METHOD FOR CATALYTIC COLD FLOW IMPROVEMENT OF HIGHER HYDROGEN FRACTIONS IN A SULFUR-CONTAINING AND NITROGEN-CONTAINING REACTION FLUID UNDER HYDRATING CONDITIONS

Abstract

The present invention relates to a reactor with an arrangement of at least two successive reactor layers comprising layers of bi-functional catalysts and mono-functional catalysts for improving the cold flow properties of higher hydrocarbon fractions, and to a process for the hydroconversion of a reaction fluid comprising a middle distillate fraction with a Watson K-factor > 11.8 in the presence of sulfur-containing organic compounds and / or nitrogen-containing organic compounds.

Description

SUMMARY

[0001] The present invention relates to a reactor with an arrangement of at least two successive reactor layers comprising layers of bi-functional catalysts and mono-functional catalysts for improving the cold flow properties of higher hydrocarbon fractions, and to a process for the hydroconversion of a middle distillate fraction with a Watson K-factor > 11.8 in the presence of sulfur-containing organic compounds and / or nitrogen-containing organic compounds. DESCRIPTION STATE OF THE ART

[0002] The conversion of carbon-containing resources into commercial products such as fuels or basic chemicals for the chemical and petrochemical industries would not be possible without catalytic hydroconversion. Sources of carbon and hydrocarbons include primarily petroleum fractions, but also distillates and condensates from the thermal treatment (coking, pyrolysis) of coal, natural gas, associated petroleum gas, crude oil, biomass, waste, and especially plastic waste.

[0003] Many of these sources still contain compounds with heteroatoms such as oxygen, nitrogen, or sulfur. Specifically in the case of sulfur-containing sources such as crude oil or coal tar, the sulfur compounds and other heteroatom compounds are typically desulfurized by hydrogenation, for example, over NiMo, CoMo, or NiW catalysts. In combination with additional, usually bifunctional, catalysts, either bond cleavage (cracking) or rearrangement reactions (isomerization) of the hydrocarbon compounds are possible under hydrogenating conditions. The reason for this further conversion is, for example, to adjust the boiling range (hydrocracking) or the viscosity or cold-flow properties (deparaffinization or dewaxing).

[0004] Beshara et al., “Successful Implementation of State-Of-The-Art ULSD / Dewaxing Technology at Irving Oil, Saint John, NB”, Catalagram 103, Spring 2008, p. 36 - 42, describes a process for reducing the sulfur content of a diesel composition by means of an upstream layer of a catalyst for hydrodesulfurization and a downstream layer of a dewaxing catalyst.

[0005] While the refining of light gasoline fractions to motor gasoline focuses on adjusting the concentration of sulfur and nitrogen compounds and achieving knock resistance, the refining of heavier fractions (gas oil, middle distillate) often places a significant emphasis on achieving certain cold flow properties, regardless of their application as turbine fuel, diesel fuel or heating oil.

[0006] For the specification of cold flow properties, for example the standardized methods and norms listed in Table 1, which are mostly used internationally, are available. Table 1: Methods and standards for specifying cold flow properties specification Typical products Standard (Example) Viscosity at -20 °C Turbine fuel EN ISO 3104 Freeze Point Turbine fuel ASTM D5972 Viscosity at 40 °C EURO VI Diesel EN ISO 3104 Cold Filter Plugging Point EURO VI Diesel EN ISO 16329 Cloud Point EURO VI Diesel EN ISO 23015 Viscosity Index (VI) Base oils DIN ISO 51563 Viscosity at 100 °C Base oils EN ISO 3104 Pour Point Intermediate ASTM D5949

[0007] Comparable specifications can also be found in the standards for heating oil (e.g., DIN ISO 51603) or heavy fuel oils for marine applications (EN ISO 8217). The cold flow properties of lubricants made from base oil blends for various automotive applications are regulated separately by the classification system of the Society of Automotive Engineers (SAE). This classification is based on the measurement of viscosities at different temperatures, in accordance with the standards mentioned above.

[0008] With the exception of viscosity, the aforementioned cold flow properties are not physical quantities, but rather measurable quantities representing the result of treating a sample in a standardized setup. Furthermore, the samples under investigation are not pure substances, but mixtures of numerous individual hydrocarbons and heteroatom compounds.

[0009] The pour point (PP) represents the temperature at which, after cooling a liquid sample, free flow is no longer possible due to an increase in viscosity. The PP is therefore closely related to viscosity. The PP itself is of minor importance for direct application in fuels and propellants. However, it is crucial for ensuring the pumpability of raw, intermediate, and finished products in the petroleum industry through pipelines in the event of a potential drop in ambient temperature.

[0010] The Freeze Point (FRP), Cold Filter Plugging Point (CFPP) and the Cloud Point (CP) are directly related to the crystallization behavior of the hydrocarbon mixture, whereby specific melting points can rarely be represented using common measurement methods and the hydrocarbon mixtures in question often exist as very highly viscous pastes (gelling or sulfiting) over long temperature ranges.

[0011] The concentration of long-chain unbranched alkanes, also known as n-alkanes, influences viscosity and the tendency to gel or even crystallize. For the purposes of this invention, long-chain unbranched alkanes are defined as hydrocarbons with a carbon number greater than 11. At sufficient concentrations, these molecules can form relatively stable agglomerates through intermolecular van der Waals interactions. For example, even low concentrations of n-dodecane in a kerosene fraction can lead to a significant deterioration of the FRP (Finished Protein Recycled).

[0012] When determining FRP, this limit temperature is reached as soon as there is no uniform liquid phase in the standardized measuring apparatus.

[0013] When determining CFPP, the quantity and size of the solids present do not allow free flow through a standardized filtration setup. Therefore, FRP largely exhibits crystallization, while CFPP determination shows complete crystallization.

[0014] When determining the critical temperature (CP), the turbidity of the sample is measured by the transmittance of a light beam during cooling. The CP, as the limiting temperature, is reached when the number of crystal nuclei increases to such an extent that the optical measuring device can register the diffraction. This makes the CP a significantly more sensitive measurement because it is related to nucleation, which is the initial stage of every crystallization process. Furthermore, a relatively high limiting temperature can be reached for a sample that appears clear and transparent when determining the CP because colloidally present paraffin crystallites restrict the optical transmittance (Tyndall effect).

[0015] For civil aviation turbine fuels, a globally usable fuel is defined for standard routes (Jet A-1 with FRP = -47 °C). For air traffic in Arctic regions such as northern Canada or Siberia, there is another standardized turbine fuel for specially designed propulsion systems (Jet B with FRP = -60 °C). In addition, there are numerous different, mostly national, turbine fuel specifications for military applications.

[0016] The CFPP (Cold Flow Pressure Point) is so important for specifying cold flow properties because modern diesel engines (e.g., with common-rail technology) have injectors with very small orifices, and due to the high pressures of up to 3000 bar on the connecting lines between the central pump and the injector system, gelling or crystallization of the fuel is never permitted. For regions with arctic climates, CP (Cold Flow Pressure) is also considered as a specifying parameter, and the range of cold flow specifications for very low ambient temperatures is further expanded with classes 0 to 4 (CFPP from -20 to -44 °C and CP from -10 to -34 °C).

[0017] The cold flow properties can be influenced by changing the concentration of long-chain unbranched alkanes or the chemical nature of the overall composition (matrix).

[0018] The following measures can be taken to influence the concentration of long-chain unbranched alkanes (population dewaxing): 1. Addition of fractions with a significantly lower boiling range to reduce the concentration of long-chain n-alkanes. This is achieved by actual admixture (e.g., adding kerosene fractions to diesel fractions) or by changing the operating conditions (cutting temperature) of the distillation column responsible for the diesel product. 2. Reduction of the proportion of high-boiling components by changing the operating conditions (cutting temperature) of the distillation column responsible for the diesel product. 3. Conversion of long-chain unbranched alkanes to shorter-chain hydrocarbons (selective cracking) or to long-chain iso-alkanes by conversion over suitable catalysts. In this context, the slight reduction in viscosity of heavier gas oils by gentle thermal cracking (viscosity breaking or visbreaking), which is used to adjust the viscosity properties of heating oils, should also be mentioned. 4. Addition of pure paraffin-rich fractions with a high proportion of iso-paraffins, such as those obtained as so-called “Hydrogenated Vegetable Oil” (HVO) using the NEXBTL process of NESTE OIL from biogenic fatty acid triglycerides (BTG). 5. Targeted removal of long-chain n-alkanes by extraction (solvent deparaffinization), which is very often used to adjust the viscosity of base oils for lubricants.

[0019] Thus, Rakoczy et al., “Consider catalytic dewaxing as a tool to improve diesel cold-flow properties”, Hydrocarbon Processing, July 2013, p. 67ff., reveals the conversion of a diesel composition using a dewaxing catalyst after a layer of a catalyst for hydrodesulfurization and before a layer of a catalyst for hydrofinishing in order to improve the cold-flow properties of the composition.

[0020] To influence the solution properties, i.e., the possible increase in supersaturation (e.g., measurable via the refractive index) by changing the chemical nature of the matrix (matrix dewaxing), the following measures can be taken: 1. Addition of intermediate fractions from non-hydrogenative cracking processes, such as the FCC process, coking, or visbreaking. After sufficient desulfurization (required sulfur content maximum 10 ppm by mass), these fractions can themselves possess sufficient cold flow properties due to their higher aromatic content and, when added to paraffin-rich fractions, have a positive effect on the matrix effect. 2. Addition of organic hetero compounds, which, as so-called additives, modify the matrix in such a way that the supersaturation is increased, and thus the crystallization is massively controlled to form very small crystal nuclei or crystallites.

