Patent Application
20100317907
This invention relates to methods for the low cost, accelerated development of plug flow catalytic hydroprocessing processes, particularly for hydroprocessing clean feeds, from discovery to commercial readiness.
The term “hydroprocessing” encompasses a large number of catalytic processes involving the hydrogen treatment of hydrocarbon streams such as hydrocracking, hydroisomerization, hydrotreating, hydrogenation, ring opening and others. For instance, the catalytic hydrocracking of paraffins involves the conversion of linear alkanes to mixtures of lower molecular weight molecules, a large fraction of which are branched isomers. In addition to the primary desired reaction, secondary reactions also occur during hydroprocessing that generate undesired species such as unsaturates, aromatics and polynuclear aromatics (PNA's) that tend to deactivate the hydroprocessing catalyst. By “clean feed” is meant a hydrocarbon feed stream that does not contain any significant amounts of deactivating species such as aromatics, PNAs or other unsaturates. Such streams may be, e.g., a slack wax, a Fischer-Tropsch wax or liquid, or another refinery stream that has been treated to remove catalyst deactivating species. Thus, in the hydroprocessing of clean feeds, the primary catalyst deactivating species are self generated by such secondary reaction paths. The ability to minimize the rate of catalyst deactivation and the production of undesired products by secondary reactions during hydroprocessing can have a substantial effect on the economic viability and profitability of a commercial scale hydroprocessing process.
In order to scale up a plug flow catalytic process, it is necessary to define the impact of time on stream, residence time, catalyst particle size, shape and other characteristics, and temperature profile on reaction rate and selectivity. The first step in a traditional scale-up program generally involves the selection, and definition of the intrinsic properties of, the catalyst. This step is typically performed isothermally with a diluted, crushed or powdered catalyst to minimize mass transfer limitations. A process variable study is performed to determine the impact of space velocity, pressure, and residence time on reaction rate and selectivity. Activity and selectivity maintenance are then determined over a six to twelve month operating period. At the end of the operation, a second process variable study is performed to determine whether these properties have changed during time on stream.
Next, a commercial form of the catalyst is tested in an isothermal reactor. The commercial catalyst is of a larger particle size than the crushed catalyst and may have a special shape to minimize pressure drop during operation. The larger particle size generally results in a lower reaction rate and a selectivity loss due to limitations on mass transfer of reactants or products in and out of the catalyst pores. Operations generally consist of performing process variable studies at the beginning and end of an activity and selectivity maintenance run. This operation can be run in a laboratory scale reactor and typically lasts approximately one year.
The final step in the scale-up process is to test the commercial catalyst under adiabatic conditions, normally in a demonstration scale reactor containing one or more reactor tubes. The tubes in the demonstration scale reactor would have internal diameter of approximately 1 inch. In some cases, to further explore heat transfer effects, a configuration containing up to about 6-8 tubes arranged at commercial spacing could be used. In an exothermic reaction, the temperature profile depends upon the degree to which heat is continuously removed, as in a tubular reactor, or the reactor is simply a fixed bed reactor without a specific heat removal capability. The temperature profile can have a significant impact on selectivity, reaction rate, and activity maintenance. The test run also provides a measure of the tendency for the catalyst to produce hot spots or temperature runaways. Here again, the operating period can exceed one year.
This sequential approach typically takes in excess of three years to complete and may not provide all of desired data for scale-up. For many catalysts, the reaction rate and selectivity may be a function of residence time as well as time on stream. This can be the result of changes in the catalyst state or form, due to exposure for extended periods of time, or it may be due to the changing gas and liquid composition from the reactor inlet to the outlet. Examples would include oxidation from water formed during conversion, formation of a support over layer, poisoning, e.g., by reaction with hydrogen sulfide and ammonia, etc. In addition, surface catalytic reactions and buildup of feed and products in the pores can result in reductions in mass transfer rate to the catalyst.
More recently, High Throughput Experimentation (HTE) techniques have been proposed as a source of data for new catalysts and processes. These HTE experiments are normally performed under conditions that minimize heat and mass transfer effects. Small volumes (less than 2 ml) of catalyst and high heat transfer rates are utilized. This approach is useful for comparing the intrinsic properties of an array of candidate catalysts but does not provide the data required for scale-up. See, for example, U.S. Pat. Nos. 6,149,882 and 6,869,799.
In accordance with the invention, there is provided a low cost, accelerated method for determining a set of operating parameters for a commercial scale plug flow catalytic process and reactor system for hydroprocessing clean feedstocks having high productivity and selectivity to desired products while minimizing to the extent practicable the catalyst deactivation, and especially that caused by the generation of deactivating species as a result of secondary reactions occurring during hydroprocessing.
The method includes supplying a clean hydrocarbon feedstock to the inlet of a composite multistage series-connected laboratory scale plug flow reactor, the stages of which each contain a catalyst suitable for the hydroprocessing of said feedstock. The feedstock is hydrocracked and isomerized in the stages of the laboratory scale reactor at a selected set of operating conditions of temperature, pressure, and reactant and reaction product flow rates, and with the catalysts in the catalyst beds of the stages of said laboratory scale reactor having selected sets of characteristics. The effluents of each of the reactor stages are sampled and the concentration of reactants and catalytic process products and byproducts in the effluents of each of the stages are measured for determining the nature of the catalytic reactions taking place in each such stage. The process is repeated at different selected sets of said operating conditions, and/or at different selected sets of characteristics of the catalysts in the catalyst beds of the stages of said laboratory scale reactor, with the results of the measurements obtained in one hydrocracking and isomerization operation being used to drive the selection of catalyst bed characteristics and operating parameters in subsequent hydrocracking and isomerization operations for improving the productivity and selectivity of the laboratory scale reactor to the desired products, while minimizing to the extent practicable the production of catalyst deactivating species in the stages of said laboratory scale reactor.
