Desulfurization Process

Abstract

In a process for hydrodesulphurization and hydrocracking of hydrocarbon oils, especially residual oils and oil fractions, with hydrogen in a fixed catalyst bed at elevated pressure and temperature, pressure drop is decreased and catalyst life improved by placing the catalyst in an annular catalyst bed and passing the reactants as a gaseous phase containing the hydrogen and a liquid phase containing the hydrocarbon oil through the catalyst bed, wherein a predominant phase of said two phases is introduced into the catalyst bed through the passage means in its outer cylindrical wall and the other phase through passage means in one of the end wall of the catalyst bed, and products are collected via the inner cylindrical wall of the catalyst bed.

Parent Case Text

This application is a continuation-in-part of Ser. No. 166,139 filed on July 26, 1971 now abandoned.

Description

The present invention relates to a process for reducing the sulphur content of heavy hydrocarbon oils such as residuum oils and crude oils by contacting a two-phase mixture of liquid hydrocarbon oils and gaseous hydrogen with a catalyst arranged in a fixed bed reactor at elevated temperature and pressure. In addition, hydrocracking of the oils to produce fractions boiling lower than the feedstock is achieved.

Depending on their origin, heavy petroleum oils contain sulphur compounds of different types. Thiols, sulphides, and disulphides predominate in the lower boiling fractions while thiophenic compounds predominate in the higher boiling fractions. In cracked petroleum oils and residual oils the sulphur compounds are in the condensed cyclic form, for instance benzothiophenes. In general, the resistance of sulphur compounds to the hydrodesulfurization process increases in the following order: thiophene, 2-ring thiophene, and thiophene compounds with 4 or more rings.

Other undesirable compounds found in heavy petroleum oils which are reduced in the hydrodesulfurization and hydrocracking process are nitrogen and oxygen compounds, commonly heteroatomic compounds.

In addition, the presence of metallic contaminants in the petroleum oils is of significant importance in the hydrodesulfurization process. The most common metals are nickel and vanadium, but also iron, copper, sodium, calcium and zinc are often present. The metals may occur as chlorides, oxides or sulphides. However, they are more often in the form of metallo-organic complexes such as porphyrins and their derivatives. These complexes are associated with asphalthenes which are non-distillable and therefore concentrate in residual oil fractions. During the hydrodesulfurization process the metal compounds are liberated from the organic compounds and most often deposited in the catalyst pores and on the outside of the catalyst particles. At the same time the high molecular weight asphalthenes coagulate and polymerize and thereby finally lead to the formation of coke deposits inside or outside the catalyst particles.

Because the sulphur and nitrogen present in petroleum fuel oils will eventually appear in the flue gases as oxides, they seriously contribute to the problems of air pollution. Therefore, processes for reducing the nitrogen and particularly the sulphur content in residuum oils and crude oils have become of increasing commercial importance. Among these processes, the hydrodesulphurization process is one in which the more resistant sulphur compounds can be removed and thereby converted to easily reclaimed sulphur compounds.

In the prior art processes of hydrodesulphurization and hydrocracking the hydrocarbon oils are reacted with hydrogen in the presence of a catalyst. In order to avoid rapid catalyst deactivation by the heavier hydrocarbons, a considerable excess of hydrogen has to be used. While the consumption of hydrogen is generally of the order of 35-350 normal liter per liter oil (Nl/l oil), the feed ratio of hydrogen to hydrocarbon oil has to be of the order of at least 150-6,000 Nl/l oil, preferably between 500 and 2,000 Nl/l oil. This is achieved by recycling the unreacted hydrogen from the product stream and adding fresh hydrogen before reintroducing it into the reactor.

The pressure at which hydrodesulphurization and hydrocracking processes are operated depends on the nature and origin of the hydrocarbon oils. While low-boiling fractions such as gas oils can be satisfactorily hydrotreated at a pressure in the range from 25 to 70 atms. abs., heavier fractions such as whole crude oils and residuum oils require a pressure up to 600 atms. abs., preferably 50 to 200 atms. abs. Also the temperature depends to some degree on the hydrocarbon oils, normally it is within the range 260.degree. to 485.degree.C, preferably 315.degree. to 430.degree.C for whole crudes and residuum oils.

By proper choice of catalyst and process conditions, the degree of hydrodesulphurization relative to the degree of hydrocracking can be varied.