[0021] Hydrocarbon mixtures with a predominant proportion of aromatic compounds exhibit unfavorable cold flow properties. This is caused by the interaction of the π-systems, and agglomeration occurs in smaller areas in the form of layers, similar to graphite. In this case, the cold flow properties can be improved by partial hydrogenation of the aromatic compounds or by adding aliphatic hydrocarbons (inverse matrix effect).

[0022] It is important to note that the nature or internal structure of the middle distillate fraction to be treated—that is, a hydrocarbon mixture containing at least 80% by mass of kerosene and gas oil and having a flash point above 30 °C—influences the choice of treatment method. For example, the application of a primarily isomerizing catalytic functionality is extremely difficult if the content of long-chain unbranched alkanes is less than 10% by mass, and the application of a primarily cleavage functionality is difficult if the content of long-chain unbranched alkanes is very high. In the latter case, a liquid sample obtained directly at the reactor outlet typically shows a significant improvement in cold flow properties. However, the middle distillate fraction obtained after distillative separation of the lower-boiling light naphtha fraction, as the final end product, shows only a slight improvement in cold flow properties.Ultimately, the formation of considerable amounts of light naphtha in the product mixture obtained directly after reactor exit is merely a solvent effect, which explains the very good cold flow properties of the entire liquid reaction product at the reactor exit.

[0023] Raw materials, intermediate products, and finished products of the petroleum refining industry are specified and characterized using various standardized measurement methods. To significantly reduce the number of measurements required, numerous mathematical models are available. These models calculate the desired measured values ​​from the density and various parameters of the distillation curve (boiling profile). A large portion of these methods are described in the API Handbook (Technical Data Book - Petroleum Refining, 6th Edition, April 1997, American Petroleum Institute, Washington, DC). Additionally, MR Riazi's "Characterization and Properties of Petroleum Fractions," First Edition, ASTM 2005, West Conshohocken, PA, is very helpful, sometimes containing further formulas.

[0024] The density that is fundamental for all calculations is the “Specific Gravity (SG),” which is determined as described in equation (1). SG=ρSample15 °CρWater15 °C ρSample15 °C = Density of the sample at 15 °C ρWater15 °C = Density of deionized water at 15 °C

[0025] Two standardized laboratory methods are established for determining the distillation curve. The determination of the boiling curve according to ASTM D-2887 represents an instrumental simulation of a laboratory distillation using gas chromatography. The distillation of approximately 150–200 cm³ 3 The sample liquid in a standardized distillation apparatus is described by ASTM D-86. ASTM D-2887 plots the required boiling point against the boiling progress in mass fractions. ASTM D-86 plots the required boiling point against the boiling progress in volume fractions. The API Handbook also describes calculation methods for converting from ASTM D-2887 to ASTM D-86.

[0026] To reduce the boiling curve of ASTM D-86 to a characteristic value, the “Mean Average Boiling Point” (MeABP) is used, the calculation of which is described in formula API 2-07 in Chapter 2 of the Technical Data Book - Petroleum Refining, American Petroleum Institute API, 6th edition, April 1997.

[0027] ASTM D-86, and the MeABP derived from it, together with the SG, serve as the basis for numerous computational models for predicting the physical properties of the hydrocarbon mixture under investigation. ASTM D-2887 is mass-based, very close to the true boiling point of real distillation, and very helpful in estimating the expected yields of fractionation products after distillation.

[0028] Based on the SG and the MeABP, the following specifications can be calculated:

[0029] The mean molecular mass (MMF) for a wide range can be calculated according to equation (2), using the MeABP in °Ra (Rankine). AMW=20.486⋅MeABP1.26007⋅SG4.98308⋅e0.000165⋅MeABP−7.78712⋅SG+0.0011582⋅MeABP⋅SG

[0030] The flash point according to Pensky-Martens closed cup (FP, EN ISO 2719) can be calculated for a wide range using equation (3). Here, the temperature at which 10 vol% of the sample has been converted is denoted as Tf. 10% The temperature is defined in °F (Fahrenheit) according to ASTM D-86, and the result is also given in °F. FP=0.69⋅T10%D−86−118.2

[0031] The Pour Point (PP) can be calculated in a simplified manner according to equation (4), where the MeABP is applied in °Ra and the result for the PP is obtained in °Ra. PP=3.85⋅10−8⋅MeABP5.49⋅10−(0.712⋅MeABP0.315+0.133⋅SG)+1.4

[0032] The Cloud Point (CP) can be calculated in a simplified manner according to equation (5), where the MeABP is applied in °Ra and the result for the CP is obtained in °Ra. log10CP=−7.41+5.49⋅log10MeABP−0.712⋅MeABP0.315−0.133⋅SG

[0033] In addition, further calculable quantities or auxiliary quantities should be mentioned that allow further conclusions to be drawn about the molecular structure.

[0034] The Watson K-factor (K w ) establishes a relationship between the SG and the distillation curve. The calculation is performed according to equation (7) and the MeABP is to be entered in °Ra without units. Kw=MeABP3SG

[0035] A low K w indicates a fraction with a high concentration of aromatics and a low concentration of aliphatics, while a high K windicates a fraction with a high concentration of aliphatics and a low concentration of aromatics. For example, a fully hydrogenated fraction from a Fischer-Tropsch process consists almost exclusively of long-chain unbranched alkanes and reaches a K w from 12.0 to 13.0. An exclusively aromatic hydrocarbon mixture achieves values ​​for K w from 7.0 to 8.0. In real-world oils with sulfur contents above 1.2% by mass, the sulfur content affects the density, and these oils reach K in the case of high paraffin concentrations. w -values ​​significantly below 11.85. Therefore, a corrected specific density SG* is introduced, which replaces SG in equation (7) and, according to equation (8), determines the sulfur mass content x S given in mass %. SG*=SG⋅(1−xS100)

[0036] The aniline point (AP) represents the temperature at which equal volumes of aniline and sample of a middle distillate fraction form a single phase. The lower the AP, the higher the aromatic content of the sample. Equation (9) is used for calculation, where the MeABP is expressed in °Ra. AP=−1253.7−0.139⋅MeABP+107.8⋅Kw+868.7⋅SG

[0037] In contrast to the K w The AP represents a measurable quantity.

[0038] The refractive index at 20 °C (nD20) is calculated in two steps via the modified Huang parameter I according to equations (10) and (11), where the MeABP is used in °Ra. nD20=(1+2i1−i)12 I=0.02266⋅e0.0003905⋅MeABP+2.468⋅SG−0.0005704⋅MeABP⋅SG⋅MeABP0.0572⋅SG−0.72

[0039] The cetane index (CI) for describing ignitability is calculated according to equation (12). The MeABP is used in °F. CI=412.26−7.673⋅API+0.168⋅MeABP+3.503⋅log10MeABP−193.816⋅log10MeABP

[0040] However, the EURO VI specification recommends the calculation according to the significantly more complex ASTM D-4737a (equations (13) and (14)), whereby the temperatures used are in °C according to ASTM D-86. CI=45.2+0.0892⋅(T10%D−86−215)+(0.131+0.901⋅B)⋅(T50%D−86−260) +(0.0523−0.42⋅B)⋅(T90%D−86−310)+0.00049⋅[(T10%D−86−215)2−(T90%D−86−310)2]+107⋅B+60⋅B2 with B=e−3.5⋅(SG−0.85)−1

[0041] To estimate the carbon-to-hydrogen mass ratio (CH), MR Riazi recommends equation (15) in Characterization and Properties of Petroleum Fractions, ASTM, 1st edition 2005. MeABP should be entered in Kelvin. CH=8.7743⋅10−10⋅e0.007176⋅MeABP+30.06242⋅SG−0.00735⋅MeABP⋅SG⋅MeABP−0.98445⋅SG−18.2753

[0042] From this, the atomic hydrogen-to-carbon ratio (HC) can be determined. atomar ) according to equation (16) HCatomar=11.9147CH

[0043] To assess the asphaltenic character of a heavier hydrocarbon fraction, the Conradson carbon residue (CCR) in parts per hundred (%) can be consulted. According to Riazi, equation (17) provides one way to estimate the CCR. CCR=148.7−86.96⋅HCatomic

[0044] By combining the above-mentioned models, further material properties related to cold flow properties can be calculated.

[0045] For the kinematic viscosity at 210 °F (v 210 ) in cSt Riazi recommends the mathematical model (18). log10v210=−0.463634−0.166532⋅API+5.13447⋅10−4⋅API2−8.48995⋅10−3⋅K w⋅API+8.0325⋅10−2⋅KW+1.24899⋅API+0.19768⋅API2API+26.786−2.6296⋅Kw

[0046] The Freeze Point (FRP) in °Ra according to the API model is calculated according to equation (19), where the MeABP in °Ra is used. FRP=−2390.42+1826⋅SG+122.49⋅Kw−0.135⋅MeABP

[0047] The following procedure is typically used to catalytically adjust the cold flow properties of a middle distillate fraction: A layer of a deparaffinization catalyst is positioned between the actual beds of the catalysts for the hydrogenation of unsaturated hydrocarbons and hetero compounds (HDX catalysts) in the reactor or in a separate reactor of a middle distillate hydrogenation plant. The extent of deparaffinization is controlled by the temperature of this catalyst layer.

[0048] For middle distillates with a high concentration of unbranched paraffins (paraffinic nature of the mixture), a deparaffinization catalyst with an additional isomerization function is required because selective cracking is not efficient. The application of this catalyst requires relatively long residence times with mass-specific space velocities (WHSV) < 3.5 h -1 However, in most cases the total accessible reactor volume does not allow the installation of such large catalyst layers for deparaffinization, because otherwise the required volumes for the HDX catalysts would be reduced so drastically that the hydrogenation effect would no longer be sufficient to achieve the desired specifications in addition to the cold flow properties of the products.