The method of the invention further includes supplying to the inlet of said one or more stages of the laboratory scale reactor a model compound not present in the clean feedstock, and measuring the amount of the model compound in the effluents of the following stages of the laboratory scale reactor for determining the rate of disappearance of the model compound. Additionally, in accordance with the invention, selected amounts of a catalyst deactivating specie can be added to the inlets of one or more of the reactor stages of the laboratory scale reactor, and information is determined concerning longitudinal deactivation phenomena occurring in each of the catalyst beds of the stages of the laboratory scale reactor based on the amounts of model compound and amounts and characteristics of reactants and reaction products and byproducts in the effluents of said stages.
The method of the invention can further include supplying a model compound to the inlet of the first stage of the laboratory scale reactor and supplying selected amounts of a deactivating specie to the inlets of the following serial-connected stages of the laboratory scale reactor so that the flow regime in each successive reactor stage contains an increasing amount of the deactivating specie, and determining information concerning longitudinal deactivation phenomena occurring in each of the catalyst beds of the stages of the laboratory scale reactor based on the amounts of model compound and amounts and characteristics of reactants and reaction products and byproducts in the effluents of the reactor stages.
The term “plug flow reactor”, as used herein refers to fixed bed reactors, packed bed reactors, trickle bed reactors and monolithic reactors operating either in a once through or a recycle mode. The term “laboratory scale plug flow reactor” as used herein, refers to a plug flow reactor in which each reactor stage has an internal diameter of less than 4 inches, preferably less than 2 inches, and more preferably less than 1 inch; a length of less than 8 feet, preferably less than 4 feet, more preferably less than 1 foot; and a catalyst charge of less than 800 grams, preferably less than 400 grams, more preferably less than 25 grams (excluding inert diluent particles charged to the reactor).
Referring to
The multistage fixed bed reactor 11 is contained in a temperature control device 33 that, for an exothermic reaction such as in hydroprocessing, could contain a material, such as circulating boiling water or a fluidized sand bath, for extracting heat from the reactor 11 in order to maintain the multistage reactor 11 at a substantially constant temperature. For an endothermic process, such as paraffin dehydrogenation or catalytic reforming, the temperature control device 33 could contain apparatus, such as an electrical heater, to supply heat to the fixed bed reactor 11 in order to maintain the substantially constant desired temperature. Alternatively, for both exothermic and endothermic catalytic processes, the temperature control device 33 can consist of a fluidized sand bath heater in which the multistage reactors are immersed. Other examples of temperature control devices 33 include circulating molten salt, and electric inductance heaters coupled with internal cooling loops.
Each of the catalyst beds 19, 21, and 23 in the reactor stages of the multistage reactor 11 replicates a longitudinal portion of the catalyst bed of a large fixed bed reactor and permits the measurement and analysis of the characteristics and performance of successive longitudinal portions of a large catalyst bed, thereby allowing determination of longitudinal gradients in reactor characteristics and performance that heretofore have been inaccessible. While reactor 11 has been shown as having three series-connected stages, it is equally possible to have a larger number of series-connected stages, e.g., four or six stages, in order to analyze the performance of the composite catalyst bed at a greater number of points along its length.
Depending on the reaction being studied and the data needed, the analysis of the feed and the effluent from the reactor stages can include, e.g., conventional GC/MS or UV or IR characterization of the reactant and product stream(s), and/or analysis of the catalyst system by XRD, diffuse reflectance IR or other spectroscopic techniques that are well known in the art. These studies would allow the performance attributes of the system to be quantified as a function of the longitudinal position in the catalyst bed. Such knowledge would allow the system to be optimized with direct knowledge of the catalytic reaction kinetics and performance attributes of each point and permit the design of catalyst systems in which, e.g., the catalyst particles may have different chemical or physical characteristics in different portions of the catalyst bed so as to operate at peak productivity or selectivity as a function of the local environment.
The catalyst beds in the reactor stages 13, 15 and 17 may be a crushed or powdered catalyst or a commercial-size catalyst. Most measurements made in gathering data for the scale up of a catalytic reactor need to be made with the reactor operating in a substantially isothermal regime. In order for the reactor stages 13, and 17 to operate in a substantially isothermal regime, the catalysts in the beds 19, 21 and 23 are diluted with an inert particulate matter, typically in a ratio of up to about 8-10 to 1. For measurements being made with the reactor operating in a substantially adiabatic regime, the catalyst in the beds 19, 21 and 23 is less diluted, and depends on the heat of reaction of the process under study and reactor diameter. The ratio of catalyst particles to diluent particles in a catalyst bed depends upon a number of factors, including the amount of heat generated by the reaction and the activity of the catalyst particles in the bed. The appropriate ratio for a given reaction, catalyst, reactor diameter and catalyst particle size can easily be determined by one of ordinary skill in the art by a simple experiment.
