DK145349B - PROCEDURE FOR PREPARING A FUEL OIL WITH REDUCED SULFUR CONTENT - Google Patents

Description

i db i (19) DENMARK V /

| J | (12) PRESENTATION WRITING ου 145349 B

DIRECTORATE OF THE PATENT AND TRADEMARKET SYSTEM

(21) Application No 2626/75 (51) (ntCI.3 C 10 6 65/04 (22) Filing day 11 · May 1973 (24) Running day 11 May 1973 (41) Available on 15 · Nov 1973 ( 44) Submitted Nov. 1, 1982 (86) International application no.

(86) International filing day (85) Continuation day - (62) Master application no. -

(30) Priority May 12, 1972, 252787, US

(71) Applicant UNIVERSAL OIL PRODUCTS COMPANY, Des Plaines, US.

(72) Inventor Newt Morrie Hallman, US: Bernhard Albert Oeltgen, US.

(74) Associate Engineering Company Budde, Schou & Co.

(54) Process for the production of a reduced sulfur fuel oil.

Catalytic desulphurisation is a well-known process thoroughly described in the field of petroleum technology, the literature on this subject being full of references to suitable desulphurisation catalysts, catalyst preparation processes and various operating conditions. By the term "desulfurization" is also meant HQ destructive removal of sulfur-containing compounds via the conversion thereof) into hydrogen sulfide and hydrocarbons, and it is often included in the broad term "hydro refining". Hydro refining processes are carried out at operating conditions and severity levels, which serve to promote primarily denitrification and of sulfur and to a lesser extent asphaltene conversion, conversion

Si 145349 2 of non-distillable hydrocarbons, hydrogenation and hydrocracking. Thus, the terms "hydro refining" and "hf sulfur" are generally used synonymously to refer to a process in which a hydrocarbon charge material is "purified" to produce a charge material suitable for final use in a subsequent hydrocarbon conversion system, or for the purpose of The product of the present invention is advantageously used to produce a fuel oil containing less than about 1.5% by weight of sulfur, and often less than about 0.5% by weight while simultaneously effecting it. at least partial conversion of the feedstock to lower boiling hydrocarbon products.

A relatively recent recognition of the need to prevent the release of various pollutants into the atmosphere has resulted in the authorities in many countries introducing control measures in a number of areas. An important area among these is the burning of high sulfur fuels, primarily coal and fuel oil, the combustion of which results in emissions into the atmosphere of excessively large amounts of sulfur dioxide. In the case of fuel oils derived from crude oil, the need for these has increased noticeably as a result of an increase in worldwide energy consumption. Of greater importance, however, is the fact that legal restrictions on the concentration of sulfur in fuel oil, calculated as a basic element, have already been introduced to a maximum of 1.5% by weight. Expert experts in this field are currently predicting that within so many years, the sulfur content of fuel oil will be limited to a maximum level of 0.5% by weight. It is this object that the present invention and the combination process embraced therein are particularly directed towards, i. preparation of fuel oils containing less than approx. 1.5% by weight of sulfur and, if necessary, less than approx.

0.5% by weight. The increasing need for low sulfur fuel oils has also necessitated the conversion of the "precipitate into the drum" with crude petroleum. In other words, the increasing need for fuel oil has necessitated the use of practically $ 100.0 of crude oil.

In accordance with the combination process of the present invention, acceptable fuel oils are obtained via desulphurization of crude petroleum, bottom products from a distillation made by atmospheric pressure, bottom products from a distillation carried out under vacuum, heavy circulation materials, crude oil residues, crude oils from which the most volatile parts are crude hydrocarbonaceous oils extracted from tar sands, etc. Crude oils and the heavier hydrocarbon fractions and / or distillates prepared therefrom contain nitrogenous and sulfurous compounds in exceedingly large quantities, the latter usually being in the range of from about. 2.5 to approx. 6.0% by weight, calculated as elemental sulfur. In addition, these heavy hydrocarbon fractions, sometimes referred to in the petroleum technique, are sometimes referred to as "black oils", organo-metallic pollutants, which include substantially nickel and vanadium in the form of higher molecular weight porphyrins and asphaltic materials. As an example of the charge materials to which the present invention may be applied, a bottom product from a distillation carried out in vacuo may be mentioned.