[0049] Therefore, there was a need to provide a process or reactor with an arrangement of reactor layers to achieve the cold flow properties of a middle distillate fraction with a Watson-K factor > 11.8 even with higher WHSV, i.e., smaller volumes of deparaffinization catalysts, as is usually the case in small-volume reactors.

[0050] This problem is solved using the reactor according to the invention, wherein the reactor comprises an arrangement consisting of at least two reactor layers, wherein supported bifunctional metal catalysts are located in the at least two reactor layers, the support of which each also has an acidic property, and wherein a monofunctional catalyst is additionally located in the upstream reactor layer, and the reactor layers are arranged such that alkanes and cycloalkanes are hydroisomerized (conditioning) in the upstream reactor layer. In the subsequent downstream reactor layer, non-hydroisomerized long-chain unbranched alkanes are selectively cleaved (final deparaffinization).

[0051] A reactor according to the present invention can be a single reactor housing. In another embodiment, the reactor can consist of several reactor housings arranged one after the other. The reactor layers can either be located in the same reactor housing or they can be located separately in reactor housings arranged one after the other. The reactor layers are defined downstream of the mass flow as the following regions within the single reactor housing or the multiple reactor housings through which the mass flow passes successively.

[0052] The catalysts within the individual reactor layers are arranged in layers.

[0053] In the first reactor layer, in addition to at least one supported bifunctional metal catalyst for conditioning, one or more monofunctional HDX catalysts are also present. The catalysts in this upstream reactor layer can be arranged in separate layers or as a mixture in a single layer within the first reactor layer. The actual positioning of the two reactor layers is determined by the fact that, in the upstream reactor layer, the conversion at the hydrocarbon skeleton is limited to hydroisomerization of the long-chain unbranched alkanes and to cycloalkanes, and possible subsequent reactions such as cleavage of the isoalkanes or the Paring reaction (cleavage of alkyl groups from alkyl methylcyclopentane) are practically impossible.

[0054] Bifunctional catalysts within the meaning of the present invention are catalysts that have two different catalytic functions.

[0055] In the case of supported bifunctional metal catalysts, the metal component applied to the support typically effects hydrogen transfer through dehydrogenation and hydrogenation reactions (first functionality). The additional catalytic component, in particular a catalytically active acidic component, which in the context of this invention is also referred to as the active component or acidic active component, in the support causes polarization and depolarization of active transition states, such as through proton transfer (second functionality).

[0056] To generate the first catalytic functionality, the metal component is usually applied by immersing the carrier in a metal-containing solution, by spraying on a metal-containing solution or suspension, or by so-called incipient wetness impregnation of a metal-containing solution.

[0057] To generate the second catalytic functionality, the support material comprises solids with acidic properties, such as zeolites or mixed oxides. In a particular embodiment, the catalytically active material is present as an inorganic or organic acid, ionic liquid, or complex compound immobilized on the acidic support material.

[0058] The carrier material can be produced by extrusion, tableting, balling, pelletizing, injection molding or 3D printing.

[0059] In one embodiment, the bifunctional metal catalyst is incorporated into a permeable polymer matrix to produce a membrane. This allows its use in a membrane reactor after the metal component has been applied to the support.

[0060] Furthermore, the bifunctional metal catalyst can be applied as a washcoat to honeycomb structures, structured metal foils, or packing materials. This allows for its use in reactive distillation or microstructured reactors after the metal component has been applied to the support.

[0061] In one embodiment, the at least one supported bifunctional metal catalyst of the first reactor layer for hydroisomerization of the mass stream is positioned directly at the reactor inlet or after a very thin layer of conventional hydrogenation catalysts, in particular NiMo, CoMo, or comparable monofunctional HDX catalysts (early conditioning). These are suitable for converting sulfur and nitrogen compounds, in particular for converting basic nitrogen compounds into ammonia or for selectively hydrogenating olefins present in the reaction fluid without simultaneously hydrogenating any aromatic components that may be present.

[0062] In one embodiment, the hydroisomerization catalyst comprises an acidic zeolite or an amorphous solid acid, each having a spaciousness index (SI) of at least 2.0 or greater. Their acidic properties are controlled by adjusting the silicon-to-aluminum molar ratio or by suitable postsynthetic methods such that, with appropriate selection of the hydrogen-transferring metal component of the bifunctional catalyst, hydrocarbons are preferentially hydroisomerized.

[0063] Due to its metal component, the catalyst is also capable of reacting mono- or polycyclic naphthenes with a carbon number greater than ten that are present in the reaction fluid. Predominantly, alkylated methylcyclopentanes or cyclohexanes are formed, some of which react further in a subsequent reaction to form monoaromatics.

[0064] However, an undesired subsequent reaction can also occur, leading to the elimination of the alkyl groups of the alkylated methylcyclopentanes or cyclohexanes. This results in the products formed no longer exhibiting the typical boiling range of a middle distillate fraction, particularly diesel or kerosene, but rather a typical boiling range of a light gasoline fraction. This reduces the yield of middle distillate (Paring reaction).

[0065] If the conditioning catalyst for hydroisomerization is positioned very early in the reactor according to the embodiment, it is possible to avoid this undesired subsequent reaction, since the catalyst also hydrogenates the sulfur-containing and nitrogen-containing compounds present in the reaction fluid by means of the metal component, and thus the metal component necessary for the undesired subsequent reaction is blocked.

[0066] A large distance between the position of the catalyst for hydroisomerization in the first reactor layer and the downstream position of the bifunctional metal catalyst for the selective cleavage of long-chain unbranched alkanes in the second reactor layer enables optimal utilization of the heat generated during hydrogenation via the HDX catalysts between the two catalyst layers.

[0067] The bifunctional metal catalyst for hydroisomerization in the first reactor layer of the reactor according to the invention typically consists of a porous support which is provided with one or more metal components by immersion or impregnation. Metals from transition metal groups IB, IIB, VIB, VIIB, and VIIIB are used as metal components, preferably the metals Mo, Cu, W, Ni, Mn, Co, Cr, Fe, or Zn, or mixtures thereof. The respective concentration of the selected metals is in the range of 0.2 to 25 wt%, preferably 0.5 to 20 wt%, and particularly preferably 1 to 15 wt%, based on the support mass after loss on ignition at 1000 °C.

[0068] The support for this catalyst is preferably selected from aluminium oxide, silicon oxide, metal- or semi-metal-doped aluminophosphates or ceramic.

[0069] The carrier for this catalyst can be produced by extrusion, tableting, pelletizing, injection molding or 3D printing processes.

[0070] Active components can include amorphous aluminosilicates, acidic clays with layered structure, mesoporous acidic layered silicates, tungstate zirconium oxide or sulfonated zirconium oxide, zeolites and related mesoporous substances with pore openings of 2 to 100 nm.

[0071] In one embodiment, the active components comprise tungstate zirconia or sulfated zirconia and are promoted with a transition element or rare earth element.

[0072] Regardless of the choice of active component, materials are suitable which, after loading with 0.5 wt% platinum in a catalytic experiment for the conversion of a C10-C13 n-alkane mixture under the following hydrogenating conditions: LHSV = 3 h -1, H2 / hydrocarbon = 400 NI / I, pressure = 40 barü, temperature range from 220 to 260 °C, cause a conversion to isoalkanes of at least 40 wt%.

[0073] Acidic clays with a layered structure are preferably talc, pyrophyllite, muscovite, phlogopite, illite, margarite, clinclinlite, montmorillonite, saponite, vermiculite, chlorite, corrensite, macaulayite, burckhardtite, surite or kegelite or a mixture thereof.

[0074] In a particularly preferred embodiment, the acidic alumina with layered structure is a mixture of these substances, wherein the proportion of montmorillonite is at least 40% by mass.

[0075] The types of zeolites typically used are those that, due to their acidity, have a primarily isomerizing property and a Spaciousness Index (SI) of at least SI = 2 or higher.

[0076] According to S. Ernst, J. Weitkamp, ​​CY Chen, “The Spaciousness Index: A Useful Catalytic Method for Probing The Effective Pore Width of Molecular Sieves”, PA Jacobs, RA van Santen (eds.) Studies in Surface Science and Catalysis 49 (1989), pages 1115-1129, the SI allows for an assessment of the pore geometry of zeolites and other microporous materials. The SI uses the reaction of n-butylcyclohexane on a bifunctional acidic zeolite (Pt) and is defined by equation (20). SI≡Yiso−ButaneYn−Butane

[0077] Table 2 from SI Zones, TV Harris Microporous and Mesoporous Materials, Issues 35-36 (2000) Pages 31-46 shows an overview of the test reactions mentioned above. Table 2: Overview of SI of selected zeolites Structure type SI CLAY ≈ 1 MTT ≈ 1 FER ≈ 1 MFI ≈ 1 MEL ≈ 1 MTW 3,0 EUO 5,0 OFF 5,5 MOR 7,5 MAZ nv LTL 16,5 BEA 13,0 - 19,0 FAU 21,0

[0078] In a catalytic experiment, commercially available bi-functional catalysts from Clariant, called HYDEX, were used. ® -C, HYDEX ® -E, HYDEX ® -G and the monofunctional catalyst named T-2559 were investigated with regard to their SI. To achieve bifunctionality in material T-2559, the catalyst was conditioned with 0.5 wt% Ni by impregnation with nickel acetate according to the incipient wetness method and subsequent calcination at 350 °C. A reactant mixture of 99.9 wt% n-butylcyclohexane with 0.1 wt% dimethyl disulfide was used for the procedure. At a WHSV of 2.5 h -1 and a hydrogen to n-butylcyclohexane ratio of 250 Ndm 3 / dm 3 At 70 bar g, the yields of isobutane and n-butane were determined in temperature ranges between 320 and 380 °C, the values ​​of the (SI) were calculated, and the following results were obtained: Ni-T-2559 SI = 0.9 HYDEX ® -G SI = 0.8 HYDEX ® -C SI = 3.0 HYDEX ® -E SI = 1.2 HYDEX ® -L SI = 0.8 HYDEX ® -T SI = 0.9

[0079] The zeolites used preferably have topologies selected from CLO, ETR, VFI, AWT, CFI, DON, OSO, SFH, SFN, UTL, AFI, AFR, AFS, AFY, ASV, ATO, ATS, BEA, BEC, BOG, BPH, CAN, CON, CZP, DFO, EMT, EON, EZT, FAU, GME, GON, IFR, ISV, IWR, IWV, IWW, LTL, MAZ, MEI, MOR, MOZ, MRT, MSE, MTW, NPO, OFF, OSI, RON, RWY, SAO, SBE, SBS, SBT, SFE, SFO, SOS, SSY, USI, VET, AEL, CGF, CGS, HEU, preferably selected from DON, AFI, BEA, EMT, FAU, LTL, MOR, MTW, OFF, AEL, CGF, HEU.