A commercial-size catalyst in a fixed bed reactor typically has particle size of about 1 to 5 mm. the catalyst particles can be in any one or more a variety of shapes, e.g., round, tubular, trilobe, toroidal, etc. The crushed or powdered catalyst, which is typically formed by crushing a commercial-size catalyst, typically has a particle size of about 0.10-0.20 mm. the crushed or powdered catalyst particles are normally preferably as small as can be obtained while still retaining a performance qualities of the catalyst. The interior diameter of a reactor stage should be about 10 times the diameter of the smaller of the diluent or catalyst particles and the minimum would typically be in the range of about 10 to 50 mm (0.4 to 2 inches) for a bed containing commercial-size catalyst particles and diluent. Crushed or powdered catalyst particles are typically more active than the commercial-size catalyst particles because of lower mass transfer resistance. Therefore, in order for a reactor containing a bed of crushed or powdered catalyst to operate at the same temperature as a similar reactor containing commercial-size catalyst, the ratio of inert diluent particles to catalyst particles in the bed of crushed or powdered catalyst particles normally needs to be higher than that of the bed containing commercial-size catalyst particles in order that the heat release per unit volume of the to catalyst beds is the same.
The interior diameter of a reactor containing crushed catalyst, can, if desired, be smaller, in the range of about 5 to 12 mm, than that of a reactor containing the commercial size catalyst. For reasons of flexibility in the use of the multistage reactor 11 in different applications, however, it may be preferable that the crushed catalyst bed have the same interior diameter as that required for a bed containing commercial-size catalyst particles. Alternatively, the interior diameter of a reactor being used with a bed of crushed or powdered catalyst particles may be reduced by the use of a thermally conductive sleeve within the reactor.
The minimum height of a reactor stage is determined either by mixing or heat release considerations. For isothermal operation, if mixing is the limiting factor, the height should be sufficient to avoid bypassing. Typically, this would be at least about 50 times the average diameter of the particles, or about 50 to 250 mm (2 to 10 inches) for a reactor stage containing a bed of commercial-size catalyst particles. Because the feed is progressively converted as it traverses the stage of the multistage reactor 11, the concentration of fresh feed in the successive reactor stages decreases from one stage to the next. If it is desired to have constant conversion in each reactor stage, the lengths of the catalyst beds 19-23 can be progressively longer in each of the successive reactor stages 13-17. If the reactor 11 is to operate in the adiabatic regime, one would tend to use a lower ratio of inert diluent and a larger diameter reactor.
Referring to
The composite reactor 35 has three series-connected reactor stages 37, 39, and 41 that contain catalyst beds 43, 45 and 47, respectively. A sampling valve 49 is connected between the output of probe reactor stage 37 in the inlet of the probe reactor stage 39 and has an output 50 for sampling the effluent from reactor stage 37. A sampling valve 51 is connected between the output of reactor stage 39 and the input of reactor stage 41 and has an output 52 for sampling the effluent from the reactor stage 39. A sampling valve 53 is connected between the output of reactor stage 41 and, e.g., a product accumulator (not shown), and has an output 54 for sampling the effluent from reactor stage 41. The fresh reactant feed from source 31 is connected to the inlet of the first probe reactor stage 37. A control and sampling valve can be connected between the source 31 in the inlet to the first probe reactor stage 37 for selectively controlling the amount of feed to the probe reactor and to permit the sampling of the feed for analysis. Also connected to the inlet to the first probe reactor stage 37 is a source 55 of a material to be controllably added to the input of the first probe reactor stage 37 for ascertaining the effect of such addition on the characteristics and performance of the stages of the probe reactor 35. A source 57 is connected to the inlet of the second probe reactor stage 39 for selectively adding a material to the input of such a second probe reactor stage for ascertaining the effect of such addition on the characteristics and performance of the second and third probe reactor stages 39 and 41. A source 59 is connected to the input of the third probe reactor stage 41 for selectively adding a material to the input of such probe reactor stage for ascertaining the effect of such addition on the characteristics and performance of the third probe reactor stage 41. In this embodiment of the invention, the catalyst beds 43, 45 and 47 of the probe reactor 35 are preferably the same as the catalyst beds 19, 21 and 23 of the multistage reactor 11, respectively.
The use of the composite multistage probe reactor 35 allows one to measure the transient response of the system to permanent or temporary changes in the feed composition at any stage of the multistage reactor 11 by comparing the characteristics and performance of the relevant stages of the probe reactor 35 over time in response to the change in input with the characteristics and performance of the corresponding stages of the multistage reactor 11. Introduction of a change in gas or liquid input to the third reactor stage of the probe reactor 35 allows one to measure the impact of the changed component on the reaction rate and selectivity of the third reactor stage catalyst bed of the multistage reactor 11 with time. Introduction of the change to the second probe reactor stage allows one to measure the impact on the second and third stage catalyst beds of the multistage reactor 11. This is equivalent to measuring the response to a change in conditions of any small segment of the catalyst bed in a commercial-size fixed bed reactor. For example, raising the gas feed rate to any reactor stage of the probe reactor 35 by having one of the sources 55, 57 or 59 and additional fresh feed into the stage of the probe reactor 35 to which it is connected, would allow the investigation of the changes in incremental performance of that stage and following stages resulting from the change in input over time.
It is also possible to use the sources 55, 57 or 59 to vary the concentrations of the trace components present in the fresh feed in a selected probe reactor stage, for instance by adding fresh reactant feed having a higher or lower concentration of such trace components, in order to quantify the effect of such trace components on various parts of the composite catalyst bed under a full range of operating conditions. By doing this it would be possible to map the critical longitudinal portions of the composite catalyst bed in a commercial system in which the catalyst is most vulnerable to poisoning or other inhibitory reactions caused by poisons or other natural byproducts of the reaction being practiced. The probe reactor 35, and other versions of probe reactor as discussed below with relation to other Figures, can also be used to investigate the transient response of a reactor to temporary changes in the composition of the feed or prior stage effluent to various points in a composite catalyst bed by temporarily adding the materials of interest to a selected stage of the probe reactor 35 and monitoring the time dependent response of that stage and following stages of the probe reactor 35 to such added materials both during and after the time that such materials are added.