In conventional hydrodesulphurization and hydrocracking processes, the catalyst is arranged in a fixed bed reactor through which the reactants flow axially downwards. Usually, the operating conditions are such that only the lighter hydrocarbons of the oil are vaporized, while the major part of the hydrocarbon oil is present as a liquid. Therefore, there is a mixed flow of two phases through the catalyst bed, a gaseous phase consisting of hydrogen and the lighter part of the hydrocarbons and a liquid phase containing the remaining heavier hydrocarbons. The liquid phase is thus dispersed on the surface of and inside the catalyst particles while it flows downwards under the influence of gravity. The gaseous phase passes through the interstitial voids between the wetted catalyst particles. Such a flow arrangement is known as a trickle flow. One disadvantage of this system is that the two phases have to pass a long way through the bed which results in a high pressure drop across the catalyst bed.

In the catalyst bed, demetallation reactions leading to liberation of metals from the metallo-organic compounds predominate at the top of the catalyst bed where the fresh hydrocarbon oil is introduced. Initially, the desulphurization and denitrification reactions take place throughout the entire catalyst bed. However, the top of the catalyst bed gradually becomes catalytically inactive for these reactions as it is being fouled by deposits. These deposits, consisting mainly of metal compounds and carbonaceous materials, accumulate in the interstitial voids, the volume of which is thereby reduced. Therefore, the free passage of the reactants and products is gradually hindered and eventually the increased pressure drop makes further operation of the process impossible. This is an additional disadvantage of the prior art processes, since the catalyst bed has to be subjected to a periodic cleaning or regeneration, and in some cases the catalyst has to be replaced either completely or partly.

In addition to the interstitial deposition of solids there is also a deposition, particularly of metals and metal compounds in the pores, inside the catalyst particles and close to their outer surface. This deposition which eventually renders the interior part of the catalyst particles inaccessible to the reactants cannot be avoided, but attempts at reducing the harmful effect have often been made by decreasing the catalyst particle size. However, for economic and technical reasons the particle size used in conventional fixed bed reactors with axial flow is limited to at least 1 mm.

An object of the present invention is to provide a process for catalytic hydrodesulphurization and hydrocracking of hydrocarbon oils in which both the pressure drop and the rate of increase of the pressure drop with time on stream are lower than in the conventional fixed bed processes with axial flow.

Another object of the present invention is to provide a process in which catalyst stability and life are improved compared with a conventional process.

According to the present invention there is provided a process for the hydrodesulfurization and hydrocracking of heavy hydrocarbon oils including crude oils and heavy residual petroleum fractions at elevated temperature and pressure in a fixed catalyst bed, in which two phases, one a gaseous phase containing hydrogen and the other a liquid phase containing the heavy hydrocarbon oil, are passed continuously through an annular fixed catalyst bed enclosed between substantially vertical cylindrical coaxial walls with passage means for liquid and/or gases and top and bottom walls with or without passage means for liquids and/or gases, the axial height of said annular bed being at least twice its width, at least one of said two phases being introduced through said passage means in the outer cylindrical wall of the said catalyst bed and substantially all of the products being withdrawn from the catalyst bed through said passage means in the inner cylindrical wall of the catalyst bed.

The width of the catalyst bed is defined as the radial distance between the inner and the outer cylindrical wall of the catalyst bed. The catalyst bed is preferably circular-cylindrical, but the cross-section of the walls may deviate from the circular-cylindrical form and, for instance, may be elliptical. In that case the width of the catalyst bed is defined as the distance between the walls, measured perpendicularly to the tangents of the cross-section. The catalyst bed will be enclosed in a pressure and temperature resistant reactor vessel. The outer cylindrical wall of the catalyst bed and its end walls may be in a spaced relationship to the cylindrical outer wall and the end walls, respectively, of the reactor vessel; or the reactor vessel walls may constitute the outer cylindrical wall and/or the top end wall and/or the bottom end wall of the catalyst bed. The said walls may be flat but are preferably arched.

When introduced into the reactor, the hydrocarbon oil will normally be preheated.