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a density of 1.021 g / cm and containing 4.05% by weight of sulfur and 23.7% by weight of asphaltenes, a Middle Eastern oil from which the most volatile parts have been removed, having a density of 0.99 g / cm and containing 10.1% by weight of asphaltenes and 5.20% by weight of sulfur and a residue from distillation made under vacuum with a density of approx. 1,009 g / cm and containing 3.0 wt% sulfur, 4,300 wt ppm nitrogen and with a 20 wt% distillation temperature of approx. 568 ° C. The use of the present process enables the maximum extraction of low sulfur fuel oil from these heavier hydrocarbon-containing charge materials.

The object of the present invention is to provide a process for desulphurising hydrocarbonaceous material, in particular a multistage process for desulphurizing heavy hydrocarbonaceous material to produce an acceptable fuel oil containing less than about 1.5% by weight of sulfur, while simultaneously maximizing recovery of fuel oil product from hydrocarbon-containing black oils at a lower level of operational severity.

Accordingly, the present invention relates to a process for the production of a fuel oil containing less than ca. 1.5% by weight of sulfur from a black oil charge material, characterized in that (a) the charge material and hydrogen are reacted in a first catalytic reaction zone under desulphurisation conditions selected to convert sulfur-containing compounds into hydrogen sulfide and hydrocarbons, (b) separating the resulting effluent from the first reaction zone into a first separation zone to provide a first predominantly vapor phase and a first predominantly liquid phase, (c) reacting at least a portion of the first liquid phase with the practically hydrogen sulfide-free third vapor phase from step (f), in a second catalytic reaction zone under desulfurization conditions selected to convert further sulfur-containing compounds into hydrogen sulfide and hydrocarbons; (d) separating the resulting residual stream from the second reaction zone into another separation zone to provide of another vapor phase and another liquid phase, (e) id a at least a portion of the second vapor phase is recycled to the first reaction zone for reaction with the charge material, and fuel oil is recovered from the second liquid phase, (f) the first vapor phase is separated into a third separation zone to provide a third vapor phase from which hydrogen sulfide is removed, and a third liquid phase, and (g) an additional amount of fuel oil is recovered from the third liquid phase.

As described below, other embodiments of the present invention are first and foremost in preferred ranges of process variables, various treatment conditions, and preferred composite catalysts for use in the fixed mass catalytic reaction zone. Particularly good results are obtained according to the invention by combining at least part of the third liquid phase with the first vapor phase prior to introduction into the third separation zone.

Embodiments of the present invention will become apparent from the more detailed description of the combination method of the present invention.

The present invention employs at least two fixed mass catalytic reaction systems in combination with a specific range of separation facilities that enable more efficient utilization of double-acting hydrogen throughout the process. It is to be understood that each reaction system may consist of one or more reaction vessels with suitable heat exchange facilities occasionally.

Briefly, the combination process is carried out by initially reacting the charge material with a relatively impure (with respect to hydrogen sulfide concentration) hydrogen flow in a first separation zone. This first reaction zone actually completes the relatively light desulfurization of the lower boiling sulfur-containing compounds. The product exit stream is separated at substantially the same temperature and pressure to provide a higher boiling normally liquid phase which is reacted in the second reaction system with a relatively pure hydrogen phase. Operating conditions in both reaction systems include a pressure of approx. 14 to approx. 210 ato, one hydrogen concentration in the range of approx. 85Ο to approx. 50 * 550 nxcP / 1000 liters, a liquid space velocity of approx. 0.25 to approx. 2.50 vol / vol / hour and a maximum catalyst mass temperature of approx. 315 to approx. 485 ° C. In order to facilitate the correct series flow of material through the process, the second reaction zone is often kept under an applied pressure of at least approx. 1.8 ato greater than that imposed on the first reaction zone. Since the more difficult desulphurisation reactions are carried out in the second reaction zone, the maximum catalyst mass temperature in the same is often about 5 ° C higher than the temperature of the first reaction zone.