Referring to
Referring to
Referring again to
Alternatively, the probe reactor 101, instead of being a fully back-mixed reactor such as discussed above, can be a two-dimensional catalyst array that comprises separate micro reactors. Such a probe reactor can be used to investigate the intrinsic characteristics of a plurality of crushed catalysts in the presence of different mixes of feed, effluent and product.
In the embodiments of
The reactant and other material feeds, and reaction products and byproducts in reactor effluents supplied or generated in the embodiments of the invention described herein may be either gaseous, liquid or mixed phase (such as e.g., gas/liquid or two or more immiscible liquids). Feeds and effluents consisting of gases can be handled using well known conventional back pressure regulators and gas flow control systems with mass flow controllers. Controlled amounts of liquids can be pumped in high-pressure environments using known pumps such as a Ruska pump or a Syringe pump. If the effluent from a reactor stage or the feed contains multiple phases, particularly if such phases are immiscible, such as water and hydrocarbons or liquid and gas, it is important to avoid slug flow. In such case, sampling valves may consist of e.g., iso-kinetic sampling valves such as available from Prosery AS, or splitters such as described in U.S. Pat. No. 4,035,168. Alternatively, the stream may be sampled immediately after a static mixer such as available from Prosery AS, which homogenizes the multiphase stream. In combining immiscible feeds or feeds and effluent to a reactor stage, or in conducting the multiphase effluent from the outlet of one reactor stage to the inlet of the following reactor stage in a series-connected multistage reactor, it is typically the practice to manifold of the streams into a line having a high Reynolds number similar in concept to a fuel injection system in an automobile engine. Alternatively, static mixers such as available from Prosery AS or from Admix, Inc., Manchester, N.H., can also be used. In this case, some simple initial testing may be desirable to confirm that the operating conditions are leading to the homogeneity of the stream passing through the device. If the gas and liquid are well mixed in a transfer line, it is possible, for instance, to take a combined liquid and gas sample in a sample bomb connected to the reactor line via double block valves. The bomb would be at atmospheric pressure or slightly above. The block valves would be opened and liquid and gas would be allowed to flow into the bomb. The two block valves would then be closed, the sample bomb removed from the reactor and the contents analyzed. The presence of a small concentration of an inert gas such as Argon in the stream can be used to allow closure of the material balance. Alternatively, if the phases are not well mixed, one could employ gas/liquid separators and analyze the gas and liquid phases separately with an internal standard such as He or Ar and overall carbon balance analysis to link the two. This could be accomplished e.g., by using a gas sample bomb attached to the top of the line and a liquid sample bomb attached to the bottom of the line.
Hydroprocessing encompasses a large number of reactions involving the reaction of the hydrocarbon stream with hydrogen over a catalyst. A characteristic example of the hydroprocessing of a clean feed is the catalytic hydrocracking of paraffins, which involves the conversion of linear alkanes to mixtures of lower molecular weight molecules, a large fraction of which are branched isomers. During hydrocracking, secondary reactions also occur which generate undesired byproducts such as aromatics, PNAs and other unsaturates that tend to deactivate the catalyst. It is possible to track of the degree of such deactivation by adding to the feed a unique single carbon number feed, such as hexadecane, that is not otherwise present in the clean feed nor generated by conversion reactions of the clean feed being hydrocracked. The measurement of the rate of decrease in the rate of disappearance of the hexadecane (or other such probe molecule) provides an indication of the degree of deactivation of the catalyst. In the case of the present invention, the system would be studied, inter alia, by use of regular feeds which contained a mixture of molecules of differing structure and average molecular weight together with a model compound such hexadecane.
A major area of concern in understanding and controlling the characteristics and performance of a plug flow reactor is the adsorption or reaction of a feed component, product or byproduct with the catalyst surface. For instance, in the hydrocracking processes, materials such as ammonia, carbon monoxide, hydrogen sulfide, can tie up active catalyst sites, reduce reaction rate and adversely impact product selectivity. The reactions caused by these materials can take time to equilibrate and can also take time to be released after removal of the material from the feed stream to the reactor.
Ammonia, or other amines, are known to react with hydrocracking catalysts, causing activity to decline and line out. In the situation where ammonia is the only nitrogen containing component in the feed, once it is removed from the feed, hydrogen can be used to remove the ammonia from the catalyst surface. In investigating the effects of ammonia on different portions of the composite catalyst bed, ammonia can be added to the inlet of any of the stages of probe reactor, thereby replicating the effect of the presence of ammonia in the feed to a selected longitudinal slice of the composite catalyst bed. By controlling the conversion level in a given catalyst slice, e.g., by adjusting temperature and/or flow rate and/or reactant partial pressures in a probe reactor stage, it is possible to define the effect of the ammonia under various operating conditions. By varying the hydrogen concentration in the feed to one or more probe reactor stages, it is possible, for example, to investigate the effect of increased hydrogen on the ammonia-contaminated catalyst in different portions of the composite catalyst bed, e.g., the bed with the greatest activity decline.
