the optimal operation of a slurry bubble column reactor requires that the solid phase be fluidized in the liquid phase over the entire height of the column . the solid phase is fluidized by upward forces caused by rising gas bubbles and acting against the tendency of the particles to settle under the downward gravitational force . generally , the catalyst powders employed in this invention comprise a group viii metal such as iron , cobalt or ruthenium or mixtures thereof on an inorganic oxide support . these catalysts may contain additional promoters comprising group i , group ii , group v , or group vii metals alone or in combination . the preferred catalyst powders of this invention comprise cobalt or cobalt and thoria on an inorganic oxide support containing a major amount of titania , silica or alumina . the catalyst may also contain a promoter metal , preferably rhenium , in an amount sufficient to provide a catalyst having a rhenium : cobalt weight ratio greater than about 0 . 01 to 1 , preferably 0 . 025 : 1 to about 0 . 1 to 1 . the catalyst contains about 2 to 50 wt % cobalt , preferably 5 to 20 wt % cobalt . the catalyst metals are supported on an inorganic refractory oxide comprising titania , silica or alumina . preferably , the support material is comprised of major amounts of titania and more preferably the titania has a rutile : anatase ratio of at least about 2 : 3 as determined by x - ray diffraction ( astm d2730 - 78 ), preferably about 2 : 3 to about 100 : 1 or higher , more preferably about 4 : 1 to 100 : 1 or higher , e . g ., 100 % rutile . the surface area of the preferred support is , generally , less than about 50 m 2 / gm ( bet ). cobalt - rhenium / titania catalysts exhibit high selectivity in the synthesis of hydrocarbon liquids from carbon monoxide and hydrogen . the catalysts employed in the practice of this invention may be prepared by techniques known in the art for the preparation of other catalysts . the catalyst powder can , e . g ., be prepared by gellation , or cogellation techniques . suitably , however , the metals can be deposited on a previously pilled , pelleted , beaded , extruded , or sieved support material by the impregnation method . in preparing catalysts , the metals are deposited from solution on the support in preselected amounts to provide the desired absolute amounts , and weight ratio of the respective metals , cobalt and rhenium . suitably , the cobalt and rhenium are composited with the support by contacting the support with a solution of a cobalt containing compound , or salt , or a rhenium - containing compound , or salt , e . g ., a nitrate , carbonate or the like . optionally , the cobalt and rhenium can be coimpregnated on the support . the cobalt and rhenium compounds used in the impregnation can be any organometallic or inorganic compounds which decompose upon heating in nitrogen , argon , helium or other inert gas , calcination in an oxygen containing gas , or treatment with hydrogen at elevated temperatures to give the corresponding metal , metal oxide , or mixtures of the metal and metal oxide phases , of cobalt and rhenium . cobalt and rhenium compounds such as the nitrate , acetate , acetylacetonate , naphthenate , carbonyl , or the like can be used . the amount of impregnation solution should be sufficient to completely wet the carrier , usually within the range from about 1 to 20 times of the carrier by volume , depending on the metal , or metals , concentration in the impregnation solution . the impregnation treatment can be carried out under a wide range of conditions including ambient or elevated temperatures . the catalyst after impregnation , is dried by heating at a temperature above 30 ° c ., preferably between 30 ° c . and 125 ° c ., in the presence of nitrogen , or oxygen , or both , or air , in a gas stream or under partial vacuum . the catalyst particles , if necessary , are converted to the desired particle size range of nominally 1 - 200 microns average diameter by crushing , ultrasonic treatment , or other methods known to those skilled in the art . the material can then be sieved , if necessary , to produce a powder that is predominantly within the desired particle size range . the slurry liquid used in the process is a liquid at the reaction temperature , must be relatively or largely or significantly chemically inert under the reaction conditions and must be a relatively good solvent for co / hydrogen and possess good slurrying and dispersing properties for the finely divided catalyst . representative classes of organic liquids which can be utilized are high boiling paraffins , olefins , aromatic hydrocarbons , alcohols , ethers , amines , or mixtures thereof . the