The catalyst bed used in the process of the present invention has an annular form which ideally is limited by two coaxial circular cylindrical surfaces and two parallel planes at right angles to the cylinder axis. Normally, the total height of the catalyst bed will be several times its outer diameter. Characteristic features of the invention are: (1) the reactants are a two-phase mixture of a gaseous hydrogen phase and a liquid hydrocarbon oil phase, (2) at least one of the two phases is introduced into the catalyst through the outside cylindrical surface of the annular catalyst bed, and (3) the products are collected through the central opening extending through the catalyst bed. In this way, the distance that the reactants have to pass through the catalyst bed has been greatly reduced, while still the time of contact between the reactants and the catalyst is high.

In the ideal case, a phase that is introduced through the outside cylindrical surface would pass strictly radially through the catalyst bed, whether it be a gaseous phase or a liquid phase. In practice, there will be deviations from this ideal situation since there is in general a tendency for a liquid phase to obtain an axial gradient of downward movement under the influence of gravity in addition to the radial gradient of inwards movement. Similarly, there is for the gaseous phase, which is present together with the liquid phase, a tendency to obtain an axial gradient of upwards movement in addition to the radial gradient of inwards movement. Which of the gradients of movements will predominate in an actual situation for each of the two phases will among other factors depend on the actual ratio of volumetric flow rates of gaseous phase to liquid phase.

An approximate value of the actual ratio of volumetric flow rates of gaseous phase to liquid phase can be calculated from the feed rates of hydrogen and hydrocarbon oil. For example, 540 Normal liters of hydrogen per liter of hydrocarbon oil at 100 atms. and 350.degree.C will correspond to 12 volumes of gaseous hydrogen per one volume of liquid hydrocarbon oil. However, this is only the approximate ratio of gaseous phase to liquid phase flow rate, since part of the hydrogen will dissolve in the liquid phase particularly at high pressures. On the other hand, lighter hydrocarbons of the hydrocarbon oil may evaporate and thus add a little to the gaseous phase volume. It is possible in each particular case to correct the calculated approximate ratio of gaseous phase to liquid phase accordingly and thus obtain the actual ratio flowing through the catalyst bed. The ratios which are mentioned in the following are assumed to be actual values corrected for dissolved gases, and they will be referred to as gas to liquid ratio.

In the process of the present invention the gas to liquid ratio may vary within wide limits. In order to illustrate qualitatively how the flow pattern in the catalyst bed will depend on this ratio, two extreme cases will be considered.

In one extreme case, it is assumed that the gas to liquid ratio is low, for example from about 0.05:1 to about 2:1. In this case, there will be a continuous liquid phase in the catalyst bed with relatively few gas bubbles. The liquid phase introduced through the outside cylindrical surface and evenly distributed over this surface will flow practically in the radial direction inwards through the annular catalyst bed. Regardless how the gas phase is introduced, the relatively few gas bubbles will flow practically axially upwards through the bed under the influence of gravity.

In another extreme case in which it is assumed that the gas to liquid ratio is relatively high, for example from about 30:1 to about 150:1, there will be a continuous gaseous phase in the catalyst bed. When introduced evenly distributed over and through the outside cylindrical surface of the annular catalyst bed this gaseous phase will flow practically in the radial direction inwards through the catalyst bed. Regardless how the liquid phase is introduced, the radial movement of the gaseous phase will tend to "blow" the relatively few liquid droplets in the radial direction. However, the influence of gravity will be strong enough to cause an additional axial and downward gradient of movement of the liquid phase in this case.

Those skilled in the art will understand that the flow pattern schematically described hereinabove will change gradually from one extreme to the other, when the gas to liquid ratio is changed from the lower extreme value to the higher extreme value. It will similarly be understood that in the extreme cases radial flow will predominate for one of the phases. In this specification it will be understood that the liquid phase is the predominant one when the gas to liquid phase is below about 2:1, while the gaseous phase is the predominant one when the gas to liquid ratio is higher than about 30:1. However, this higher limit cannot be very well defined, since it will change with the precise flow arrangements and it may be possible to make such arrangements that the gaseous phase is the predominant one, even if the gas to liquid ratio is as low as 10:1.

The flow restrictions in the catalyst bed will always be more serious for the predominant phase. However, this phase will always pass radially through the annular bed and thus have to move through the shorter distance. Therefore, the main objective of the present invention, a reduction of the total pressure drop, has been achieved.

