally no recycle of the liquid effluent from gas separator 8 to either reactor.
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Liquid effluent obtained from the second reactor is fed to product stabilization section 3 The stabilization section is maintained under process conditions sufficient to produce products substantially free of hydrogen and light hydrocarbons. In case of a solvent/middle distillate feed, the overhead product of the distillation column 3 is divided after condensation 10 in the reflux drum 11 into non-condensable gases, which are used as fuel gases, and light liquid hydrocarbons, which can be used as naphta. The dearomatized solvent or middle distillate product is removed as the bottoms product.
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Figure 2 depicts a similar but simplified process configuration which is particularly suitable for hydrogenation of naphtha-type feedstocks. The process comprises one hydrogenation reactor 12 with a stabilization section 13, 19 and 20 as explained above.
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This embodiment is operated in the same way as the above: The petroleum distillate feedstock is fed to the hydrogenation reactor 12 with the aid of a pump 15. The liquid stream is combined with makeup hydrogen and recycle hydrogen feeds and then heated in a heat exchanger 14 to feed temperature. The liquid effluent of the reactor is cooled in heat exchangers 14 and 16 and the gas phase is separated from the liquid phase in a gas separator vessel 17. A part of both the cooled liquid effluent and the hydrogen gas is recirculated to the reactor 12 inlet.
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The dearomatized liquid effluent is conducted to a distillation column 13, in which hydrogen gas and light hydrocarbons are separated from the dearomatized product, which comprises, e.g., (cyclo)hexane or methylcyclohexane or naphtha.
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Figure 3 shows a process scheme for the alternative embodiment, wherein the petroleum feedstock is first dearomatized in a hydrogenation section similar to the one described in Figure 1 and then distilled in a distillation section to produce different cuts.
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Reference numerals 21 and 22 relate to two hydrogenation reactors in cascade. Numeral 23 designates the distillation column for splitting the dearomatized feed from the dearomatization section. Like in the embodiment of Figure 1, the petroleum feedstock is fed to the first dearomatization reactor 21 with the aid of a pump 25. Before the reactor the liquid stream is combined with hydrogen feeds and heated in a heat exchanger 24 to proper feed temperature. The temperature rise across the reactor is controlled by liquid recycle from gas separator 26 to reactor 21 inlet. The second reactor 22 is operated in the same way as the first. The hydrogen-rich gas separated from the liquid effluent in gas separator 28 is recycled back to the inlet of reactor 21 via a compressor 29 in the gas line.
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The liquid effluent from gas separator 29 is fed to the distillation column 23 via a heat exchanger 30. The distillation section (reference numerals 23 - 40) comprises a distillation column 23, side strippers 32 - 34 and an overhead accumulator 41. The distillation column is provided with a reboiler (heater) 35 at the bottom and another in the middle 31. The side strippers 32 - 34 comprise liquid/gas separators wherein the condensed liquid stream is separated from the gas, which is recycled to the distillation column. Each liquid stream is cooled in a heat exchanger 37 - 39 to form distillation cuts A to C with different boiling point ranges. The bottom product of the distillation column forms the heaviest cut and it is also cooled in a separate heat exchanger 36. The liquified overhead is collected with a fan 40 in a reflux drum 41. In the drum 41 light hydrocarbons are separated from water.
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If the dearomatized feed of the distillation section comprises a hydrocarbon fraction with a boiling point range of about 20 to 300 °C, the following products can be recovered: middle distillation products as bottoms products, three solvent products as side drawoffs A to C, and naphtha and fuel gas from the overhead.
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The embodiment according to Figure 3 is described in more detail in Example 9.
Example 1
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Several dearomatized solvents and diesel were prepared in the same process unit in which two trickle-bed reactors were in series. The process configuration corresponds to that depicted in Figure 1. Testruns were carried out with a capacity of 20 tons/hr.
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For processing solvents, a feedstock containing 22 vol.-% aromatics was hydrogenated in the dearomatization unit. The density of solvent feed was 780 kg/m
3 at 15 °C and its b.p. range was
| ASTM D86 distillation, vol.-% |
| IBP | 150 °C |
| 10 | 161 °C |
| 30 | 169 °C |
| 50 | 173 °C |
| 70 | 179 °C |
| 90 | 190 °C |
| 95 | 193 °C |
| FBP | 197 °C |
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Make-up hydrogen feed, mol. wt 4.8, used contained
| hydrogen | 90.9 mol-% |
| methane | 3.2 mol-% |
| ethane | 3.0 mol-% |
| propane | 1.6 mol-% |
| butane | 0.6 mol-% |
| pentane | 0.7 mol-% |
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Solvent feedstock containing less than 5 ppmwt sulphur was heated to 130 °C with hydrogen mixture containing make-up hydrogen and recycle hydrogen. Heated two phase mixture was fed into the first reactor of the dearomatization unit, in which aromatics were hydrogenated from the hydrocarbon feedstock in the presence of an active HTC-400 type nickel catalyst. Reactor pressure was 40 bar. Temperature rise across the reactor was 30 °C. Temperature rise across the reactor was controlled by liquid recycle comprising recycling cooled liquid effluent from a gas/liquid separator to the first reactor. Thus, the first reactor outlet was cooled to 130 °C in heat exchanger and liquid recycle was separated in a high pressure separator and recycled back to the inlet of the first reactor.