high boiling paraffins include c 10 - c 50 linear or branched parafinic hydrocarbons ; the olefins include poly x - olefin liquids ; the aromatic hydrocarbons include c 2 - c 20 single ring and multi and fused ring aromatic hydrocarbons ; the ethers include aromatic ethers and substituted aromatic ethers where the ether oxygen is sterically hindered from being hydrogenated ; the amines include long chain amines which can be primary , secondary , and tertiary amines , wherein primary amines preferably contain at least a c 12 alkyl group in length , secondary amines preferably contain at least two alkyl groups being c 7 or greater in length , and tertiary amines preferably contain at least three alkyl groups being c 6 or higher in length . the slurry liquid can contain n and o in the molecular structure but noted s , p , as or sb , since these are poisons in the slurry process . representative examples of specific liquid slurry solvents useful are dodecane , tetradecane , hexadecane , octadecane , cosane , tetracosane , octacosane , dotriacontane , hexatriacontane , tetracontane , tetratetracontane , toluene , o -, m -, and p - xylene , mesitylene , c 1 - c 12 mono - and multi - alkyl substituted benzenes , dodecylbenzene , naphthalene , anthracene , biphenyl , diphenylether , dodecylamine , dinonylamine , trioctylamine , and the like . preferred liquid hydrocarbon slurry solvent is octacosane or hexadecane . ( most preferred really is hcs wax , i . e ., the product of the ft reaction .) in addition to co hydrogenation , the improved process of the present invention can be applied to hydrogenation , aromatization , hydrodesulfurization , hydrodenitrogenation , resid hydroprocessing , hydroforminglation and related reactions . these are described in more detail in “ applied heterogeneous catalysis ”, j . f . lepage et al , editions tecnip paris ( 1987 ). the height to which the catalyst can be fluidized is given by d /( u s - u l ) where d is the dispension coefficient for the particles , u s is the particle settling velocity ( see example 2 ) and u l , is the liquid velocity along the column . this liquid flow along the column may be produced by removing liquid from the top of the reactor at a rate equal to the liquid product production rate plus any extra liquid which is injected back into the bottom of the reactor as liquid recycle . while such liquid recycle can enhance the fluidization height it is usually an undesirable operating mode because the pumps and added filters greatly add to the construction and operating costs and to the complexity of operating procedures of the reactor . if this fluidization height exceeds the actual physical height of the reactor , the catalyst concentration becomes more uniform than is necessary for efficient operation . hence , excellent reactor performance can be achieved when the reactor is designed and the operating conditions are picked such that h = d /( u s - u l ) or d = h ( u s - u l ), where h is the expanded height of the liquid in the column . this ability to fluidize the particles arises from mixing induced by the gas bubbles and , therefore , at the expense of some backmixing of the reactants . plug - flow behavior can still be achieved , however , because for gas mixing the importance of the dispersion must be compared to u g , the gas velocity , rather than to the particle setting velocity u g . specifically , the degree of backmixing of the gas is measured by the peclet number , defined by pe = u g h / d . if the dispersion is maintained at that just required to fluidize the particles ( i . e ., if the reactor conditions are such that d = 0 . 5 h ( u s - u l ), then pe = u g /( u s - u l )& gt ;& gt ; 1 and plug flow behavior will prevail . in practical terms , this requires a design with intermediate values of the reactor diameter and gas velocity . low particle settling velocities increase the window of acceptable operating parameters but is not by itself a necessary conditions for good performance ( see example 7 ). it is the object of the present invention to avoid the need for substantial quantities of diluent in the reactant feed gas stream , while simultaneously operating with a well suspended catalyst bed in a plug flow mode . this will not only lead to substantially higher productivity per unit volume of reactor space , but also eliminate the need to first produce and then separate diluent from products and various reactant gas recycle streams . all of these factors lead to energy and investment savings , thereby improving the overall economic attractiveness of the process . while these are the primary constraints for efficient reactor performance , there are auxiliary conditions important in