However, the more the actual situation in the catalyst bed resembles one of the extreme cases referred to above, the less satisfactory and intimate is the mixing of the two phases in the case when both phases are introduced evenly distributed over the outside cylindrical surface of the catalyst bed. This is because of the tendency to separation of the two phases that is a result of the additional axial gradient of movement of one of them, namely the one that is far from being the predominant one. However, it is within the scope of the present invention to compensate for this tendency by altering the rates at which the two phases are introduced into the catalyst bed at various positions.

Thus, when the gas to liquid ratio is low and consequently the gas bubbles move almost entirely upwards through the catalyst bed, there will be a relatively higher concentration of gas phase in the upper part of the catalyst bed unless a greater part of the gas phase is introduced closer to the bottom of the catalyst bed. Therefore, in this case the rate of hydrogen introduction through the outside cylindrical surface should be increased with increasing distance from the top of the bed, e.g., by having the gas inlets near the bottom situated closer to each other in the outer cylindrical wall compared to those at the top of the said wall. It may even be desirable to introduce all the hydrogen at a position close to the bottom or through the bottom.

Similarly, when the gas to liquid ratio is high and, consequently, the liquid phase tends to move axially downwards, there will be a scarcity of liquid phase in the upper part of the catalyst bed, particularly at some distance from the outside surface. It is possible to compensate for this scarcity by introducing all or part of the liquid phase through the top of the catalyst bed.

In both cases one of the phases, namely the one that is not the predominant one, will have to move over a longer distance through the catalyst bed, namely almost entirely in the axial direction. However, this does not significantly increase the total pressure drop, since the major contribution to pressure drop relates to the other phase, namely to the predominant phase. Therefore, both kinds of variations are within the scope of the present invention.

The preferred operating pressure for the process is in the range 50-200 atms. abs., and the preferred temperature is in the range 315.degree.-430.degree.C.

The catalyst used in the process of the invention may be of any conventional type used commercially in hydrodesulphurization and hydrocracking of hydrocarbon oils. Such catalyst generally contain oxides of nickel, cobalt, molybdenum, and/or tungsten, normally carried on a support material. The support material is a highly porous oxidic material such as alumina. Other types of support materials belong to the group of zeolites. Still other types of support materials can be used. The catalyst is used in the form of discrete particles of regular or irregular shape. Several factors have to be taken into account in the selection of particle size and shape, such as the allowable pressure drop, the desired degree of desulphurization and the type of hydrocarbon oil.

The catalyst particle size is in the range 0.1-10 mm, preferably 0.2-1 mm. The particle shape may be a conventional shape such as a cylinder or a sphere. However, catalyst shapes having a much higher surface to volume ratio are preferred. Such catalyst shapes are irregularly shaped particles, for example crushed or flake-shaped catalysts.

The catalyst is arranged in an annular bed which may be contained in a basket having inside and outside cylindrical surfaces. At least one of the two phases, the gaseous hydrogen or the liquid hydrocarbon oil phase, is introduced into the catalyst bed through the outside cylindrical surface. This introduction may take place through a system of nozzles or orifices which are designed to disperse one or both phases in the desired manner through the catalyst bed. Alternatively, the outside cylindrical surface of the catalyst basket may be a perforated plate or a metal screen which is surrounded by a greater diameter vessel so that there is an annular space between the reactor vessel and the catalyst basket. This latter alternative is particularly suitable when the gas to liquid ratio is low, i.e., when the liquid phase is the predominant one, since the liquid phase will then fill the annular space around the perforated basket and successively flow through the catalyst bed in a radial direction. In this case the gaseous hydrogen phase may be introduced through pipes connected to or extending into the catalyst basket.

It may be desirable to vary the gas or liquid inlet rate continuously from bottom to top of the catalyst bed. Thus, when the gaseous phase is the predominant one it may be advantageous to place the nozzles, orifices or inlet apertures according to their individual capacities (which may vary, since they may have varying effective passage area) more or less densely in such a manner that the gas passage capacity per unit area of the outer cylindrical catalyst bed wall increases gradually or stepwise from the upper to the lower end of said cylindrical wall; in that case the inlet orifices, nozzles or apertures for the liquid are preferably all situated at the top end wall of the bed.

Conversely, when the liquid phase is the predominant one, the nozzles, orifices or inlet apertures for the liquid may be positioned more or less densely in accordance with their individual passage capacity in such a manner that the liquid passage capacity per unit area of the outer catalyst bed wall decreases gradually or stepwise from the upper to the lower end of said cylindrical wall; in that case the inlet orifices, nozzles or apertures for admitting the gas phase are preferably all positioned either in the bottom end wall of the bed or quite near thereto.