Carbon monoxide is tightly held on a hydrocracking catalyst, which can reduce available surface for hydrogen, thereby making hydrogen the rate limiting step. By varying the concentrations of carbon monoxide and hydrogen in the feed to selected stages of the probe reactor 35 or 101 and comparing performance of the relevant probe reactor stages with the corresponding stages of the multistage reactor 11, it is possible to determine the impact of carbon monoxide and hydrogen concentration on reaction rate and selectivity. The use of a multi-stage probe reactor allows for testing of the impact at various conversion levels by e.g., by adjusting temperature and/or flow rate and/or reactant partial pressures.
The amount of Conradson carbon is usually utilized in correlations for hydrotreater performance. In general, carbon deposits on a catalyst inhibit the diffusion of reactants to the catalyst surface and the removal of reaction products from the catalyst surface. This tends to lead to activity reduction via unwanted side reactions with deposits on the catalyst surface or with the diffusion limited reactants or both. In the case of beds containing commercial-size catalyst particles where the diffusion path is the longest, this sort of diffusion limitation can limit overall catalyst life and require costly steps to maintain system performance. Adding different molecular weight fractions of these materials to a selected stage of the probe reactor 35 or 101 would allow the determination of what portion of the composite catalyst bed is impacted the most. The effects of various regeneration techniques such as by the addition of hydrogen, water, or a light solvent can also be determined my controlling the feeds to the relevant stages of the probe reactor 35 or 101, thereby to define the preferred rejuvenation technique. These issues will be particularly important in processing of heavy feeds from tar sands, shale, heavy oil deposits, and coal. These feeds are known to carry many contaminants that can lead to catalyst poisoning, and in situ regeneration, in order to avoid the cost of frequent replacement with fresh unused catalyst, is frequently the only means to make the overall process economically viable.
Polynuclear aromatics are also known to inhibit a catalyst by forming carbonaceous overlayers on catalyst sites that reduce selectivity and activity of hydroprocessing catalysts. The effect of the presence of polynuclear aromatics in the feed at various longitudinal portions of a composite catalyst bed of a fixed bed reactor can be determined by adding the polynuclear aromatics to a selected stage of the probe reactor 35 or 101 and comparing the characteristics and performance of the relevant stages of the probe reactor with the corresponding stages of the multistage reactor 11. This can be used to help define in what portion of the composite catalyst bed the polynuclear aromatics have their greatest impact, and what can be done to improve the process design and catalyst performance, by comparison of the performance and characteristics of the relevant stages of the probe reactor 35 or 101 with the corresponding stages of the multistage reactor 11.
Referring now to
Referring now to
The arrangements of
Having a plurality of composite multistage series-connected reactors disposed in a common temperature environment, such as constituted by the fluidized sand bath 537, or as described above with relation to
Other experiments to be performed that aid in the scaling up of a catalytic process include, for example, investigating the characteristics of a plurality of different catalysts simultaneously. Alternatively, a crushed catalyst in the catalyst beds of one multistage series-connected reactor could be compared with a plurality of different shapes or sizes of commercial-size versions of the catalyst in the catalyst beds of other multistage series-connected reactors, all disposed in a common constant temperature department. In an alternative arrangement, it is also possible to have different catalysts in different reactor stages of the multistage series-connected reactor 11 for testing the catalysts in series. Using such an arrangement, one can design a layered composite catalyst bed in which the intrinsic behavior of each catalyst layer is matched to the local kinetic and mass transfer environment, so that the overall response of the system is varied longitudinally so as to obtain behavior characteristics in each longitudinal portion of the composite reactor that are optimum for process performance. If a plurality of multistage series-connected fixed bed reactors is disposed in separate, independently controllable temperature control devices, a plurality of different heat removal levels can be investigated in parallel.
Referring now to
Each of the reactor stages 151-1 through 151-n contain a catalyst bed 153-1 through 153-n. The modules 155 and 157 can be identical to the module 151, and contain a plurality of parallel fixed bed reactor stages 155-1 through 155-n and 157-1 through 157-n, respectively. Each of the parallel reactor stages in the modules 155 and 157 contain catalyst beds 159-1 through 159-n and 161-1 through 161-n, respectively. In the illustrated embodiment, the outlet of each of the reactor stages in module 151 is connected to the inlet of the corresponding reactor stage in module 155, and the outlet of each of the reactor stages in module 155 is connected to the inlet of the corresponding reactor stage in module 157. Thus, the series connected reactors stages 151-1, 155-1 and 157-1 form a composite multistage series-connected fixed bed reactor. Similarly, the other sets of series connected reactor stages in the modules 151, 155 and 157 also form composite multistage series-connected fixed bed reactors. The modules 151, 155 and 157 may contain any desired number of parallel reactor stages depending upon the application. For instance, each module might contain four or eight or even 16 parallel reactor stages. Is also possible to have additional modules of parallel reactors stages, with each of said parallel reactors stages being connected in series with the corresponding reactor stages of the preceding and succeeding modules. For instance, there might be four or six modules in a given application.
The modules 155 and 157 are surrounded by temperature control devices 158 and 160, respectively, that may be the same as, or common with, the temperature control device 152 that surrounds the module 151. Sampling valves 163-1 through 163-n are connected between the outlet of each reactor stage in the module 151 and the inlet of the corresponding reactor stage in module 155. Sampling valves at 165-1 through 165-n are connected between the outlets of each of the reactor stages in module 155 in the inlet of the corresponding reactor stage in module 157. Fresh reactant feed is fed from a source 167 through control valves 169-1 through 169-n to the inlets of each of the reactor stages 151-1 through 151-n of module 151 for supplying controlled amounts of reactant feed to the inlets of the respective reactor stages. The fixed bed reactor 171 also receives fresh reactant feed gas from the source 167 at its inlet, and has its outlet connected to the inlets of the reactor stages 151-1 through 151-n through control valves 173-1 through 173-n, respectively, for supplying controlled amounts of effluent from the reactor 171 to the reactors 151-1 through 151-n.