scale - up . for example , the total reactor volume must be controlled to yield the desired conversion rate of reactants to products for the intrinsic catalyst activity . this normally sets a minimum height for the reaction zone ( i . e ., minimum liquid and fluidization height ) for economic reactor operation . also , while making the catalyst particles smaller improves the fluidization it also greatly increases the difficulty in separating them from the liquid product stream . hence , catalyst particle diameters of less than 5 microns should be avoided . a more preferred diameter is greater than 30 microns . finally , if the effective reactor diameter is too small , the flow regime has large gas slugs ( i . e ., slug flow ) rising through the column rather than a dispersion of small gas bubbles typical of bubbly or churn tubulent flow which give better mass transfer performance . hence , bubble column diameters of less than 10 cm should not be used for either obtaining data for scaleup or for commercial units ( see gas - liquid - solid fluidization engineering liang - shih fan , butterworths , boston ( 1989 )). reactors with small effective diameters also can lead to unreasonable complexity and construction cost when designed for systems large enough to convert commercial feed rates of reactants . comparison of reactor volume needed for a given conversion in plug flow and well mixed reactors for isothermal reactors with constant catalyst density being supplied with reactant at a volumetric rate f , the reactor volume ( v ) required for given exit concentration or conversion is : backmixed : ⁢ ⁢ v 1 f = c o - c exit r exit = c o ⁢ j e r exit ( 1 ) plug - flow ⁢ : ⁢ - v 2 f = ∫ c o c exit ⁢ d ⁢ ⁢ c r = c o ⁢ ∫ o j e ⁢ d ⁢ ⁢ j r ⁢ ⁢ where ⁢ ⁢ r = c o - c c o ( 2 ) is the conversion , j g is the exit conversion , r is the reaction rate c is the concentration c o is the inlet concentration . assuming first order kinetics ( the hcs reaction order in total pressure is about 0 . 7 ), we find : backmixed : ⁢ ⁢ v 1 f = j e k ⁡ ( 1 - j e ) ( 1 ) plug - flow ⁢ : ⁢ - v 2 f = - 1 k ⁢ ln ⁢ ⁢ ( 1 - j e ) ( 2 ) where k is the reaction rate constant . therefore , the ratio of volume requirements for a given conversion is : v 1 v 2 = j e ( 1 - j e ) ⁢ ln ⁢ ⁢ ( 1 - j e ) as shown in fig1 , the required reactor volume is much less for the plug flow system at moderate or high conversions . a detailed model is available for calculating the effects of non - uniform catalyst distribution at intermediate mixing conditions for arbitrary kinetic expressions in specific applications . a given reactor generally exhibits behavior intermediate between plug - flow and well - mixed . for reaction rates that are first order in the concentration of reactants , the reactant concentration , c , at any vertical distance from bottom of the reactor , x , can be determined by solving the convection - diffusion equation : u g ⁢ ∂ c ∂ x - d ⁢ ∂ 2 ⁢ c ∂ x 2 = kc where u g is the average velocity of the gas along the reactor , d is the gas phase dispersion coefficient , and k is the reaction rate constant . defining z = x / h where h is the height of expanded liquid zone , this equation can be rewritten as : ∂ c ∂ z = 1 pe ⁢ ∂ 2 ⁢ c ∂ z 2 = kh u g ⁢ c where pe = hu g / d is the peclet number . as pe approaches zero the dispersion dominates and the reactor is well - mixed , while as pe approaches infinity , the dispersion becomes negligible and plug - flow behavior is achieved . this model can be solved to determine the reactor volume necessary for a given conversion as a function of peclet number . the results are plotted in fig2 . if the peclet number is above 10 , we see that effectively plug flow behavior is achieved and that the reactor volume requirements become independent of peclet number . example 2 demonstrates the importance of the peclet number in determining the reactor performance . while calculating or measuring the reaction zone height and the gas velocity is relatively straightforward , it is difficult to know a priori the dispersion coefficient . the dispersion coefficient depends on the gas throughput velocity as well as on the reactor geometry . as the reactor diameter is increased , the dispersion increases rapidly . the problem is complicated by the need to provide internal reactor structures within the reactor in order to improve heat removal . a general correlation for the dispersion coefficient as a function of geometric configuration is not possible because different internal configurations will produce different and poorly understood mixing patterns . in fact , the geometry of the internal design is the key factor the designer can use to control the mixing behavior of