The main advantage of the process of the present invention compared to hitherto used processes for hydrodesulphurization and hydrocracking of hydrocarbon oils is related to the predominantly radial flow of at least one of the two phases. The total pressure drop over the catalyst bed is primarily associated with the flow of the predominant phase which is introduced at the higher volume rate. Since this phase is flowing predominantly radially, a considerable reduction of the overall pressure drop has been achieved by the present invention. In contrast to this flow pattern, both phases in conventional hydrodesulphurization processes flow axially through the catalyst bed and thus have a much longer distance to pass through.

In hydrodesulphurization processes, the rate of feeding the hydrocarbon oil to the catalyst bed is usually defined by the liquid hourly space velocity (LHSV), which is the volume of liquid hydrocarbon oil per volume of reactor per hour. For the process of the present invention, the LHSV will be similar to the LHSV for a conventional hydrodesulphurization process. Consequently, the LHSV will be in the range 0.1 to 10, preferably 0.3 to 3.0. The term "reactor volume" used in this specification is understood to be the volume actually occupied by the catalyst bed.

The rate of hydrogen introduction is defined by the hydrogen to hydrocarbon oil ratio. This ratio is usually expressed as the normal volume of hydrogen (i.e., gaseous volume at 0.degree.C and at atmospheric pressure) per liquid volume of hydrocarbon oil. In a conventional hydrodesulphurization process, the hydrogen to oil ratio is normally of the order from 150 to 5,000 normal liter H.sub.2 per liter oil (Nl/l oil). In the process of the present invention it may be within the same range. The actual consumption of hydrogen depends particularly on the type of hydrocarbon oil, but typically it may be from 30 to 300 Nl/l oil. However, a certain excess of hydrogen is required to eliminate carbon deposition on the catalyst. Part of the hydrogen will become dissolved in the hydrocarbon oil. The solubility which is much dependent on operating pressure and temperature, will typically be of the order 8.0 to 34 Nl/l oil. Under the preferred process conditions, excess of hydrogen required to suppress carbon deposition is higher than the amount that can be dissolved in the hydrocarbon oil and there will always have to be a gaseous phase present in any part of the catalyst bed.

Thus, of the total amount of hydrogen introduced into the catalyst bed together with the hydrocarbon oil, a part will become dissolved in the oil, another part will be consumed by the hydrodesulphurization and hydrocracking reactions, while a third part remaining as a free gaseous phase will serve to prevent carbon deposition. The amount of hydrogen required for each of the three purposes will vary considerably, particularly with the operating temperature and pressure and with the type and required degree of conversion of hydrocarbon oil. However, as exemplified previously in the present specification, it is possible to make an approximate evaluation of the actual volume of the gaseous phase present in the catalyst bed. This will be further illustrated in the examples given later on in the specification (particularly in Example 2).

One further advantage of the radial flow process of the present invention is that much smaller catalyst particles can be used compared with the conventional axial flow process, while still keeping the pressure drop low. In the conventional process, the size of catalyst particles is limited to at least 1mm. In the radial flow process the particle size can be as low as 0.2mm or even less and still the pressure drop will be low compared with the conventional process. In the smaller catalyst particles a greater portion of a particle is accessible to the reactants, whereby a higher catalytic activity is achieved. Furthermore, when small particles are used, the deposition of metal sulphides and coke in the pores close to the outer surface of the catalyst particles will take place in a much larger portion of the total catalyst volume. The harmful effect of deactivation is thereby much delayed, and hence also an improved catalyst life is achieved in the present invention.

A major problem arises in conventional axial flow reactors during hydrogen processing of residual and crude oils owing to the deposition of iron, nickel and vanadium sulphides also in the interstitial void space between the catalyst particles together with deposition of coke. The iron, nickel and vanadium are contained in organometallic constituents of the oil. The consequence of the interstitial deposition is that the pressure drop across the axial flow reactor soon rises beyond tolerable commercial limits. It is a further advantage of the radial flow process of the present invention, particularly in the embodiment in which the hydrocarbon oil is introduced through the cylindrical surface of the catalyst basket, that the deposition of solids in the interstitial voids becomes less harmful, because they are being spread over a large area. Therefore, the rate of pressure drop buildup in the hydrodesulphurization process of the present invention is low compared to the corresponding rate of pressure build-up in the conventional hydrodesulphurization process. In consequence, the useful life time of the catalyst will be increased.