The advantages of the present process are numerous.

However, the most significant of these is a significant reduction in the maximum catalyst mass temperature to achieve the desired degree of desulfurization. Although the maximum catalyst mass temperature can be in the range of approx. 515 to approx. 485 ° C, it is thus not uncommon to carry out the above-mentioned combination process at maximum catalyst mass temperatures of less than ca. 455 ° C. Furthermore, the combination process of the present invention permits the use of a relatively impure high flow stream to effect the desulfurization of lower boiling sulfur-containing compounds in the initial reaction zone. Another advantage consists in increasing the effective acceptable life of the composite catalyst used in both reaction zones.

Although the composite catalysts may have different physical and chemical characteristics in many situations, they may be identical. Notwithstanding this, the composite catalysts used in the combination process of this invention comprise metal components selected from the metals of Groups VI-B and VIII of the Periodic Table and compounds thereof. In accordance with 145349 6

The Periodic Table of The Elements, E.H. Thus, Sargent & Co., 1964, suitable metal components are components selected from chromium, molybdenum, tungsten, iron, ruthenium, osmium, cobalt, rhodium, iridium, nickel, palladium and platinum. Furthermore, recent studies have indicated that composite catalysts, especially for use in the conversion of feed materials of extremely high sulfur content, are improved by incorporating a zinc and / or bismuth component. In the present specification and claims, it is intended that the use of the term "component" when referring to the catalytically active metal (s) comprises the presence of the metal as a compound such as an oxide, sulfide, etc., or as an element . No matter how it is, the concentrations of the metal components are calculated as if the metal is present as the element in the composite material. Although neither the exact composition nor the method of preparing the various composite catalysts is considered essential to the present invention, certain aspects are preferred. For example, since the charge material of the present process is generally of a high boiling nature, it is preferred that the catalytically active components tend to effect a limited degree of hydrocracking while at the same time promoting the conversion of sulfur-containing compounds to hydrogen sulfide and hydrocarbons. . The concentration of the catalytically active metal component (s) is primarily dependent on the specific metal as well as the physical and / or chemical characteristics of the charge materials. The metal components of Group VI-B are e.g. usually present in an amount in the range of about 4.0 to approx. 0.0% by weight, the iron group metals in an amount in the range of about 2.0 to approx. 10.0% by weight, while the Group VIII noble metals are preferably present in an amount in the range of about 10% by weight. 0.1 to approx. 5.0% by weight. When a zinc and / or bismuth component is used, it is usually present in an amount in the range of from about. 0.01 to approx. 2.0% by weight. All concentrations are calculated as if the components are present in the composite catalyst as element.

The porous support material with which the catalytically active metal component (s) is combined is constituted by a refractory inorganic oxide of the kind thoroughly described in the literature. When it is of the amorphous type, alumina or alumina in combination with 10.0 - about 10 is preferred. 90.0% by weight silica. When treating heavier charge materials containing a significant amount of hydrocarbons having boiling points at normal oxygen at a temperature of approx. 510 ° C, it may be convenient to use a support material comprising a crystalline aluminum silicate or a zeolitic molecular sieve. In most cases, such a support material is used in the treatment of the partially desulfurized feed material to the second reaction zone. A suitable zeolite material includes mordenite, faujasite, type A or type U molecular sieves, etc., and these can be used in practically pure form. However, it is to be understood that the zeolite material may be contained in an amorphous matrix such as silica, alumina and mixtures of alumina and silica. It is further intended that a halogen component may be incorporated into the composite catalysts, such component being selected from fluorine, chlorine, iodine, bromine and mixtures thereof. The halogen component is bonded to the carrier in such a manner as to produce a composite final catalyst containing from ca. 0.1 to approx. 2.0% by weight.