a slurry bubble column reactor for a given outside diameter and height that are determined by volume requirements ( i . e ., by requirements to achieve a desired conversion ). the peclet number of a given reactor can be determined from a tracer test for the reactor or a geometrically equivalent non - reactive mockup unit at milder but well defined conditions . inert gas is fed to the bottom of the reactor and after allowing the system to equilibrate a small concentration of a tracer gas is added to the feed stream as a step function ( cf . fig3 a ). then by measuring the shape of the concentration profile of the tracer in the gas outlet stream , the peclet number can be determined for the reactor by matching experimental results to model calculations . the model describing this test using a tracer that is insoluble in the liquid is : ∂ c ∂ t + ∂ c ∂ z - 1 pe ⁢ ∂ 2 ⁢ c ∂ z 2 = 0 where t is the time measured from the tracer injection multiplied by u g h / e with e being the gas holdup in the column . the peclet number is the only parameter in this equation . for infinitely large peclet numbers ( i . e ., plug - flow ), the output is a delayed step function ; and as the pe number is decreased ( backmixing increases ), the output response is more spread out in time . if the response pulse falls outside of the shaded area in fig3 b , the peclet number is less than 1 , and the reactor begins to give decreased conversion because of backmixing . the same tests can be performed using a soluble tracer , but the required model and its interpretation of results are more complicated . the required procedure is obvious to those knowledgeable of the art and can yield additional information not germane to this invention . the catalyst distribution is determined by a balance of gravitational settling of the particles and dispersion created by the upward flow of the gas bubbles . this balance results in a profile of the catalyst concentration , c p = a ⁢ ⁢ exp ⁡ [ - x ⁢ u g - u l d ] where u g is the particle settling velocity , u l is the liquid velocity along the reactor , d is the dispersion coefficient of the liquid , x is the vertical distance from the bottom of the reactor , and a is a constant that depends on the total volumetric solids concentration in the reactor . the solids concentration decreases by a factor of 2 . 7 each time the height in the slurry increases by an amount equal to d /( u g - u l ). while u g is given primarily by the composition of the liquid and by the size and density of the catalyst particles , d is governed by the effective reactor diameter and by the gas velocity . fig4 shows the decay length of the catalyst concentrations as a function of gas velocity as measured in a 6 inch diameter by 5 meter tall non - reactive bubble column using fischer - tropsch wax and titania particles . these data demonstrate a ten - fold change in bed height ( i . e ., dispersion coefficient ) over the velocity range of interest in commercial reactors ( i . e ., from 2 - 25 cm / sec ). as discussed in example 3 , the qualitative form of this curve will remain the same but the decay lengths will be different for reactors with different diameters and internal structures because of their marked effect on the dispersion coefficient . the solids distributions of glass beads and of titania particles in hcs wax were determined in a 6 ″ diameter non - reactive bubble column by taking samples from the vessel at 1 meter intervals . the temperature was 400 ° f . and the pressure was 280 psig for gas velocity below 8 cm / sec and 150 psig for gas velocities about 8 cm / sec . the decay length was obtained by taking the slope of a line segment joining the data points when plotted as the logarithm of solids concentration versus height . in fig5 , the decay length in each zone is plotted versus the average concentration in the zone for superficial gas velocities of 2 - 16 cm / sec . d / u g ( feet )= 0 . 2 ( 1 + 20c p 2 + 3000c p 4 i / u o ( cm / sec ) for u g & lt ; 4 cm / sec d / u g ( feet )= 1 . 