In order that the invention should be better understood, it will now be described in further detail in the following Examples 1 and 2 and with reference to the attached drawings.

In the drawings,

FIG. 1 shows an axial section of a first embodiment of a reactor to be used in the process of the invention, and

FIG. 2 an axial section of a second embodiment of a reactor to be used in the process of the invention.

The figures will be explained in detail in the respective examples.

EXAMPLE 1

This example illustrates on the basis of design data that the useful life of a desulphurization catalyst is the longer, the smaller the particle size of the catalyst. The example will furthermore illustrate that the pressure drop for a given particle size is very much smaller in the radial flow process of the invention than in the conventional axial flow process. Consequently, it is possible to reduce the catalyst particle size in the process of the invention, thus achieving an increased catalyst life. The design data given in Table I (first columm) relate to a hydrodesulphurization process with radial flow in accordance with the invention, while design data for a conventional axial flow process are given (for comparison) in the same Table (second column). Under these conditions a Kuwait Long Residue oil containing 4 wt% sulphur is desulphurized to 2 wt% sulphur, while the original content of metals in the oil, 0.0065 wt% Fe + Ni + V, is reduced to about the half.

In the conventional axial flow process, both the liquid phase and the gaseous phase flow axially downwards through the catalyst bed. In the radial flow process of the invention conducted under the conditions described in Table I, both phases are introduced through the outer cylindrical wall and pass more or less completely radially through the annular bed from where they are collected through the inner cylindrical wall.

Two cases are considered for each of the two process of Table I, both carried out with spherical catalyst particles. In one case, case A, the catalyst particle size is 1.7 mm, while in the other case, case B, the catalyst particle size is 0.5 mm. Table II gives for both cases the initial pressure drop and the metal loadings which is the amount of metals (Fe + Ni + V) finally deposited on the catalyst. The pressure drop was calculated by well known methods and found to be in accordance with experimentally determined values. Values for the metals loading are known from the conventional desulphurization process, and it is assumed that the same values will be obtained in the process of the invention for a given catalyst particle size. The catalyst utilization, which is the amount of oil that can be processed per unit volume of catalyst, can be derived from the figures for metals loading and for amount of metals removed from the oil.

Table I ______________________________________ Design Data ______________________________________ Radial Axial Flow Flow Process Process ______________________________________ Oil feed rate m.sup.3 /day 4,100 4,100 Hydrogen feed rate Nl/l oil 825 825 Pressure atm. abs. 100 100 Temperature .degree.C 370 370 Gas to liquid ratio m 20:1 20:1 Height of catalyst bed m 20 20 Outer diameter of m 3.40 3.30 catalyst bed Inner diameter of m 0.85 0.00 catalyst bed Volume of catalyst bed m.sup.3 170 170 ______________________________________

Table II ______________________________________ Radial Flow Axial Flow Process Process ______________________________________ Case A: Catalyst particle size mm 1.7 1.7 Initial pressure drop atm. 1.1 .times. 10.sup.-.sup.3 0.5 Metals loading, 0.23 0.23 g Fe+Ni+V/ml cat. Catalyst utilization, 7.7 7.7 m.sup.3 oil/l cat. Case B: Catalyst particle size mm 0.5 0.5 Initial pressure drop atm. 1.1 .times. 10.sup.-.sup.2 1.7 Metals loading, 0.78 0.78 g Fe+Ni+V/ml cat. Catalyst utilization, 26.0 26.0 m.sup.3 oil/l cat. ______________________________________

As can be seen from Table II, the pressure drop for a given catalyst particle size is very much smaller in a radial flow process than in an axial flow process. It is furthermore seen from the table that the metals loading increases with decreasing particle size. This is because the metals deposit inside the catalyst particles close to the outside surface of each particle. Consequently, the smaller particles can accept more metal per unit volume of catalyst bed than can the larger particles. The catalyst utilization or catalyst life is correspondingly improved from 7.7 m.sup.3 oil per liter catalyst of 1.7 particle size to 26.0 m.sup.3 oil per liter catalyst of 0.5 mm particle size. In other words, more than three times as much oil can be processed on the smaller catalyst particles than on the larger ones before they have obtained the maximum amount of metals. However, the use of the smaller catalyst particles size is only possible in the radial flow process, because otherwise the pressure drop will be too high. In case B (0.5 mm catalyst particle size), the critical pressure drop was found to be 1.7 atms., which is commercially unacceptable, because it will soon exceed the maximum allowable value as a result of interstitial deposition of metals and coke.