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Hydrocarbons and hydrogen-rich gas were mixed and fed to the second reactor of the unit, in which dearomatization reactions were completed in the presence of an active HTC-400 type nickel catalyst. Reactor pressure was 36 bar and temperature rise in the reactor was 8 °C.
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Reactor outlet was cooled to 40 °C in a heat exchanger. Hydrogen-rich gas containing 87 mol-% hydrogen was separated from the liquid effluent in a gas separator and recycled back to the first reactor inlet. The liquid stream was fed to a stabilization section where light components were removed from the dearomatized product. The stabilization section consisted of a distillation column, a reboiler, a condenser and an overhead drum. In the stabilization column hydrogen and light hydrocarbons were distilled from dearomatized solvent product in 200 kPa pressure. Distillation feed temperature was 160 °C. Solvent product b.p. point range was 150 to 200 °C, sp.gr. 0.780 and aromatics content was 0.5 vol.-% or less.
Example 2
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Another testrun was performed in the same unit with heavier solvent feedstock, the sp.gr of which was 0.810. The feedstock used was containing
aromatics 25 vol.-% and
sulphur 5 ppmwt. The b.p. range of the feedstock was
| ASTM D86 distillation, vol.-% |
| IBP | 198 °C |
| 10 | 209 °C |
| 30 | 212 °C |
| 50 | 215 °C |
| 70 | 217 °C |
| 90 | 230 °C |
| 95 | 239 °C |
| FBP | 248 °C |
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A solvent feedstock was heated to 130 °C with hydrogen mixture containing make-up hydrogen and recycle hydrogen. The heated two phase mixture was fed into the first reactor containing an HTC-400 type nickel catalyst. Reactor pressure was 40 bar. Temperature rise across the reactor was 32 °C and it was controlled by liquid recycle.
-
The first reactor outlet was cooled to 132 °C in the heat exchanger and a liquid stream was separated in gas separator and recycled back to the first reactor inlet.
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Hydrocarbons and hydrogen-rich gas were mixed and fed into the second reactor filled with an HTC-400 type nickel catalyst. Reactor pressure was 36 bar and temperature rise in the reactor was 9 °C.
-
The outlet of the second reactor was cooled to 40 °C in a heat exchanger. Hydrogen-rich gas containing 87 mol-% hydrogen was separated from the liquid effluent in the gas separator and recycled back to the first reactor inlet. The liquid stream was fed to the stabilization section where light components were removed from the dearomatized solvent product in 200 kPa pressure. Distillation feed temperature was 200 °C. Product aromatics content was 0.5 vol.-% or less.
Example 3
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A third testrun was done in the same unit with the feedstock containing middle distillates. The feedstock used contained
aromatics 24 vol.-% and
sulphur 5 ppmwt. The feedstock, sp.gr. 0.820, had a b.p. range
| TBP distillation, vol.-% |
| IBP | 129 °C |
| 5 | 178 °C |
| 10 | 191 °C |
| 30 | 212 °C |
| 50 | 230 °C |
| 70 | 250 °C |
| 90 | 270 °C |
| 95 | 279 °C |
| FBP | 320 °C |
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Middle distillate feedstock was heated to 170 °C with hydrogen mixture contained make-up hydrogen and recycle hydrogen. Heated two phase mixture was fed into the first reactor contained an HTC-400 type nickel catalyst. Reactor pressure was 40 bar. Temperature rise across the reactor, controlled by liquid recycle, was 34 °C.
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The first reactor outlet was cooled to 174 °C in the heat exchanger and liquid recycle was separated in the gas separator and recycled back to the first reactor inlet. Hydro-carbons and hydrogen-rich gas were mixed and fed into the second reactor containing an HTC-400 type nickel catalyst. Reactor pressure was 36 bar and temperature rise in the reactor was 3 °C.
-
Reactor outlet was cooled to 40 °C in the heat exchanger. Hydrogen-rich gas containing 87 mol-% hydrogen was separated from the liquid effluent in the gas separator and recycled back to the first reactor inlet. Liquid stream was fed to the stabilization section, wherein light components, e.g. hydrogen, C1 to C5 hydrocarbons and naphtha components, were removed from the dearomatized middle distillate product. Distillation temperature was 240 °C. Product aromatics content was 5 vol.-%.