[0080] When using zeolites with FAU topology, the acidity can be adjusted by exchanging divalent or trivalent cations according to the Hirschler-Plank mechanism, as described by A.E. Hirschler, J. Catal. 2 (1963), pages 428-439. Additionally, the number of acidic centers, and thus the activity and pore size, and therefore the SI, can be adjusted by exchanging further cations.

[0081] In one embodiment, the carrier has a pore volume, determined by Hg porosimetry according to DIN 66133, of at least 0.05 mm³. 3 / g, preferably at least 0.1 mm 3 / g and preferably at least 0.3 mm 3 / g. In a further embodiment, the carrier has a pore volume, determined by mercury porosimetry according to DIN 66133, of at most 0.7 mm³. 3 / g, preferably of no more than 0.5 mm 3 / g. In another embodiment, the carrier has a pore volume in the range of 0.1 to 0.7 mm³. 3 / g, preferably in the range of 0.3 to 0.5 mm 3 / g on.

[0082] In a further embodiment, a layer of catalyst for the hydrogenation of sulfur-containing organic compounds and nitrogen-containing organic compounds is arranged upstream in the first reactor layer. This is followed by a layer of a bifunctional catalyst for the hydroisomerization of the current, and subsequently by a layer of catalyst for hydrogenation (late conditioning).

[0083] In the so-called contact conditioning process, the catalyst layer for the hydroisomerization of the long-chain unbranched alkanes is located directly in front of the selectively cleaving bifunctional metal catalyst of the second reactor layer, with which the alkanes of the reaction fluid present after the first reactor layer are selectively cleaved and optionally converted into monoaromatics by dehydrocyclization.

[0084] Due to the functionality of the metal component, the hydroisomerization catalyst is also capable of reacting mono- or polycyclic naphthenes with a carbon number greater than ten that are present in the reaction fluid, which can lead to a reduction in the number of polyaromatics. Preferentially, alkylated methylcyclopentanes or cyclohexanes are formed, some of which react further in a subsequent reaction to form monoaromatics.

[0085] However, an undesired subsequent reaction can also occur, resulting in the elimination of the alkyl groups of the alkylated methylcyclopentanes or cyclohexanes, which means that the products formed no longer have the typical boiling range of a middle distillate fraction, especially of diesel or kerosene, but rather a typical boiling range of a light gasoline fraction, thereby reducing the yield of middle distillate.

[0086] In this case, particularly in the case of late or contact conditioning, the active component of the bifunctional metal catalyst is therefore selected to suppress the undesired subsequent reaction described above, namely the elimination of the alkyl groups of the alkylated methylcyclopentanes or cyclohexanes. This is achieved by using an acidic active component with primarily isomerizing properties and an SI < 2.

[0087] The zeolites used preferably have topologies selected from AEL, EUO, IMF, ITH, MTT, NES, RRO, SFG, STF, STI, TER, TON, TUN and WIE, preferably selected from AEL, EUO, MTT, NES and TON.

[0088] Following the first reactor layer, a catalyst is positioned in the second reactor layer, with which the alkanes of the reaction fluid present after the first reactor layer are selectively cleaved and optionally converted into monoaromatics by dehydrocyclization.

[0089] The catalyst of the downstream second reactor layer of the reactor according to the invention is a bifunctional metal catalyst consisting of a porous acidic support and one or more transition metal components. The one or more transition metal components are applied by immersion or impregnation with metal salt solutions. The transition metal components are selected from metals of subgroups IB, IIB, VIB, VIIB, and VIIIB, preferably the metals Cu, Ni, Mn, Cr, Fe, and Zn. The respective concentration of the selected metals is in the range of 0.1 to 10 wt% based on the support mass after loss on ignition at 1000 °C.

[0090] The support for this catalyst consists of an active component and a binder. Aluminum oxide, silicon oxide, or aluminas serve as binders. The catalyst support can be produced by extrusion, tableting, pelletizing, injection molding, or 3D printing.

[0091] Active components can include amorphous aluminosilicates, acidic clays with layered structure, mesoporous acidic layered silicates, tungstate zirconium oxide or sulfonated zirconium oxide, zeolites and related mesoporous substances with pore openings of 2 to 100 nm.

[0092] In one embodiment, the active components comprise tungstate zirconia or sulfated zirconia and are promoted with a transition element or rare earth element.

[0093] In one embodiment, zeolites are used as the active component, which, due to their topology, primarily effect the cleavage conversion of long-chain n-alkanes and have an SI < 2.

[0094] In one embodiment, the zeolites used preferably have topologies selected from ERI, FER, IMF, ITH, MEL, MFI, PON and SFF, particularly preferably selected from ERI, FER, MEL and MFI.

[0095] In another embodiment, wide-pore zeolite structures are used, which are selected by known postsynthetic methods, such as silanization. One such postsynthetic method is treatment with supernatant liquor after completion of the zeolite synthesis, as described in patent application WO 2009 / 0716154 A1.

[0096] In one embodiment, the carrier has a pore volume, determined by Hg porosimetry according to DIN 66133, of at least 0.05 mm³. 3 / g, preferably at least 0.1 mm 3 / g and preferably at least 0.3 mm 3 / g, on.

[0097] In one embodiment, a further layer containing a catalyst for final hydrogenation is positioned behind the second reactor layer. This catalyst serves to subject the reaction fluid after the second reactor layer to additional hydrogenation in order to convert any olefins that may be present.

[0098] An increased concentration of basic nitrogen compounds or the presence of catalyst poisons such as As, Sb, Cd, Hg, V, Po, Pb, Se, and Te can severely restrict the hydroisomerization properties of the bifunctional metal catalyst in the first reactor layer by inhibiting the acid function or poisoning the metal function (hydrogen transfer). The upstream use of a supported pure nickel catalyst or even a nickel precipitation catalyst with nickel contents above 10 wt% and a WHSV space velocity greater than 20 h⁻¹ -1 enables the chemisorption of the metals on the nickel surface and extensive hydrogenation of the basic nitrogen compounds to the corresponding alkanes and ammonia, which has practically no inhibitory effect on the acid centers of the hydroisomerization catalyst of the first reactor layer.

[0099] Particularly in the utilization of pyrolysis oils from biomass or plastic waste, which can contain high concentrations of polyaromatics, enlarging the first reactor layer allows for a significant reduction in the polyaromatic concentration through ring opening and increased formation of alkylated methylcyclopentanes or cyclohexanes. A final deparaffinization via the downstream second reactor layer improves the flowability of the resulting hydrocarbons.

[0100] In the combined hydrogenation of fossil gas oil fractions with biogenic triglycerides from fats and oils, such as animal and vegetable oils, waste from the food industry or catering, it is possible to achieve good cold flow properties by selecting the fossil gas oil mixture component with a Watson K-factor of less than 11.5 and using a smaller catalyst layer of HYDEX. ® -G and a WHSV > 8 h -1to achieve this after hydrogenation, in particular deoxygenation, of the combined gas oil product. However, if the available fossil gas oil fraction in use achieves a Watson K-factor above 11.5, the hydrogenated gas oil product obtained with the process according to the invention can exhibit sufficient cold flow properties, which is particularly advantageous with regard to its use as fuel in jet engines or fuel in diesel or lean-burn engine units.

[0101] In one embodiment of the reactor according to the invention, the composition of the bifunctional metal catalyst of the second reactor layer is as follows, wherein the values ​​refer to a composition annealed at 1000 °C: Metal component: 0.5 - 5.0 wt% Ni; 0.0 - 2.0 wt% Fe or Zn; binder content 15 - 40 wt% aluminum oxide or silicon oxide; acidic zeolite with MFI topology with a stoichiometric Si / Al2 ratio in the range of 15 to 400 and a content of 55 - 80 wt%.

[0102] This material shows the following results in a catalytic experiment under sulfur-containing hydrogenation conditions: Middle distillate section (boiling point 120 °C, boiling point 440 °C, 0.8 wt% sulfur); pressure = 60 bar g, LHSV = 3 h -1 , H2 / starting oil = 400 h -1 in a temperature range of 350 to 360 °C, an increase in the proportion of monoaromatics by at least 1 to 2 mass% and an increase in the proportion of diaromatics by at least 0.5 to 1 mass% when comparing the liquid hydrocarbon mixture at the reactor inlet to the liquid hydrocarbon mixture at the reactor outlet. In one embodiment of the invention, when using a reaction fluid with a sulfur content of less than 3000 ppm by mass, platinum is used as the metal component of the bifunctional catalysts.