FIG. 1 illustrates the design of a reactor for conducting the radial flow process for which design data are given in Table I. The gas to liquid feed ratio is about 20:1. Since some hydrogen is consumed by the reaction, the ratio will be a little lower at the outlet than at the inlet. However, throughout the bed, the ratio is between the described two extreme ranges, from 0.05:1 to 2:1 and from 30:1 to 150:1, for which reason it is proposed in the present specification that only one of the two phases should pass radially through the annular catalyst bed. Therefore, both the liquid oil phase and the gaseous hydrogen phase can be introduced through atomizing nozzles distributed on the outside cylindrical surface of the bed.

An annular space 11, is substantially filled with a bed of cobalt-molybdenum oxide catalyst (not shown) on an alumina support having a particle size of 0.5 mm (case B, Table II). The space having a height of 20 meters and an outer diameter of 3.40 meters, is contained in a circular-cylindrical pressure vessel 12 with a removable cover 13. A central tube 14 extending axially through space 11 has a diameter of 0.85 meter. It rests loosely on the bottom of the pressure vessel 12 and is kept in its central position by an annular bottom wall 15 and an annular top wall 16. The bottom wall rests on a ring 17 secured to the inside of the pressure vessel and on a ring 18 secured to the outer surface of the central tube 14. A sufficient gas tightness at these places is ensured by appropriate sealings. The slightly conical top wall 16 is kept between flanges of the vessel 12 and the cover 13.

The liquid hydrocarbon oil and hydrogen are both introduced into the catalyst bed through atomizing nozzles 19 distributed on and extending through the cylindrical surface of vessle 12. In order to simplify the drawing, only a few of these nozzles have been shown. The central tube 14 is provided with apertures 20 through which the liquid and gaseous products leave the catalyst bed. Only a few of these apertures are shown in the drawing. Since part of the gaseous hydrogen phase may move to the top of the catalyst bed, there is a clearance 21 between the top wall 16 and the central tube 14 so that the gas can escape this way. The slightly conical form of the top wall ensures that there will be no permanent accumulation of gas in the top of the catalyst bed.

The liquid oil phase and the gaseous hydrogen phase will be separated in the central tube 14 where the liquid phase will move toward the bottom and leave pressure vessel 12 through a tube 22. Similarly, the gaseous phase will move upwards and leave the pressure vessel through a tube 23 on the cover 13.

In order to compensate for the tendency of the gaseous phase to move upwards and of the liquid phase to move downwards in the catalyst bed, the gas to liquid ratio of the feed introduced through the nozzles is adjusted accordingly, so that the ratio is somewhat increased near the bottom and decreased somewhat near the top of the catalyst bed, the average gas to liquid ratio still being 20:1.

EXAMPLE 2

A Kuwait Long Residue oil containing 4 wt% sulphur is desulphurized to 1 wt% sulphur at a temperature of 350.degree.C and under a pressure of 150 atms. abs.. At these conditions, the solubility of hydrogen in the hydrocarbon oil is 17 Nl/loil, while the consumption of hydrogen for the hydrodesulphurization and hydrocracking reactions is abt. 85 Nl/l oil. The total amount of the hydrogen introduced together with the hydrocarbon oil is 153 Nl/l oil. Consequently, there is available as free gaseous hydrogen in the feeds 136 Nl/l oil and in the products 51 Nl/l oil. Since 1 Normal liter of gas at 350.degree.C and 150 atms. abs. will be reduced to 0.015 liters, the gas to liquid ratios as previously defined will be 2:1 and 0.8:1, respectively.

FIG. 2 illustrates how the desulphurization process can be carried out under these conditions in accordance with the present invention. Since the gas to liquid ratio is low (between about 2:1 and 0.8:1), the liquid phase is introduced through the outside cylindrical surface of the catalyst bed and passes radially through the bed, while the hydrogen mainly is introduced near the bottom of the catalyst bed. Some hydrogen may also be introduced in the form of hydrogen dissolved in the liquid hydrocarbon oil.