Example 4
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A testrun for benzene hydrogenation of hexane feed was performed in the Figure 1 type unit without recycle gas and with only one reactor. Benzene was hydrogenated to cyclohexane in a trickle-bed reactor containing an HTC-400 type nickel catalyst. Sulphur content of the feed was less than 5 ppmwt.
-
The hexane feed used, contained
| isopentane | 0.5 wt-% |
| n-pentane | 3.0 wt-% |
| cyclopentane | 4.0 wt-% |
| methylcyclopentane | 9.0 wt-% |
| cyclohexane | 7.0 wt-% |
| benzene | 8.6 wt-% |
| n-hexane | 24.0 wt-% |
| other C6 hydrocarbons | 39.6 wt-% |
| C7 hydrocarbons | 4.3 wt-% |
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The make-up hydrogen used, mol. wt 5.24, contained
| hydrogen | 90.8 mol-% |
| methane | 2.8 mol-% |
| ethane | 2.5 mol-% |
| propane | 2.0 mol-% |
| isobutane | 0.5 mol-% |
| n-butane | 0.4 mol-% |
| pentane | 0.4 mol-% |
| hexane | 0.6 mol-% |
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The hexane feed was heated to 90 °C with make-up hydrogen in the heat exchanger and fed into the reactor. Reactor pressure was 2100 kPa. Benzene hydrogenation to cyclohexane is an exothermic reaction and temperature rise across the reactor was 31 °C. Temperature rise across the reactor was controlled by liquid recycle in which hydrogenated liquid effluent was recycled back to the reactor inlet.
-
Reactor outlet was cooled in the heat exchanger. Hydrogen-rich gas was separated from the liquid stream in the gas separator. Hydrogen-rich gas containing 87.8 mol-% hydrogen was fed to hydrogen net. In the gas separator liquid recycle was separated from the hydrogenated liquid stream and fed to the reactor inlet.
-
The hydrogenated product was fed to the product stabilization section where light hydrocarbons were distilled from hexane product at 600 kPa pressure. Benzene content of the hexane product was 0.1 wt-%.
Example 5
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A second testrun for benzene hydrogenation was carried out with reformate feed having a b.p. range of 55 to 222 °C, sp.gr. of 0.831 and a benzene content of 4.0 wt-%.
-
Reformate feed was heated to 90 °C with make-up hydrogen in the heat exchanger and fed into the trickle-bed reactor filled with HTC-400 type nickel catalyst. Reactor pressure was 2100 kPa. Temperature rise across the reactor was 19 °C and it was controlled by liquid recycle.
-
Reactor outlet was cooled in the heat exchanger. Hydrogen-rich gas was separated from the product in the gas separator. Hydrogen-rich gas was fed to hydrogen net and liquid recycle was to the reactor inlet.
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The hydrogenated liquid stream was fed to the product stabilization section where light hydrocarbons were distilled from reformate product at 600 kPa pressure. Product benzene content was 0.1 wt-% or less.
Example 6
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A third testrun for benzene hydrogenation was constituted by light solvents containing 3.9 wt-% benzene. The b.p. range of the feedstock used was 88 to 105 °C and sp.gr. 0.720.
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The solvent feed was heated to 110 °C with make-up hydrogen and fed into the trickle-bed reactor containing HTC-400 type nickel catalyst. Reactor pressure was 2100 kPa. Temperature rise across the reactor, controlled by liquid recycle, was 17 °C.
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Reactor outlet was cooled in the heat exchanger. Hydrogen-rich gas was separated from the product in the gas separator. Hydrogen-rich gas was fed to hydrogen net. In the separator liquid recycle was separated from the hydrogenated liquid stream and fed to the reactor inlet.
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The hydrogenated product was fed to the product stabilization section where light hydrocarbons were distilled from solvent product at 600 kPa pressure. Product benzene content was 0.01 wt-%.
Example 7
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A fourth testrun for benzene hydrogenation was done with light solvents contained 7.6 wt-% benzene. The density of the feedstock used was 0.680 and its b.p. range was 65 to 75°C.
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The solvent feed was heated to 110 °C with make-up hydrogen in the heat exchanger. Heated mixture was fed to the trickle-bed reactor filled with HTC-400 type nickel catalyst. Reactor pressure was 2100 kPa. Temperature rise across the reactor was 28°C and it was controlled by liquid recycle.
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Reactor outlet was cooled in the heat exchanger. Hydrogen-rich gas was separated from the liquid effluent in the gas separator. Liquid recycle was separated from the hydrogenated liquid effluent and fed to the reactor inlet.