The metal components can be incorporated into the composite catheter lysator in any convenient way, e.g. co-precipitation or co-formation with the support material, ion exchange or impregnation of the support material or during a co-extrusion process. After the incorporation of the metal components, the support material is dried and subjected to a high temperature calcination or oxidation at a temperature of from ca. 400 to approx. 705 ° C. When a crystalline aluminum silicate is used in the support material, the upper limit of calcination is preferably approx. 5 ^ 0 ° C.

As regards the operating conditions imposed on the catalytic reaction zones, they are primarily selected so as to effect the conversion of sulfur-containing compounds to hydrogen sulfide and hydrocarbons. As stated above, the operating conditions imposed on the second reaction zone often result in greater operational severity, although such a technique is not essential. The variation in operating severity levels between the two reaction zones can be obtained by adjusting pressure, maximum catalyst mass temperature and liquid room velocity variables. The operation at the higher severity of the second reaction zone is usually carried out at an elevated pressure, an elevated maximum catalyst mass temperature, and at a somewhat reduced liquid space velocity or some combination thereof.

With the exceptions listed above, suitable ranges for the different operating variables are usually the same for both reaction systems. The pressure is thus in the range of approx.

8 to 14 approx. 210 ato, the hydrogen concentration is from approx. 850 to approx. 50,550 nm / 1000 liters, the maximum catalyst mass temperature is in the range of approx. 315 to approx. 485 ° C and the liquid space velocity varies from approx. 0.25 to approx. 2.50 vol / vol / hour. Given that the reactions carried out are predominantly exothermic, an increasing temperature gradient is observed in both reaction zones as the reactants pass through the catalyst mass. Preferred operating conditions dictate that the increasing temperature gradient should be limited to a maximum of ea. 40 ° C, and for controlling the temperature gradient, it is within the scope of the present invention to use quenching streams, either normally liquid or usually gaseous, introduced at one or more intermediate sites in the catalyst mass. Before the present invention is further described, and particularly with respect to the embodiment illustrated in the drawings, various definitions are considered necessary to obtain a clear understanding. In the present specification and claims, it is intended that a "pressure substantially the same as" or a "temperature substantially the same as" means the pressure or temperature of the container subsequent to the flow, except from the normal pressure drop due to fluid flow through the system and the normal temperature drop due to transfer of material from one zone to another. For example, when the pressure at the approach to the first reaction zone is approx. 162 ato and the temperature is approx. 370 ° C, the first separation zone operates at practically the same pressure and temperature of approx. 155 at 37 and 37 ° C. Similarly, it is intended that the use of the term "at least part of" when referring to either a predominantly vapor phase or a predominantly liquid phase comprises both an aliquot portion and a selected fraction. Thus, at least part of the first predominantly vapor phase is introduced into the second reaction zone after the removal of hydrogen sulfide therefrom, while at least part (in this case an aliquot part) of the first predominantly liquid phase can be recycled for combination with the fresh feed-batch material.

Other conditions and preferred operating techniques are given in connection with the description below of the process of the present invention. In describing this combination desulfurization process, reference is made to the drawing which illustrates a specific embodiment. In the drawing, the embodiment is illustrated by means of a simplified flow diagram in which details such as compressors, pumps, heating and cooling means, instrumentation and control means, heat exchange and heat recovery circuits, valves, commissioning lines and similar apparatus have been omitted as non-essential for an understanding of the technique involved. The use of such mixed accessories for modifying the method is obvious to one skilled in the art, and its use does not remove the resulting method from the scope of the claims.

For demonstration of the illustrated embodiment, the drawing is described in connection with the conversion of a gas oil charge material into a commercial scale unit. It is to be understood that the charge material, stream compositions, operating conditions, design of fractionating means and separators and the like are given by way of example only and may be widely varied within the scope of the present invention as defined by the claims.