2 ( 1 + 3c p 2 + 550c p 4 )/ u o ( cm / sec ) for u g & gt ; 8 cm / sec where c p is the volume fraction of solids in the slurry and u o is the storkes settling velocity defined by u o = 1 18 ⁢ d p 2 ⁢ p 2 - p 1 μ ⁢ g where d p is the diameter of the particle , ρ s is the density of the particle , pl is the density of the suspending liquid , μ is the viscosity of the liquid , and g is the gravitational constant . for velocities between 4 and 8 cm / sec a linear interpolation works well . this demonstrates a marked increase in the fluidization height when the solids loading is increased above 20 %. most of this increase is due to a reduction in the particle settling velocity as the solids concentration increases ( see r . h . davis and a . acrivos , annual review of fluid mechanics 17 , 19 , 1985 ). method for predicting solids distribution in a slurry bubble column reactor the data of example 5 allow one to predict the catalyst distribution in a slurry bubble column reactor via the following algorithm . suppose the reactor is charged with n species of particles each of which has a stokes setting velocity u i and average concentration c io . a given species may or may not be catalytic . the concentration of each species must satisfy the differential equation . ⅆ c i ⅆ x = - u i ⁢ f ⁡ ( u g , c p ) ⁢ c i where x is the vertical distance from the bottom of the reactor , f ( u g , c p ) is the function d / u s given in example 5 , u g is the gas velocity at height x , and c p is the local value of the total solids concentration ( i . e ., the algorithm begins by guessing the values of all the c i &# 39 ; s a the bottom of the reactor and then the equation for the c i &# 39 ; s are integrated numerically until the top of the slurry is reached . the gas velocity is computed by requiring a given overall conversion , j , and assuming the extent of reaction at any height is proportional to the fraction of the total catalyst inventory below that height . when the integration reaches the top of the slurry , the total predicted inventory for each solid species is compared to the known charge ; i . e ., we check that c io ⁢ ∫ o h ⁢ ( 1 - e ) ⁢ ⅆ x = ∫ o h ⁢ ( 1 - e ) ⁢ c i ⁢ ⅆ x where e is the volume fraction of gas ( i . e ., the gas holdup ) at vertical distance x and h is the height of the slurry . if these equalities are not satisfied , the assumed concentrations at the bottom of the reactor were not correct and they are adjusted using a newton - raphson interation technique . the equations are then integrated starting from the new values and the interations are continued until convergence is obtained . this procedure can be easily extended to include the use of liquid flow along the reactor . the right hand side of the above equations are modified by subtracting the liquid velocity from the setting velocity of species i . as pointed out above , the dispersion coefficient should be large enough to fluidize the particles adequately over the height of the reactor , e . g . : 0 . 5 ⁢ ⁢ h ≦ d u i - u l ⁢ ⁢ or ⁢ ⁢ d h ≧ 0 . 5 ⁢ ⁢ ( u g - u l ) but small enough that the gas flow remains plug flow , e . g . : . 2 ≦ pe = u g ⁢ h d ⁢ ⁢ or ⁢ ⁢ d h ≦ 5 ⁢ ⁢ u g together , these require an intermediate dispersion coefficient that satisfies the conditions : 5 ⁢ ⁢ u g ≧ d h ≧ 0 . 5 ⁢ ⁢ ( u i - u l ) ⁢ ⁢ or ⁢ ⁢ 5 ≧ d hu g = 1 pe ≧ 0 . 5 ⁢ ⁢ ( u i - u l ) u g hence for optimum performance , the reactor geometry and operating conditions must be chosen such that they lie within the shaded triangle shown in fig6 . for the catalyst / wax system being considered , the reynolds number for the settling particles is small and u s is given by stokes law times a monotonically decreasing function [ f ( c p )] which varies between 1 and 0 as the volume fraction of catalyst in the slurry is increased from 0 to the concentration at maximum packing . the “ hindered settling function ”, is described in r . h . davis and a . acrivos , annual review of fluid mechanics 17 ( 1985 ). hence , for the catalytic system of greatest interest . u s = 1 18 ⁢ d p 2 ⁢ p i - p 1 μ ⁢ gf ⁡ ( c ) = 1 . 1 × 10 - 4 ⁢ d p 2 ⁢ f ⁡ ( c ) where d p is the particle diameter , ρ s is the effective density of the particles ( approximately 2 . 7 gf / cm 3 ), ρ l is the density of the wax ( 0 . 7 gr / cm 3 ), μ is the viscosity of the wax ( 0 . 01 gr - cm / sec ), g is the gravitational constant , c is the volume fraction of solids in the slurry , and f ( c ) is 1 for a dilute slurry and a monotonic decreasing function of solids volume fraction at higher solids loading . hence , for our system . assuming the gas phase peclet number must be at least 2 to maintain plug flow behavior , the acceptable operating range is shown in fig7 for the case where u g = 5 cm / sec and no liquid recycle is used . it is clear that while the d / h design range is quite broad for small particles because of our ability to fluidize such particles without extensive backmixing , this range narrows significantly for large particles . particles with greater than 100 μm diameters cannot be effectively fluidized without a backmixing debit on the kinetic driving force . practical restrictions on the size of catalyst particles that can be effectively separated from the liquid medium by filtering decantation , or other methods considerably narrow the allowable range of particle sizes and therefore design d / h parameters .