[0103] The use of microstructured reactors for the catalytic conversion of hydrocarbons in a wide variety of reaction matrices is very promising. The project described in German patent application DE 10 2012 100 344 A1 by P. Pfeiffer, K. Schubert, and R. Dittmeyer, concerning the design, fabrication, and operation of microstructured reactors for catalytic reactions, is mentioned as representative of many excellent approaches.

[0104] For the isomerization of higher paraffins, such as those found in kerosene or diesel fractions, WO 2004 / 087315 A1 describes the use of a zeolite of type MTW as an acidic component in a bi-functional catalyst system.

[0105] Patent EP 2 200 736 B1 discloses the use of rapid prototyping methods or 3D printing methods for the production of catalysts.

[0106] Layers of inert materials can be positioned above, between, and / or below the catalyst layers. These can be solid reactor internals or packed beds of inert material. These layers can serve to improve the distribution of the reaction fluid components within the reactor or to prevent catalyst material from falling out. In one embodiment, these layers are positioned above the first upstream catalyst layer as an inlet layer and below the last downstream catalyst layer as an outlet layer.

[0107] The invention also relates to a process for the hydroconversion of a middle distillate fraction with a Watson K-factor > 11.8 in the presence of sulfur-containing and nitrogen-containing compounds, with which the cold flow properties of this hydrocarbon mixture are improved, and for which the reactor according to the invention is used.

[0108] The hydroconversion within the scope of the present invention comprises a step of hydroisomerization and a step of selective cleavage, and optionally a step of hydrogenative removal of aromatics, olefins and / or hetero compounds.

[0109] In one embodiment, the process is a hydroconversion of aromatics into alkylated methylcyclopentanes. In another embodiment, the process is a method for modifying the boiling curve and density of a hydrocarbon mixture by cracking or rearrangement reactions.

[0110] The aforementioned quantities Watson K-factor (K W Aniline point (AP), cetane index (CI), Conradson carbon residue (CCR) or the carbon-to-hydrogen ratio are used to assess the internal structure of the oils used, in particular the Watson K-factor (K W ).

[0111] An improvement in cold flow properties is reflected in a change in at least one of the parameters listed in Table 1.

[0112] The inlet temperature is the temperature of the hydrocarbon mixture upon entering the reactor. This is typically in the range of 280 to 390 °C, for example, in the range of 300 to 385 °C.

[0113] The outlet temperature is the temperature of the product stream after exiting the reactor. This is typically in the range of 350 to 420 °C, preferably in the range of 360 to 410 °C, particularly preferably in the range of 365 to 400 °C, and most preferably in the range of 370 to 390 °C.

[0114] The reaction fluid introduced into the reactor comprises a mixture of hydrocarbons from a middle distillate fraction with a Watson K-factor > 11.8.

[0115] The hydrocarbon mixture entering the reactor contains, in addition to the hydrocarbons to be hydroisomerized, impurities of hetero compounds in the form of basic nitrogen compounds with concentrations in the range of 50 to 2000 ppm by mass and / or sulfur compounds with concentrations in the range of 400 to 30000 ppm by mass.

[0116] Furthermore, the mixture typically contains other impurities of heterocompounds such as oxygenates with concentrations in the range of 100 to 30,000 ppm by mass, as well as other impurities such as: Alkali and alkaline earth metals with concentrations in the range of 0.1 to 200 ppm by mass. Higher elements of groups 13 to 17 with concentrations in the range of 1 to 10 ppb by mass (total). Vanadium, cadmium, mercury with concentrations in the range of 1 to 10 mass ppb total

[0117] Before use in the process according to the invention, the catalysts are subjected to sulfidation. Typically, a sulfur-containing gas stream is passed over the catalysts so that they are at least partially converted into the corresponding sulfur compounds. The sulfidation can take place outside the reactor. Preferably, the sulfidation takes place in the reactor after the catalysts have been introduced.

[0118] In the process according to the invention, hydrogenation of the impurities and hydroisomerization of the existing long-chain unbranched alkanes take place in the first reactor layer; in the second reactor layer, selective cleavage of the unreacted long-chain unbranched alkanes and optional dehydrocyclization into monoaromatics take place. This can be carried out with commercially available catalysts such as HYDEX. ®-C or HY-DEX®-E for the first reactor layer and with commercially available catalysts such as HYDEX®-G, HYDEX®-L or HYDEX®-T for the second catalyst layer.

[0119] The described invention is particularly suitable when the nature of the feed oil requires prior dewaxing by isomerization, but the space in the entire reactor is insufficient to place a relatively large layer of an isomerization catalyst between the hydrogenating catalysts. Experience has shown that these purely isomerizing layers require mass-based space velocities (WHSV) in the range of approximately 1 to 5 h⁻¹. -1 depending on the applicable temperature and the desired cold flow improvement. The advantage of the invention lies in the fact that a relatively small layer with a WHSV of at least 6 h -1 , preferably of at least 8 hours -1 , preferably of at least 10 h -1the hydroisomerizing catalyst can be used for pure conditioning of the hydrocarbon stream, the conversion level of which is set so high that the subsequent reaction on a relatively small layer with a WHSV of at least 8 h -1 , preferably of at least 10 hours -1 , preferably of at least 12 h -1 a selectively cleaving and dehydrocyclizing catalyst is applicable in the deparaffinization layer.

[0120] In one embodiment, the reactor is characterized in that the WHSV of the catalyst layer for conditioning and the catalyst layer for deparaffinization last at least 6 hours. -1 , preferably at least 8 hours -1 , preferably at least 10 hours -1 amounts.

[0121] One embodiment of the process can be reactive distillation. In this case, the at least two reactor layers are located in separate columns or separated as packing material in the apparatus for reactive distillation.

[0122] Another embodiment of the method can involve the use of a microstructured reactor. In this case, the at least two reactor layers are separated within a microstructured reactor or in separate microstructured reactors.

[0123] In another embodiment, the at least two reactor layers are in the form of a catalytically active membrane in a membrane reactor. Fig. Figure 1 illustrates the differences in the sequence of reactor layers and catalyst layers of the individual embodiments, with early conditioning shown on the left, late conditioning in the middle and contacting conditioning on the right. Fig.Figure 2 shows a schematic representation of an experimental apparatus for carrying out the method according to the invention. EXAMPLES Experimental apparatus

[0124] To carry out the comparative examples and examples according to the invention, an experimental apparatus such as that described in Fig. Figure 2 shows a schematic representation. The apparatus has four reactors connected in series (REACTOR I to REACTOR IV). The setup was designed to enable near-adiabatic behavior of the reactors. Each reactor was dimensioned to provide a total catalyst volume of at least 2500 cm³. 3 It can accommodate and is suitable for operating pressures in the range of 10 to 150 bar. Supply gases were fed directly from a gas supply and adjusted to the desired operating pressure using appropriate cylinder pressure reducing valves. Commercially available electronic mass flow controllers were used for precise control of the volume flows.

[0125] Nitrogen was used solely for rinsing the system to prevent the formation of explosive air-hydrogen or air-hydrocarbon mixtures.

[0126] The feed oil was placed in a sealed, pressure-equalized container positioned on a scale and pumped into the cross-flow micro heat exchanger I (KSMW I) along with the hydrogen. The cross-flow micro heat exchangers (KSMW I to IV) were designed to heat a reactant flow of 0.5 to 4.0 kg / h within the pressure range specified above to a maximum temperature of 400 °C within a maximum of eight hours (minimum heating capacity of 5 kW). The pipe inlets and outlets were heated by a temperature control system to ensure the desired reactor inlet temperature was maintained. An additional thermocouple was located at each reactor outlet to determine the reactor outlet temperature.

[0127] Connecting the reactors in series allowed for the implementation of different catalyst loading concepts and sequences. The addition of hydrogen and subsequent temperature adjustment via a cross-flow micro-heat exchanger enabled, on the one hand, the simulation of potential intermediate cooling (often called quenching), and on the other hand, the simulation of actual thermal gradients if the available or measured heat of reaction and the resulting temperature gradient across the upstream reactor did not correspond to the expected or actually observed values ​​due to the lack of adiabatic behavior of the setup (overall configuration).

[0128] Furthermore, a bank of high-pressure valves allowed for the bypass of the reaction fluid between the individual reactors. This enabled more precise monitoring of the behavior of the individual catalyst layers via the intermediate products obtained. The downstream reactors could be independently purged of reactants using hydrogen, and the catalyst layers could be brought into a suitable standby mode or standby conditions. The last reactor had a separate reactant discharge (FINAL) to allow for the implementation of cyclic processes, if necessary.

[0129] For discharge, the entire reactor outflow was directed into a temperature-controlled buffer vessel, where the liquid phase separated due to gravity and was discharged at the bottom of the vessel against atmospheric pressure. The collected liquid phase was then transferred to the appropriate product collection vessel. The gaseous component of the reactor outflow was discharged overhead using a backpressure control valve. The mass flow rate was determined using suitable mass flow meters and analyzed by online gas chromatography.

[0130] The gas phase was analyzed by introducing the gas phase, reduced to ambient pressure, through a sampling loop in a standard 6 / 2-way valve in a gas chromatograph. Since the system had two gas phase streams—the stream from the bypass and the stream from the final reactor—a coupled circuit of two 4 / 2-way valves allowed the two streams to be directed to the desired outlets, either via the GC sampling loop or directly to the product collection tank for the respective product stream, i.e., bypass or final. The gas phase stream was cooled again via a cross-flow micro heat exchanger to condense any entrained higher hydrocarbons and then fed into the sample collection tank positioned on a balance.The remaining gas phase was again expelled from the sample collection container upside down, and the mass flow rate was determined using a commercially available mass flow meter and directed into an exhaust pipe.

[0131] The valves in the reactor downstream area were embedded in aluminum blocks with heating elements for temperature control; the pipe connections were also heated using heating cables where necessary. Temperature control was achieved with thermocouples and electronic controllers with PD (phase-dependent) behavior.