An annular space 31 for a catalyst bed (not shown) has a height of 20 meters, an outer diameter od 3.40 meters and an inner diameter of 0.85 meter, which gives a catalyst bed volume of 170 m.sup.3. This catalyst bed space is contained in a pressure vessel 32 having a removable cover 33. The vessel is provided with pipes and auxiliaries for introducing the reactants and collecting the products. It is furthermore provided with the necessary means for controlling the operating conditions such as temperature, pressure and flow rates. These additional means and auxiliaries are not shown in the drawing.

The catalyst bed space 31 is kept between outer and inner cylindrical walls. The outer cylindrical wall 34 is a perforated plate. Its diameter is smaller than the diameter of the pressure vessel 32, so that it is surrounded by an annular space 46 having a width of a few cms. The inner cylindrical wall 35 is a perforated tube. Furthermore, the catalyst bed is supported by a bottom wall 36 resting on a ring 37 secured to the side of the pressure vessel 32. There is an appropriate sealing material between the bottom wall 36 and the ring 37. The bottom wall 36 is furthermore secured gas-tight to and thus carried by the inner cylindrical tube 35 resting loosely on the bottom of the pressure vessel 32. Finally, the catalyst bed has a removable cover 39.

The gaseous hydrogen phase is introduced through perforated pipes 40 extending into the bottom of the catalyst bed through the vessel and through the cylindrical wall of the catalyst bed. The liquid hydrocarbon oil is heated and introduced into the annular space 46 surrounding the catalyst bed through one or more tubes 41. The oil may in advance have been completely or partly saturated with dissolved hydrogen. From the annular space 46, the liquid oil enters the catalyst bed 31 through apertures 42, of which only a few are shown in the drawing, passes radially through the bed in space 31, and is collected in central tube 35. This tube or inner cylindrical wall is provided with apertures 43 distributed so that an almost equal distribution of the oil flowing through the catalyst bed is achieved. Only a few of these apertures are shown in the drawing.

In the central tube 35, the two phases separate. The liquid oil phase moves downwards and leaves the pressure vessel through tube 45, while the gaseous hydrogen phase leaves the vessel through tube 44.

The arrangement of the pipes 40 for hydrogen introduction may not always be practical, since they require pressure-resistant packing 47 through the vessel 32 and have to pass through holes 48 into the catalyst bed 31. In an alternative arrangement, the gaseous hydrogen phase may be introduced into the empty space 38 under bottom wall 36, from where it enters the catalyst bed through apertures (not shown) in the bottom wall 36. In this arrangement, the central tube 35 rests on a gas-tight packing on the bottom of the pressure vessel 32.

Claims

1. A process for the hydrodesulphurization and hydrocracking of heavy hydrocarbon oils including crude oils and residual petroleum fractions at elevated temperature and pressure in a fixed catalyst bed, comprising the steps of

a. continuously passing at elevated temperature and pressure a gaseous phase containing hydrogen through the fixed catalyst bed,

b. simultaneously therewith passing at elevated pressure and temperature a liquid phase containing the heavy hydrocarbon oil through the fixed catalyst bed, said catalyst bed being placed in an annular space between upper and lower end walls and substantially vertically and coaxially disposed, substantially cylindrical walls, at least said cylindrical walls having passage means for fluids, the axial height of said annular space being at least twice its width as measured radially between the two cylindrical walls, wherein a predominant phase of said two phases is introduced into the catalyst bed through the passage means in its outer cylindrical wall and the other phase through passage means in one of the end walls of the catalyst bed,

c. recovering the product hydrocarbon oil via the passage means in the inner cylindrical wall of the catalyst bed, and

d. removing gases formed in the process and unreacted feed gases via passage means in any wall of the catalyst bed other than the bottom end

2. The process of claim 1, wherein the liquid hydrocarbon-containing phase is introduced into the catalyst bed through the passage means in its outer cylindrical wall, and the gaseous hydrogen-containing phase through passage means situated at selected positions in its lower end wall and the

3. The process of claim 2, wherein the gas to liquid ratio is in the range

4. The process of claim 1, wherein the liquid hydrocarbon-containing phase is introduced into the catalyst bed through passage means situated at selected positions in its top end wall and the upper part of its outer cylindrical wall, and the gaseous hydrogen-containing phase through the

5. The process of claim 4, wherein the gas to liquid ratio is in the range 10:1 to 150:1.