In a commercial-size plug flow reactor, the proportion of fresh feed and reaction products and byproducts varies continuously along the length of the catalyst bed. At the inlet there is 100% fresh reactant feed and zero reaction products and byproducts. As the fresh feed is consumed in the catalyst bed of the reactor, the proportion of fresh feed decreases and the proportion of reactant products and byproducts increases longitudinally along the catalyst bed. The multiple parallel-serial reactor arrangement of
The investigation in accordance with the method of the present invention of catalyst deactivation phenomenon in the hydroprocessing of clean feeds can involve a number of different experiments. In one important experiment, a composite multistage series-connected laboratory scale plug flow reactor system similar to that described in connection with
Other experiments can be conducted to probe the kinetics of key reaction steps such as dehyrocycloaromatization, i.e. conversion of paraffins to aromatics, and the desired reverse step of aromatic saturation and ring opening, i.e. the conversion of aromatics to paraffins. In this manner, one can better determine the key structure function relationships of catalyst properties with kinetics of desired reaction steps; and then to define more effective catalysts for those steps. In these experiments on could use model compounds to monitor the rate of dehydrogenation, cyclization and aromatization; and to develop overall kinetic models of the key steps of the integral full scale reactor system.
Kinetics
Heretofore, it has been the practice to measure the kinetics of a plug flow catalytic system only by measurements taken at the inlet and the outlet of the catalyst bed, so that the measurements are averaged over the length of a catalyst bed. In analyzing the kinetic performance of such a reactor, it was necessary to make assumptions concerning the kinetic order of the reaction. Typically, it was assumed that the order of the reaction remained constant along the length of the catalyst bed in the reactor. Applicants have found that this assumption was in many cases incorrect. With the use of the multistage series-connected plug flow reactor of the present invention as described above with relation to any of the
Using the multiple parallel-serial reactor arrangement illustrated in
In order to determine the integral kinetics of the catalytic system formed by a composite multistage series-connected fixed bed reactor as a function of longitudinal positions in the catalyst bed, it is necessary to analyze the inlet feed stream and composition and outlet feed stream and composition, normalized, for instance to STP per standard liter of feed, at each of the successive longitudinal slices of the catalyst bed. For instance in a Fischer-Tropsch reaction, one would measure how many moles of H2 and CO were consumed and how much product and byproduct were produced in each reactor stage. The conversion, or an equivalent quantity, such as the remaining concentration of fresh feed, is then plotted versus the residence time, which corresponds to successive longitudinal positions along the catalyst bed as the reactant feed traverses the catalyst bed. The slope at each point along the resulting curve is equal to the Reaction Rate for the system. The reaction rate is then plotted on a log-log plot versus the concentration of the fresh feed along the reactor catalyst bed. If the resulting curve is a straight line, the integral kinetics of the system is a constant along the length of the catalyst bed. If the line is horizontal, the system has first-order kinetics. If the line has a positive slope, the system has positive order kinetics greater than one. If the line has a negative slope, the system has negative order kinetics.
If the resulting curve on the log-log plot is not a straight line, then the integral kinetics of the system varies along the length of the reactor catalyst bed. In this case, it is necessary to do a regression analysis to fit the curve to an equation relating the reaction rate to the concentration of feed. Differentiating that equation, either graphically or mathematically, gives the Rate Model Correlation as a function of longitudinal position along the catalyst bed. A representative graphic technique is discussed in Graphical Methods for Data Analysis, John M. Chambers, Chapman and Hall, May 1983, ISBN: 0412052717.
In order to determine the effects of temperature and pressure on the integral kinetics of the system, the above-described experiment can be run at different temperatures and at different pressures. The experiment can also be run using different size catalysts. For example, the experiment can be run using the intended commercial size and shape catalyst and also with a diluted crushed or powdered catalyst.
The intrinsic and differential kinetics, free of mass transfer and heat transfer effects, of the composite multistage series-connected fixed bed catalytic system of the invention can also be investigated for purposes of scale-up to a commercial system using the systems depicted in
While these kinetics measurements have been described with relation to
Mass Transfer
Methods of investigating the mass transfer characteristics of a catalytic process in a plug flow reactor, such as a fixed bed reactor, typically involve a comparing the conversion versus residence time characteristics at a given set of operating conditions of a finely crushed with that of a commercial-size catalyst. The crushed catalyst is screened to a narrow particle size range, preferably one that is close to the minimum obtainable catalyst particle size that still retains its catalytic properties. This minimum catalyst particle size depends on the characteristics of the specific catalyst being used, and can be determined by simple experimentation. In the more simple method for determining the mass transfer characteristics, the finely crushed and screened catalyst is assumed not to have any mass transfer limitations, so that any difference in the conversion versus residence time characteristics between the crushed catalyst and the commercial-size catalyst is assumed to be the result of mass transfer limitations. For a given feed, the effluent of the two reactors is sampled to determine the amount of conversion. Alternatively, the input flow rates of the two reactors can be adjusted (i.e., the input flow rate to the crushed catalyst in reactor is increased, or the input flow rate to the commercial-size catalyst reactor is decreased) so that each of the reactors has the same percentage conversion, and that difference in residence times is attributed to mass transfer limitations in the commercial-size catalyst.