[0132] To regulate the reactor inlet temperatures, the temperatures of the heating media (thermal oil) were controlled in a suitable manner via electronic controllers with PID behavior. Reactor loading

[0133] To optimize the flow to the catalyst layers, a bed of lumpy silicon carbide with diameters of 2 to 4 mm was layered in the upper and lower sections of each reactor casing as inlet and outlet layers, respectively. The individual catalyst layers were mixed in a catalyst-to-silicon carbide volume ratio of 2:1. Additionally, the mass of the catalyst material used was determined, and the mass-related space velocity was expressed as Weight Hourly Space Velocity (WHSV) in h⁻¹. -1 specified. Upgrading the catalysts

[0134] Before conducting the catalytic experiment, the catalyst layers were prepared. The following steps were performed: 1) Dry step: The complete apparatus loaded with catalysts was heated with a nitrogen stream at a GHSV of at least 500 h⁻¹ based on the reactor volume. -1The fluid flowed through the reactors against ambient pressure (i.e., without pressure). The valves were positioned so that all reactors were flowed through in series. All cross-flow micro-heat exchangers were controlled to achieve a heating rate of 0.5 °C / min in the heat transfer oil. Once 150 °C was reached at the reactor outlet thermocouples, this temperature setting was maintained for at least five hours. The switching mechanism for the gas analysis feed was configured so that the gas phase of the effluent from REACTOR IV was routed directly to the product feed tank. The temperature of the cross-flow micro-heat exchanger was set to a maximum of 10 °C in the flow path. This ensured that all water in the product feed tank condensed. The drying step was complete when the measured mass increase in the product feed tank was less than 20 g / min. 2) Wetting step: For wetting and sulfidation, a heavy gas oil containing dimethyl disulfide with the following properties was used: • Density (SG according to ASTM D-854) 0.84 to 0.88 kg / dm³ 3 • Boiling range (ASTM D-86) IBP > 165 °C; FBP < 400 °C • Total sulfur content 1.5 to 2.5% by mass

[0135] Under the operating conditions of the dry step, a target flow rate was set at a mass flow rate increase of 0.05 kg / min to achieve a WHSV of 3.5 h⁻¹ based on the total charged catalyst mass excluding silicon carbide. The wetting step was completed as soon as all temperature measurement points were stable – i.e., nearly identical – for half an hour. Sulfurization step: Under the operating conditions of the wetting step, a target flow rate was set at a mass flow rate reduction of 0.05 kg / min to achieve a WHSV of 1.0 h⁻¹. -1The target temperature was achieved based on the total charged catalyst mass, excluding silicon carbide. Subsequently, the process gas was switched from nitrogen to hydrogen. The hydrogen mass flow rate was adjusted to achieve a hydrogen-to-feed oil ratio of 0.04 kg hydrogen per kg feed oil. Once the mass flow rates of the reaction fluids were constant, the back pressure regulator was adjusted to achieve an operating pressure of 50 bar atmospheric gauge. All four cross-flow micro-heat exchangers were heated at a rate of 1 °C / min to achieve a temperature of 280 °C across all reactors. During the heating phase, the switching mechanism for the gas analysis feed was configured to direct the gas phase of the effluent from REACTOR IV via the sample loop to the product feed tank. The concentration of the hydrogen sulfide produced was regularly monitored using a suitable rapid GC program.After achieving a constant hydrogen sulfide breakthrough, the temperature of 280 °C was maintained for at least another hour. Subsequently, the hydrogen flow was increased by a factor of 3.5 at a rate of 2 kg / h, and then the feed oil flow was increased by the same factor at the same rate. This resulted in a WHSV of 3.5 h at the same hydrogen-to-feed oil ratio. -1 The process was stopped. To continue the sulfidation, the temperature was increased to 320 °C across all reactors at a heating rate of 0.5 °C / min. This temperature was maintained for three hours. Afterwards, a temperature of 250 °C was set at a cooling rate of 30 °C / h.

[0136] 3) Stabilization: To stabilize the system, the desired parameters were set on all units using the desired operating oil over a period of approximately one hour, after which the measurement was started. COMPARISON EXAMPLES 1 - Oil used with a balanced paraffin, aromatics and naphthene content

[0137] For the comparative examples 1a and 1b, a native (straight-run) gas oil mixture with the following specifications was used: Specific gravity (SG according to API): 0.852

[0138] Boiling curve according to ASTM D2887 (simulated distillation) Boiling point 141 °C 5 mass % 215 °C 10% by mass 241 °C 30% by mass 276 °C 50% by mass 302 °C 70% by mass 324 °C 90% by mass 355 °C Boiling 391 °C Sulfur content 1.06% by mass nitrogen content 128 ppm by mass. Aromatic content Monocyclic 16.9% by mass, polycyclic 12.0% by mass

[0139] The properties were determined or calculated using common standard methods as described above and are summarized below. Average molecular mass: 236.3 kg / kmol Flash point: 113 °C Carbon-hydrogen mass ratio: 7,09 Hydrogen-carbon atom ratio: 1,68 Conradson carbon residue: 2.5% by mass Viscosities: 4,5 100°F (1,6 210°F ) cSt Aniline point: 72 °C Freeze Point: -7,5 °C Cloud Point: 1,0 °C Cetan Index: 56,8 Mean Boiling Point (MeABP): 299,2 °C Refractive index: 1,473.

[0140] With a K w With a value of 11.85, this oil used can still be considered balanced with regard to its paraffin, aromatics and naphthene content, especially its paraffinicity.

[0141] COMPARISON EXAMPLE 1a: Deparaffinization by primarily selective cracking using a first reactor layer consisting of an upstream layer of a commercially available supported cobalt-molybdenum HDMax®-250 catalyst from Clariant, upstream of a layer of a metal-free, monofunctional, commercially available zeolite-containing catalyst T-2559 from Clariant, and a subsequent second reactor layer of the HDMax®-250 catalyst. In the comparison examples, hydrogenation is largely limited to denitrification and desulfurization; therefore, cobalt-molybdenum catalysts are used. Reactor loading: REAKTOR I: HDMax ® -250: SiC = 1 : 1 (mass-based) REACTOR II: T-2559: SiC = 1 : 1 (mass-based) REAKTOR III: HDMax ® -250: SiC = 1 : 1 (mass-based)

[0142] COMPARISON EXAMPLE 1b: Deparaffinization by primarily selective cracking using a first reactor layer consisting of an upstream layer of Clariant's supported cobalt-molybdenum HDMax®-250 catalyst in front of a layer of a bi-functional catalyst HYDEX ® -G of the Clariant and a subsequent reactor layer of the supported cobalt-molybdenum HDMax®-250 catalyst. Reactor loading: REAKTOR I: HDMax ® -250: SiC = 1 : 1 (mass-based) REAKTOR II: HYDEX ® -G: SiC = 1 : 1 (mass-based) REAKTOR III: HDMax ® -250: SiC = 1 : 1 (mass-based) COMPARISON EXAMPLES 2 - Oil with a highly paraffinic internal structure

[0143] For the comparative examples 2, a native (straight-run) gas oil mixture with the following specifications was used: Specific gravity (SG according to API): 0.822 Boiling curve according to ASTM D-86: Boiling point 163 °C 5 volume % 1 81 °C 10% by volume 204 °C 30% by volume 257 °C 50% by volume 301 °C 70% by volume 325 °C 90% by volume 349 °C Boiling 381 °C Sulfur content: 0.03% by mass Nitrogen content: 90 ppm by mass. Aromatic content: Mono 4.06 mass %Di+ 0.53 mass %

[0144] The properties were determined or calculated using common standard methods as described above and are summarized below. Average molecular mass: 224.4 kg / kmol Flash point: 157 °C Carbon-hydrogen mass ratio: 6,37 Hydrogen-carbon atom ratio: 1,86 Conradson carbon residue: not determined Viscosities: 3,2 100°F (1,2 210°F ) cSt Aniline point: 78 °C Cloud Point: 17,4 °C Cetan Index: 62 Average boiling point (MeABP): 282,3 °C Refractive index: 1,458.

[0145] The structure can be considered highly paraffinic, which is expressed, among other things, in a Kw of 12.4.

[0146] The details of the corresponding operating conditions applied are summarized in Table 3.

[0147] The reaction products were characterized by determining the specific density, simulated distillation using gas chromatographic methods, and the cloud point.

[0148] COMPARISON EXAMPLE 2a: Deparaffinization by primarily selective cracking using a first reactor layer consisting of an upstream layer of Clariant's supported cobalt molybdenum HDMax®-250 catalyst in front of a layer of Clariant's metal-free monofunctional zeolite-containing catalyst T-2559 and a subsequent reactor layer of the supported cobalt molybdenum HDMax®-250 catalyst. Reactor loading: REAKTOR I: HDMax ® -250: SiC = 1 : 1 (mass-based) REACTOR II: T-2559: SiC = 1 : 1 (mass-based) REAKTOR III: HDMax ® -250: SiC = 1 : 1 (mass-based)

[0149] COMPARISON EXAMPLE 2b: Deparaffinization by primarily selective cracking using a first reactor layer consisting of an upstream layer of Clariant's supported cobalt-molybdenum HDMax®-250 catalyst in front of a layer of the bifunctional catalyst HYDEX ® -G of the Clariant and a subsequent reactor layer of the supported cobalt-molybdenum HDMax®-250 catalyst. Reactor loading: REAKTOR I: HDMax ® -250: SiC = 1 : 1 (mass-based) REAKTOR II: HYDEX ® -G: SiC = 1 : 1 (mass-based) REAKTOR III: HDMax ® -250: SiC = 1 : 1 (mass-based)

[0150] The operating conditions were chosen to highlight the differences in functionality during deparaffinization. Details of the specific operating conditions applied are summarized in Table 3.