With reference to the accompanying drawings, the illustrated embodiment is described in connection with a commercial scale unit designed to produce maximum amounts of a fuel oil containing less than approx. 1.0% by weight of sulfur

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a gas oil charge material having a density of approx. 0.957 g / cm. Other relevant properties of the gas oil charge material include a sulfur concentration of 4.0 wt%, 60.0 wt ppm vanadium and nickel, 2.4 wt% heptane insoluble material, a Conradson carbon factor of 8.2 and a 65.0 wt% distillation temperature of approx. 565 ° C. The charge material is fed into the process via a line 1 in an amount of approx. 740.55 moles per hour. The batch material is mixed with a hot (365 ° C) recycled gas phase from a conduit 2, the latter containing approx. 76.0% by volume of hydrogen. The mixture proceeds through conduit 2 and is passed through it into a reactor 3 at a temperature of approx. 347 ° C and a pressure of approx. 158 ato.

In the reactor 3, a composite catalyst with 2.0 wt.% Nickel and 16.0 wt.% Molybdenum combined with an amorphous support material consisting of 88.0 wt.% Alumina and 12.0 wt.% Silica, is placed. 1.0 vol / vol / hour. The discharge current from the reactor zone 3 1653A9 is taken out via a conduit 4 at a temperature of ea. 365 ° C and a pressure of approx. 151 ato and introduced into a heat separator 5 at practically the same pressure and temperature. The heat separation zone 5 serves to provide a predominantly liquid phase in a conduit 6 and a predominantly vapor phase in a conduit 7. A component analysis of the discharge stream from the first reaction zone (conduit 4), the liquid phase from the heat separator (conduit 6) and the steam phase from the heat separator (conduit) 7) is given in Table I below.

Table I

Product distribution from the reaction zone 3

Component mol · / hour Wire 4 Wire 6 Wire 7

Water 8 ^ 6.89 - 856.89

Hydrogen sulfide 350.87 16.75 334.12

Hydrogen 9663.08 313.60 9349.48

Methane 1358.04 49.68 1308.35

Ethane 139.89 10.26 129.63

Propane 111.16 · 9.15 102.01

Butanes 68.01 6.88 61.13

Pentans 24.50 3.11 21.39

Hexanes 18.30 2.79 15.51

Heptane - 190 ° C 73.05 17.07 55.98 190 - 365 ° C 106.05 58.64 47.42

Light fuel oil 284.18 157.12 127.05

Heavy fuel oil 54o, o6 527.72 12.33 Total 13574.04 1172.77 12401.27

The predominantly vapor phase in conduit 7 is mixed with 1420.00 moles per ml. hour of a cold flash evaporative liquid recycling stream, the source of which is described below. The mixture proceeds through conduit 7 and is passed through it into a cold separator 10 at a pressure of approx. 148 ato and a temperature of approx. 60 ° C. A hydrogen-rich vapor phase is taken out of the cold separator 10 via a conduit 11 and passed through it into a hydrogen sulfide removal system 12. A predominantly liquid phase is taken out via a conduit 15 and fed through it into a cold lightning evaporator separator 16 at practically the same temperature. The cold lightning evaporation zone 16 operates at a pressure of approx. 14 ato and a temperature of approx. 57 ° C to provide a vent gas stream in a conduit 17 and a predominantly liquid product phase in a conduit .18, a portion of which is recycled for combination with the predominantly gaseous phase in conduit 7. Component analyzes of the hydrogen-rich stream in conduit 11, the vent gas stream in the conduit. 17 and the normally liquid material in conduit 18 sent to stabilization facilities for product recovery are set forth in Table II below.

Table II

Power analysis, cold separation system

Component mol / hour Wire 11 Wire 17 Wire 18

Water 14.76 ----

Hydrogen sulfide 273.51 46.04 8.21

Hydrogen 9158.03 189.42 1.14

Methane 1219.47 83.69 2.93

Ethane 111.17 14.97 1.97

Propane 79.87 13.40 4.92

Butanes 41.19 7.29 7.16

Pentans 10.47 1.82 5.14

Hexanes 5.00 0.85 5.44

Heptane - 190 ° C 2.52 0.39 29.74 190 - 365 ° C 0.02 - 26.72

Light fuel oil 0.05 0.01 71.64

Heavy Fuel Oil 0.05 - 7.00 Total 10916.04 357.86 172.01 12 US $ 49

As indicated below, a wash water stream is fed into the process via a conduit 8. This water is eventually removed from the process system using a "dip-leg" not shown in the cold separator 10.