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The hydrogenated product was fed to the product stabilization section where light hydrocarbons were distilled from solvent product at 600 kPa pressure. Benzene content of light solvent was 0.01 wt-%.
Example 8
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Hydrogenation of toluene was studied by simulating a trickle-bed reactor model. The simulation model performed an industrial scale trickle-bed reactor filled with the hydrogenation catalyst. The diameter of the reactor was 1.5 m. The properties of the catalyst required for the mass transfer and pressure drop calculations, are shown in Table 1.
Table 1. | Properties of the catalyst particles |
| Particle diameter |
| | 1 mm |
| Particle length |
| | 3 mm |
| Void fraction of the bed | 0.40 |
| Bulk density | 830 kg/m3 |
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The liquid feed of the simulated reactor was 100 tons/h and it consisted of toluene (5.0 w-%), methylcyclohexane 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 tons/h. The temperature of the liquid and the gas feed was 110 °C. The operating pressure was 3100 kPa.
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The results of the reactor model suggested that the 85 % conversion of toluene was reached with the reactor lenght of 2 m and the reactor length of around 4 m was sufficient for the 99 % conversion of toluene. The increase in temperature was about 38 °C. The computed pressure drop in the reactor was 80 kPa.
Example 9
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The following example relates to a dearomatization process depicted in Figure 3, in which the petroleum feedstock is first fed into a trickle-bed reactor and then the liquid stream was split into dearomatized petroleum distillates; solvents, naphtha and middle distillate products. Feed capacity of the unit was 50 tons/hr.
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The characteristics of the petroleum feedstock were as follows: density at 15 °C is 808 kg/m3 and b.p. range (ASTM D86 distillation) 145 to 320 °C. It contained 20 vol-% aromatics and 5 ppmwt sulphur.
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The make-up hydrogen feed used, mol. wt 4.1, contained
| hydrogen | 93.0 mol-% |
| methane | 3.1 mol-% |
| ethane | 1.5 mol-% |
| propane | 1.3 mol-% |
| butane | 0.8 mol-% |
| pentane | 0.3 mol-% |
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The petroleum feedstock was heated to 150 °C with a hydrogen mixture containing make-up hydrogen and recycle hydrogen. The heated two-phase mixture was fed into the reactor, in which aromatics were hydrogenated from the hydrocarbon feedstock in the presence of an active HTC-400 type nickel catalyst. Reactor pressure was 40 bar. Temperature rise across the reactor was 34 °C and it was controlled by liquid recycle in which cooled liquid effluent was recycled from separator to the reactor inlet
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Reactor outlet was cooled to 40 °C in a heat exchanger. Hydrogen rich gas containing 87 mol-% hydrogen was separated from the liquid effluent in a gas separator and recycled back to the reactor inlet. Liquid effluent was fed to the distillation section, wherein the liquid stream was split into petroleum distillates.
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In the distillation section the liquid feed was heated to 175 °C in a heat exchanger and fed into a distillation column in which liquid stream was distilled into dearomatized naphtha, solvent and middle distillate products at a pressure of 160 kPa. The column diameter was 3 m and the height 50 m. The column consisted of 72 trays and pump-around with cooler of 2.3 MW from tray no 38. The feed was into 68 th tray.
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In addition, column consisted of three side-draw-off trays no 58, 38 and 18, from which three solvent products A, B and C with different b.p. range were removed. Each solvent product was stabilized in steam stripper with 8 trays. Sidestrippers for solvents A and B diameter was 1.0 m and height 6.0 m. Solvent C sidestrippers diameter was 1.1 m and height 6.0 m. Distilled products are specified in table 1.
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Column bottom temperature was 310 °C and the column bottom was reboiled by heater (duty 10 MW). Column bottom product was middle distillate.
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Top temperature of the column was 120 °C and overhead condenser duty was 10 MW. The overhead product was unstabilized naphtha which was fed to stabilization section where light hydrocarbons were removed from the dearomatized naphtha product. It is also possible to feed distillation feedstock first to the product stabilizer and then, after stabilization, to conduct the feed to the distillation column in order to split petroleum distillates.
Table 2. | Properties of distillated dearomatized product |
| Product | Rate, tons/hr. | Aromatics, vol-% | B.p. range, °C | Sp. gravity | Flash point, °C |
| Solvent A | 1.7 | ≤ 0.5 | 213 - 270 | 0.823 | 91 |
| Solvent B | 11.0 | ≤ 0.5 | 188 - 246 | 0.810 | 72 |
| Solvent C | 16.0 | ≤ 0.5 | 170 - 221 | 0.797 | 54 |
| Unstabilized naphta | 9.4 | ≤ 0.2 | - 184 | 0.777 | |
| Middle distillate | 11.8 | ≤ 1.0 | 241-328 | 0.845 | |