[0151] The reaction products were characterized by determining the specific density, simulated distillation using gas chromatographic methods, and the cloud point. COMPARISON EXAMPLES 3 - Clarification of selective hydroisomerization

[0152] For the execution of comparative examples 3, a mixture of C10-C13 n-paraffins with the following specifications was used as a starting oil: Specific density: 0,749 Boiling range: 190 - 222 °C

[0153] C10-n-paraffins 13.57 wt%, C11-n-paraffins 39.87 wt%, C12-n-paraffins 10.67 wt%, C13-n-paraffins 35.86 wt%, and total aromatics < 2 wt%.

[0154] The starting oil mixture of C10-C13 n-paraffins was blended with 2-methylnaphthalene in a mass ratio of 1:10. Additionally, dimethyl disulfide (DMDS) was added to the mixture to achieve a total sulfur content of 0.1 wt%.

[0155] After adding 2-methylnaphthalene, a CP of -15 °C was determined.

[0156] COMPARISON EXAMPLE 3a: Deparaffinization by primarily non-selective hydroisomerization using a first reactor layer consisting of an upstream layer of a commercially available supported nickel-molybdenum HDMax®-350 catalyst from Clariant in front of a layer of a bifunctional catalyst HYDEX ® -C of the Clariant and a subsequent reactor layer of the supported nickel-molybdenum HDMax ®-350 catalyst. The use of nickel-molybdenum catalysts enables, under the given reaction conditions, the almost complete saturation of the added 2-methylnaphthalene. Reactor loading: REAKTOR I: HDMax ® -350: SiC = 1 : 1 (mass-based) REAKTOR II: HYDEX ® -C: SiC = 1 : 1 (mass-based) REAKTOR III: HDMax ® -350: SiC = 1 : 1 (mass-based)

[0157] COMPARISON EXAMPLE 3b: Deparaffinization by primarily selective hydroisomerization using a first reactor layer consisting of Clariant's supported nickel-molybdenum HDMax®-350 catalyst prior to a bifunctional catalyst HYDEX ® -E of the Clariant and a subsequent reactor layer of the supported nickel-molybdenum HDMax®-350 catalyst. Reactor loading: REAKTOR I: HDMax ® -350: SiC = 1 : 1 (mass-based) REAKTOR II: HYDEX ® -E: SiC = 1 : 1 (mass-based) REAKTOR III: HDMax ® -350: SiC = 1 : 1 (mass-based)

[0158] Table 3: Summary of operating conditions for comparison examples 1a to 3b (space velocities, WHSV based on the undiluted catalyst masses specified in the text): COMPARISON EXAMPLES REACTOR I 1a 1b 2a 2b 3a 3b P in barü 60 60 60 60 60 60 H 2 / Öl Ndm 3 / dm 3 500 500 500 500 500 500 WHSV h -1 1 1 1 1 1 1 T in °C 300 300 300 300 300 300 T out °C 310 310 310 310 310 310 REACTOR II WHSV h -1 3 3 3 3 3 3 T in °C 350 350 350 350 350 365 T out °C 354 348 356 350 355 369 REACTOR III WHSV h -1 6 6 6 6 6 6 T in °C 354 348 356 350 350 355 T out °C 355 350 357 351 350 355

[0159] The results obtained from comparison examples 1a, 1b, 2a, 2b, 3a, and 3b are summarized in Table 4. Table 4 shows the CP values ​​for both the starting oil and the liquid sample obtained in each comparison example, the so-called True Liquid Product (TLP), as well as the gas oil sample stabilized by subsequent distillation with an intersection point of 145 °C. Furthermore, the calculated product distribution of the integral sample from data of a simulated distillation (ASTM D2887) minus the NH3, H2S, and trace amounts of hydrogen formed is presented. Table 4: Results of comparison examples 1a, 1b, 2a, 2b, 3a and 3b COMPARISON EXAMPLES 1a 1b 2a 2b 3a 3b CP (OPERATION) °C 1 1 17,4 17,4 -20 -20 CP (TLP) °C -11 -18,2 11,5 7,4 -55 -45 CP (145°C+) °C -6,5 -17,5 14,3 12,5 -32 -38 Product distribution hydrocarbons C1-C4 Mass % 3,9 2,6 4,4 3,2 2,3 2,5 C5 - 145 °C Mass % 8,1 5,5 9,9 7,9 15,0 3,5 145 °C+ Mass % 88,0 91,9 85,7 88,9 82,7 94,0

[0160] The comparative examples 1a and 1b clearly highlight the effect of the additional matrix dewaxing (1b) to the standalone population dewaxing (1a) based on the achieved cold flow improvement and the necessary yield losses in the 145°C+ product gas oil fraction.

[0161] Comparison examples 2a and 2b illustrate the problems associated with dewaxing higher paraffinic gas oils and the solvent effect caused by the light naphtha formed (fraction of C5 hydrocarbons and higher-chain hydrocarbons up to a boiling point of 145 °C: C5-145°C+) and the relatively large difference in CP values ​​between TLP and the 145 °C+ gas oil fraction. With a constant increase in the inlet temperature at reactor II, a limiting cloud point for the 145 °C+ diesel fraction is reached for both catalysts. Despite increasing conversion, this results in a CP of 14.3 °C for T-2559 and a CP of 14.3 °C for HYDEX. ® -G CP = 12.5 °C is not undercut.

[0162] In comparative example 3a, the experiment was deliberately conducted in such a way that the introduced 2-methylnaphthalene (10 wt% in an n-paraffin section) was initially almost completely saturated. Due to the relatively high porosity of the zeolite component of HYDEX ® -C allows for massive isomerization of 2-methyldecalin to alkylated methylcyclopentanes, which are subsequently converted to light naphtha molecules (C5-145 °C) via further cleavage reactions (Paring reaction). This is ultimately evident in the massive formation of light naphtha. Due to the lower porosity in the zeolite component of HYDEX ® -E does not produce high levels of light naphtha (C5-145 °C) despite the presence of 2-methyldekalin, and the occurrence of a possible pairing reaction is limited to the outer surface of the active component due to shape selectivity. Examples of invention

[0163] For the implementation of the examples according to the invention, the same higher paraffinic feed gas oil was used as in COMPARISON EXAMPLES 2. The following examples A to C were carried out under varying inlet temperatures T. in and outlet temperatures T out . Example A: Early Conditioning (EC)

[0164] This involved a combined deparaffinization process using conditioning of the feed oil through maximum achievable hydroisomerization, employing a first reactor layer consisting of an upstream layer of the non-selective HYDEX. ®The -C catalyst from Clariant is used directly at the reactor inlet. Subsequently, hydrogenation took place in this first reactor layer using the supported nickel-molybdenum HDMax®-350 catalyst, followed by final deparaffinization in a second reactor layer through conversion on a HYDEX layer. ® -G catalyst followed by a layer of supported cobalt molybdenum HDMax®-250 catalyst for final hydrogenation. REAKTOR I: HYDEX ® -C: SiC = 1 : 1 (mass-based) REAKTOR II: HDMax ® -350: SiC = 1 : 1 (mass-based) REAKTOR III: HYDEX ® -G: SiC = 1 : 1 (mass-based) REACTOR IV:HDMax ® -250: SiC = 1 : 1 (mass-based) Example B: Late Conditioning (SC)

[0165] This involved a combined deparaffinization process using conditioning of the feed oil through maximum achievable hydroisomerization using a first layer of a non-selective HYDEX. ®A catalyst from Clariant, positioned between two layers of a supported cobalt-molybdenum HDMax®-250 and a nickel-molybdenum HDMax®-350 catalyst, facilitates a hydrogenation reaction in the first reactor layer. Subsequently, deparaffinization occurs in the second reactor layer through conversion in a layer of HYDEX. ® -G catalyst and a subsequent final hydrogenation in a layer of a supported cobalt-molybdenum HDMax®-250 catalyst.

[0166] The layers of the HDMax®-350 and the HYDEX ® -G were arranged one after the other in the reactor housing REACTOR III in this example. REAKTOR I: HDMax ® -250: SiC = 1 : 1 (mass-based) REAKTOR II: HYDEX ® -C: SiC = 1 : 1 (mass-based) REAKTOR III: HDMax®-350 HYDEX ® -G: SiC = 1 : 1 (mass-based) REACTOR IV:HDMax ® -250: SiC = 1 : 1 (mass-based) Example C: Contact Conditioning (CC)

[0167] In this process, hydrogenation was achieved in a first reactor layer consisting of a downstream layer of a supported nickel-molybdenum HDMax®-350 catalyst. Subsequently, the feed oil was conditioned in this first reactor layer by achieving maximum isomerization using a layer of selective HYDEX. ® -E catalyst immediately in front of a layer of HYDEX ® -G catalyst for deparaffinization of the reactant stream in a second reactor layer, which was then hydrogenated in a subsequent layer of a supported cobalt molybdenum HDMax®-250 catalyst. REAKTOR I: HDMax ® -350: SiC = 1 : 1 (mass-based) REAKTOR II: HYDEX ® -E: SiC = 1 : 1 (mass-based) REAKTOR III: HYDEX ® -G: SiC = 1 : 1 (mass-based) REACTOR IV:HDMax ® -250: SiC = 1 : 1 (mass-based)