The hydrogen sulphide removal system 12 serves to increase the hydrogen concentration of the recycle gas in a conduit 13 to a level of approx. $ 85.9 · Complementary hydrogen to compensate for what is consumed in the overall process as well as dissolution losses are introduced via line 22. The hydrogen-rich phase is mixed with the predominantly liquid phase in line 6 to which 809.72 mol / hour water is added. via conduit 8, the mixture proceeding through conduit 13 into a second reaction zone 14 at a pressure of about 165 ato and a temperature of ca. 361 ° C. Reaction zone 14 contains a composite catalyst comprising 1.8 wt.% Nickel and 16.0 wt.% Molybdenum combined with a carrier consisting of 63.0 wt.% Alumina and 37.0 wt.% Silica.

1.5 vol / vol / hour. The product exit stream from the reaction zone 14 is taken out via a conduit 19 and fed into a heat separator 20 at a pressure of approx. 158 ato and a temperature of approx. 365 ° C. A predominantly vaporous phase is taken out via line 2 and recycled for combination with the charge material in line 1. A normally liquid phase is taken out via line 21 and combined with the normally liquid phase in line 18, which product from the process is sent to suitable separation facilities for extraction. of the desired fuel oil. Component analyzes of the total feed material for the reactor 1.4 (line 13) and the discharge current thereof (line 19) are given in Table III below.

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Table III

Component analyzes, reactor l4

Component Feeding material Exit current mol / hour wire 13 wire 19

Water 824.47 829.13

Hydrogen Sulfide 47.17 92.46

Hydrogen 11,703.15 11,275.72

Methane 1,367.19 1 * 379.76

Ethane 127.92 135.59

Propane 9 ^, 106.27

Butanes 51, l8 61.21

Pentans 14.43 19.32

Hexanes 8, l8 12.10

Heptane - 190 ° C 19.77 34.63 190 - 365 ° C 58.66 78.62

Light Fuel Oil 157.18 161.37

Heavy fuel oil 527.72 513.13 Total 15.001.25 14.697.30

The separation performed in the cold separator 20 is illustrated in Table IV below.

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Table IV

Component analyzes from the heat separator 20

Component mol / hour Wire 2 Wire 21

Water 829.13

Hydrogen sulfide 88.86 3 * 60

Hydrogen 10,979.32 294.40

Methane 1,338.71 41.05

Ethane 127.46 8.14

Propane 99.08 7.19

Butanes 56.09 5.11

Pentanes 17.27 2.04. Hexanes 10.55 1.55

Heptane - 190 ° C 27.79 6.97 190 - 365 ° C 38.10 40.51

Light fuel oil 78.21 83.16

Heavy fuel oil 11.97 501.16 Total 13,702.54 994.88

The total product distribution and component yields are given in Table V. These include the material withdrawn in the hydrogen sulfide removal system 12, the vent gas flow in conduit 17 and the predominantly liquid product discharge stream in conduits 18 and 21.

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Table V

Total product distribution and yields

Component Mol / hour Volume percentage

Hydrogen sulphide 30 ^> 01

Hydrogen 523> 07

Methane 128.85

Ethane 25.08

Propane 25.51

Butanes 19> 56 0.56

Pentanes 9.00 0.32

Hexanes 7.84 0.32

Heptane - 190 ° C 37> 10 2.00 190 - 365 ° C 67.23 5.71

Total fuel oil 662.97 92.52

The heptane-190 ° C fraction within the naphtha boiling range has a density of 0.765 g / cnP and contains less than 0.05 wt% sulfur. The 190-365 ° C intermediate distillate fraction has a density of 0.847 g / cm and contains less than ca. 0.1% by weight of sulfur.

The total fuel oil boiling from approx. 365 ° C, has a density of approx. 0.925 g / cm and contains 0.3% by weight of sulfur.

The foregoing description and illustrative embodiment disclose the process by which the present invention is practiced and the advantages provided by its use in producing maximum amounts of a fuel oil product containing less than about 10%. 1.5% by weight of sulfur.