In a more rigorous and technically exact method of determining the mass transfer characteristics of a commercial-size catalyst, the finely crushed catalyst is not assumed to have zero mass transfer limitations, and the Thiele Modulus of the commercial catalyst is determined from the ratio of the observed reaction rates of the crushed and commercial-size catalysts and the ratio of their particle sizes. The Effectiveness Factor for the commercial-size catalyst can then be determined from a plot of the effectiveness factor versus the Thiele Modulus. This method is described in Hougen and Watson, Chemical Process Principles, Part III, Kinetics and Catalysts, pp. 998-1000, Wiley, March 1966, which is incorporated herein by reference.
A problem with both of these methods is that they does not give any information concerning longitudinal variations in mass transfer performance along the reactor catalyst bed and basically assumes that the mass transfer characteristics are uniform from input to output. This assumption is incorrect for many catalytic systems, and the inability to investigate the longitudinal variations in mass transfer characteristics in a fixed catalyst bed has meant that information which would allow the optimization of the catalyst bed along its length has not been available.
In accordance with the present invention, the catalyst beds of the fixed bed reactors are segmented longitudinally into at least three series-connected stages and the effluent of each of the stages is sampled to determine the amount of conversion occurring in each longitudinal segment of the catalyst bed. Referring again to
The catalyst beds 43, 45 and 47 in reactor stages 37, 39 and 41 of reactor 35 are composed of commercial-size catalyst particles that also may be mixed with a lesser percentage of inert diluent particles so that the operation of reactor 35 is also essentially isothermal. To investigate the longitudinally-dependent mass transfer characteristics of the commercial-size catalyst in accordance with the method of the invention, each of the reactors 11 and 35 receive the identical reactant feed from the source 31 and the pressure and the feed rate for each of the two reactors is held constant. The conversion versus residence time relationship is obtained for each stage of the reactors 11 and 35 from the difference in the amount of reactant feed at the inlet and outlet of each reactor stage and the flow rate, for a given set of operating conditions.
In the simplified method of determining mass transfer limitations, the Effectiveness Factor for the commercial-size catalyst is obtained for the commercial-size catalyst at each stage of the reactor 35 by taking the ratio of the Observed Reaction Rates of the commercial-size catalyst and the crushed catalyst for each reactor stage. The Observed Reaction Rate is obtained for each reactor 11 and 35 by plotting the cumulative conversion of reactant and corresponding cumulative appearance of the product and byproducts (if any) versus residence times at the outputs of the reactor stages of each reactor and fitting curves to the data using well-known techniques. See, e.g., Graphical Methods for Data Analysis, John M. Chambers, Chapman and Hall, May 1983, ISBN: 0412052717. See also related studies such as, A Mechanistic Study of Fischer-Tropsch synthesis using transient isotopic tracing. Part-1: Model identification and discrimination, van Dijk et al., Sections 3, 5 and 5.2. &
If the Hougen and Watson method is used, the Ko of the crushed catalyst is not assumed to be equal to the K. According to this method, it is possible, using the graph of
Because the Effectiveness Factor is the ratio of Ko to the Ki, it is possible to calculate the Ki for a catalyst from the Effectiveness Factor and the Ko for a given longitudinal point along the catalyst bed. Since Ki is the same for the crushed and commercial-size catalysts, the Effectiveness Factor for the commercial-scale catalyst at any point along the catalyst bed can be determined from the Ko for the crushed catalyst at that point and the K.
For reactions in which different reaction pathways are possible in different longitudinal portions of the catalyst bed of the fixed bed reactor, e.g., conversion of sulfur or nitrogen containing feedstocks, carburization, or the production of methane via hydrogenolysis, it is important also to characterize the behavior of the different kinetic pathways producing the product and various byproducts that can exist for the system as they vary along the length of the composite catalyst bed of the reactor in order to explore the longitudinally dependent kinetic and mass transfer space for the system, and to distinguish between the occurrence of mass transfer and kinetic effects in the system. When this space has been explored, the mass transfer performance of reactant to product for the system operating at a given set of conditions that involve an optimal set of trade-offs for the particular catalyst can be investigated.
An example of the opportunity to optimize the longitudinal characteristics of a catalyst bed of a fixed bed reactor afforded as a result of the data obtained by the method of the present invention is illustrated in connection with the graph in FIG. 10 of the drawings. This graph depicts what is believed to be a typical relationship between the Effectiveness Factors and conversion rates for crushed and commercial-size catalysts in a fixed bed reactor working at a given set of operating conditions of temperature and pressure and with a common reactant feed. Mass transfer limitations are clearly present up to the point in each of the fixed bed reactor catalyst beds at which about 50 to 60% conversion has occurred, but are not present at the portions of the catalyst beds at which greater than about 70% conversion has occurred. The greater mass transfer limitations, evidenced by of the lower Effectiveness Factor, of the bed containing the commercial-size catalyst particles is believed to reflect the differences in the lengths of the reaction pathways in the crushed and commercial-size catalyst particles. This suggests that buildup of material, such as wax, in the catalyst pores is present at the portion of the reactor catalyst bed at which lower conversion has occurred, i.e., close to the inlet of the reactor bed where the catalyst experiences almost entirely fresh feed, but not present at lower portions of the catalyst bed at which higher conversions have occurred.
In a hydroprocessing reactor, a lower Effectiveness Factor results in an undesirable higher methane and/or carbon over-layer formation. Thus, particularly in the reactor containing the commercial-size catalyst bed, the upper portions of the catalyst bed would be producing substantial amounts of methane. This results in a much lower diffusivity of the reactant gases in such pores, so that the active sites within the catalyst become starved of reactants and begin generating large amounts of methane. In order to optimize the catalyst bed structure of the fixed bed reactor to avoid the undesirable high methane make in the inlet portions of the catalyst bed, it is possible, for instance, to use a less active catalyst in that portion of the reactor bed, which would generate lesser amounts of methane.