[0168] Table 5: Summary of operating conditions for examples A1, A2, A3, B1, B2, B3, C1, C2 and C3 (space velocities, WHSV based on the undiluted catalyst mass specified in the text) EXAMPLE A EXAMPLE B EXAMPLE C REACTOR I 1 2 3 1 2 3 1 2 3 P in barü 60 60 60 60 60 60 60 60 60 H2 / Öl Ndm 3 / dm 3 500 500 500 500 500 500 500 500 500 WHSV h -1 3 3 3 1 1 1 1 1 1 T in °C 290 290 290 290 290 290 300 300 300 T out °C 298 298 298 300 300 300 308 308 308 REACTOR II WHSV h -1 2 2 2 5 5 5 3 3 3 T in °C 300 300 300 300 300 300 325 325 325 T out °C 309 309 309 305 305 305 330 330 330 REACTOR III WHSV h -1 3 3 3 3 / 3 3 / 3 3 / 3 3 3 3 T in °C 340 350 360 335 345 355 340 350 360 T out °C 340 348 355 340 348 355 340 348 355 REACTOR IV WHSV h -1 6 6 6 6 6 6 6 6 6 T in °C 335 335 335 335 335 335 335 335 335 T out °C 335 338 352 335 338 352 335 338 352

[0169] The results obtained from the examples according to the invention are summarized in Table 6. Table 6 shows the CP values ​​for both the feed oil and the liquid sample obtained in each example, the so-called True Liquid Product (TLP), and the gas oil sample stabilized by subsequent distillation with an intersection point of 145 °C. Furthermore, the calculated product distribution of the integral sample from data of a simulated distillation (ASTM D2887) minus the NH3, H2S, and trace amounts of hydrogen formed is shown. The following can be observed: The achievable cold flow improvement with high-paraffinic feedstocks is significantly lower compared to the exclusive use of T-2559 and HYDEX. ® -G does not reach a limiting cloud point; values ​​above 35 °C can be achieved. Additionally, improved product gas oil yields of 145 °C+ are attained. Table 6: Results of examples A1, A2, A3, B1, B2, B3, C1, C2 and C3 EXAMPLE A EXAMPLE B EXAMPLE C 1 2 3 1 2 3 1 2 3 CP FEED °C 17,4 17,4 17,4 17,4 17,4 17,4 17,4 17,4 17,4 CP TLP EXREACTOR I °C 15,9 15,9 15,9 17,5 17,5 17,5 18,4 18,4 18,4 CP TLP EXREACTOR II °C 16,3 16,3 16,3 15,9 15,9 15,9 15,4 15,4 15,4 CP TLP EXREACTOR III °C -0,9 -8,0 -17,8 -1,3 -8,4 -18,2 -1,8 -8,9 -18,7 CP TLP EXREACTOR IV °C -0,6 -7,7 -17,5 -1,1 -8,0 -17,9 -1,5 -8,6 -18,4 CP 145 °C EXREACTOR IV °C -0,4 -7,4 -17,0 -0,9 -7,7 -17,4 -1,3 -8,3 -17,9 C1-C4 Mass % 1,2 1,6 1,9 1,3 1,7 1,9 1,1 1,5 1,7 C5-145 °C Mass % 5,2 6,5 7,7 5,1 6,4 7,6 5,0 6,3 7,2 145 °C+ Mass % 93,6 91,9 90,4 93,6 91,9 90,5 93,9 92,3 91,1 QUOTES INCLUDED IN THE DESCRIPTION

[0000] This list of documents cited by the applicant was automatically generated and is included solely for the reader's convenience. The list is not part of the German patent or utility model application. The DPMA accepts no liability for any errors or omissions. Cited patent literature

[0000] WO 2009 / 0716154 A1

[0095] DE 10 2012 100 344 A1

[0103] WO 2004 / 087315 A1

[0104] EP 2 200 736 B1

[0105] Cited non-patent literature

[0000] Beshara et al., “Successful Implementation of State-Of-The-Art ULSD / Dewaxing Technology at Irving Oil, Saint John, NB,” Catalagram 103, Spring 2008, p. 36 - 42

[0004] DIN ISO 51603

[0007] Technical Data Book - Petroleum Refining, 6. Auflage, April 1997, American Petroleum Institute, Washington, D.C

[0023] M. R. Riazi, „Characterization and Properties of Petroleum Fractions“, First Edition, ASTM 2005, West Conshohocken, PA

[0023] M. R. Riazi in Characterization and Properties of Petroleum Fractions, ASTM, 1. Auflage 2005

[0041] S. Ernst, J. Weitkamp, C. Y. Chen, „The Spaciousness Index: A Useful Catalytic Method for Probing The Effective Pore Width of Molecular Sieves“, P.A. Jacobs, R.A. van Santen (Herausgeber) Studies in Surface Science and Catalysis 49 (1989), Seiten 1115-1129

[0076] S.I. Zones, T.V. Harris Microporous and Mesoporous Materials, Ausgaben 35-36 (2000) Seiten 31-46

[0077] A.E. Hirschler, J. Catal. 2 (1963) auf den Seiten 428-439

[0080] DIN 66133 [0081, 0096]

Claims

Reactor for the hydroconversion of a middle distillate fraction with a Watson K-factor > 11.8 in the presence of sulfur-containing organic compounds and / or nitrogen-containing organic compounds, wherein the reactor comprises an arrangement consisting of at least two reactor layers, wherein the at least two reactor layers contain supported bifunctional metal catalysts whose supports additionally exhibit acidic properties, and which are arranged to hydroisomerize alkanes and cycloalkanes and hydrogenate unsaturated hydrocarbons and heterocompounds in the upstream reactor layer and to cleave alkanes and optionally convert them into monoaromatics by dehydrocyclization in the downstream reactor layer, and wherein a monofunctional catalyst is additionally present in the upstream reactor layer. Reactor according to claim 1, wherein the supports of the bifunctional catalysts independently comprise aluminium oxide, silicon oxide, a metal foam, ceramic, alumina or a temperature-stable polymer. Reactor according to one of claims 1 or 2, wherein the active components of the bifunctional catalysts independently comprise an amorphous aluminosilicate, zeolite, chlorinated aluminum oxide, tungstate zirconia or sulfonated zirconia, preferably zeolite. Reactor according to claim 3, wherein the active components of the bifunctional catalysts independently comprise tungstate-modified zirconium oxide or sulfated zirconium oxide and are promoted with a transition element or rare earth element. Reactor according to one of claims 1 to 4, wherein the active components of the bifunctional catalysts independently comprise an immobilized acid or ionic liquid on a porous support material. Reactor according to one of claims 1 to 5, wherein the active components of the bifunctional catalysts are embedded independently of one another in a temperature-stable organic, ceramic or metallic matrix by using a 3D printing process (rapid prototyping). Reactor according to one of claims 1 to 6, wherein the bifunctional metal catalyst of the upstream reactor layer comprises a zeolite with a Spaciousness Index SI ≥ 2.0 as the active component and the bifunctional metal catalyst of the downstream layer comprises a zeolite with a Spaciousness Index SI < 2.0 as the active component. Reactor according to one of claims 1 to 6, wherein the upstream reactor layer is positioned very close to or directly in front of the downstream reactor layer in the last third of the reactor and comprises as an active component a zeolite with a Spaciousness Index SI in the range of 1.0 to 1.

6. Reactor according to one of claims 1 to 8, wherein the supported bifunctional metal catalyst of the first reactor layer has a content of metals of the transition groups IB, II B, VI B, VII B or VIII B in the range of 0.2 to 25 wt%, preferably 0.5 to 20 wt%, particularly preferably 1 to 15 wt%. Reactor according to any one of claims 1 to 9, wherein the supported bifunctional metal catalyst of the first reactor layer has a pore volume in the range of 0.1 to 0.7 mm3 / g, preferably in the range of 0.3 to 0.5 mm3 / g. Reactor according to one of claims 1 to 10, wherein the supported bifunctional metal catalyst of the second reactor layer has a content of metals of the transition groups IB, II B, VI B, VII B or VIII B, preferably of the elements Cu, Ni, Mn Cr, Fe and Zn, in the range of 0.1 to 10 wt%. A process for the hydroconversion of a reaction fluid comprising a middle distillate fraction with a Watson K-factor > 11.8 in the presence of sulfur-containing organic compounds and / or nitrogen-containing organic compounds, wherein the process comprises the following steps: - providing a reactor according to any one of claims 1 to 11 - introducing a reaction fluid comprising a middle distillate fraction with a Watson K-factor > 11.8 into the reactor - hydroconversion of the reaction fluid - discharge of the product stream from the reactor Method according to claim 12 for the hydroconversion of aromatics into alkylated methylcyclopentanes. Method according to claim 12 for changing the boiling curve and density of a hydrocarbon mixture by cracking reaction or rearrangement reactions. Method according to one of claims 12 to 14, wherein a catalyst layer for conditioning is located in the upstream reactor position and a catalyst layer for deparaffinization is located in the downstream reactor position. The method of claim 15, wherein the WHSV of the catalyst layer for conditioning and the catalyst layer for deparaffinization is at least 6 h-1, preferably at least 8 h-1, more preferably at least 10 h-1. Method according to any one of claims 12 to 16, wherein the inlet temperature is in the range of 280 to 390 °C, preferably in the range of 300 to 385 °C and the outlet temperature is in the range of 350 to 420 °C, preferably in the range of 360 to 410 °C, particularly preferably in the range of 365 to 400 °C, most preferably in the range of 370 to 390 °C. Method according to one of claims 12 to 17, wherein the at least two reactor layers of the reactor are in separate columns or separated as packing material in a distillation plant for reactive distillation. Method according to one of claims 12 to 18, wherein the at least two reactor layers of the reactor are separated in a microstructure reactor or in separate microstructure reactors. Method according to one of claims 12 to 19, wherein at least one of the supported bifunctional metal catalysts is present in one of the reactor layers in the form of a catalytically active membrane in a membrane reactor. Method according to any one of claims 12 to 20, wherein the reaction fluid contains aromatics, olefins, carbon monoxide, carbon dioxide, carbonyl sulfide or carbon disulfide.

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