As an alternative to using crushed in commercial size catalyst particles of different sizes in an investigating the mass transfer characteristics of the catalyst bed in a fixed bed reactor, is possible to use the same size particles with different levels of catalyst loading. The particles would be made up of finely crushed or powdered catalyst dispersed and inert diluent such as alumina or silica. The powder or finely crushed catalyst is uniformly mixed with the finely crushed inert diluent, formed into particles of a given size and sintered. Particles in which the catalyst concentration is selected to be relatively low can correspond to the crushed catalyst in the method described above. Particles in which the catalyst concentration is relatively higher can correspond to the commercial-size catalyst. The concentration of catalyst within the particles appropriate for the particles to correspond to crushed catalyst or commercial-size catalyst depends on the activity of the catalyst and the nature of the reaction.
In scaling-up a reactor to commercial size, is preferable to confirm the mass transfer characteristics determined under isothermal conditions in the manner described above in an adiabatic reactor. In an adiabatic reactor, the amount of diluent for the commercial-size catalyst is reduced and the tube diameter is controlled so that its thermal performance mirrors that expected for the commercial-size reactor.
In investigating mass transfer effects in a hydroprocessing reactor, as an alternative to plotting the reaction rate versus conversion or residence times, is to plot the methane selectivity versus conversion. Methane selectivity is greater when mass transfer limitations exist. Mass transfer is an issue only in those parts of the reactor where methane selectivity is widely different for the commercial and crushed catalyst. Between about 35% and 80% conversion, the methane selectivity is very low. In this region, mass transfer is not an issue. In the portion of the catalyst bed where above about 80% conversion has occurred, the methane selectivity increases rapidly and the reaction rate slows down for both the crushed and the commercial catalyst. This is an indication that something other than mass transfer effects is limiting the catalyst activity and increasing the methane selectivity.
Heat Transfer Effects
Understanding the heat transfer performance of a plug flow reactor is critical to maximizing the productivity at which the reactor can be run. For exothermic reactions, such as Fischer-Tropsch, the reaction rate is higher at higher temperatures. However, if the temperature is allowed to become too high, there is a danger of temperature runaways. The temperatures in the catalyst bed of a fixed bed reactor can vary both longitudinally and laterally within the catalyst bed. For exothermic reactions in a fixed bed reactor, excess heat must be removed through the walls of the reactor to a medium such as circulating boiling water or a fluidized sand bath. For endothermic reactions, the problem is the opposite; there can be the need to get heat into cold spots in the catalyst bed or the reaction may shut down.
The reactor system illustrated in
Temperature and product measurements are preferably a repeated for different reactor flow rates, pressures, and productivities, both at Start of Run and during the reactor's time on stream as the reactor lines out. The effect on heat transfer characteristics and other process parameters, such as conversion, selectivity and kinetics, of using catalyst particles of various sizes and shapes in the catalyst bed can also be investigated using the method of the invention. The data obtained from such measurements permits one to investigate and gain an understanding of how the heat transfer properties of the reactor system affect reactor performance over the entire multivariable space in which the commercial-size reactor might operate.
Referring now to
If the reactor 201 is operated at a given conversion level, e.g. 80%, the input to the individual reactor stages 207-1 through 207-n can represent any degree of conversion from zero to 80% by using the control valves 209-1 through 209-n and 213-1 through 213-n to adjust the ratio of reactor 201 effluent to fresh feed being supplied to the individual reactor stages 207-1 through 207-n. Thus, if the valves 209-1 and 213-1 are adjusted such that reactor stage 207-1 receives only effluent from the reactor 201, and the thickness of the catalyst bed 211-1 is such that it performs an additional 5% conversion on such effluent, the catalyst bed 211-1 is equivalent to a cross-sectional slice of a fixed bed reactor in which the conversion between 80 and 85% takes place. Similarly, if the valves 209-2 and 213-2 are adjusted such that the input to reactor stage 207-2 is equivalent to the effluent of a reactor operating at 40% conversion, and the thickness of the catalyst bed 211-2 is such that it performs an additional 5% conversion on such effluent, the catalyst bed to an 11-2 is equivalent to a cross-sectional slice of a catalyst bed in which the conversion between 40 and 45% takes place. Thus, the catalyst beds 211-1 through 211-n can replicate the performance of a cross-sectional slice of a fixed bed reactor positioned at any longitudinal position along the catalyst bed.
The catalyst beds 211-1 through 211-n need not all have the same composition. For instance, the beds 211-1 and 211-2 could contain crushed and commercial-size catalyst particles, respectively, in each case mixed with an amount of inert diluent particles such that the beds operate in isothermal mode. In this case the mass transfer, heat transfer and kinetics characteristics of a cross-sectional slice of a catalyst bed located at any longitudinal position in the catalyst bed can be investigated. In a different application, the catalyst beds 211-1 through 211-n could contain catalyst particles of different chemical or physical composition. In order to prevent heat loss or gain in the effluent from the reactor 201 being fed to the reactor stages 207-1 through 207-n, the connecting tubing and valves are preferably surrounded by insulating material and the entire system comprising the reactor 201 and the reactor stages 207-1 through 207-n can be surrounded by a temperature control device, or alternatively, the reactor 201 and reactor stages 207-1 through 207-n can be surrounded by separate temperature control devices, depending on the needs of the application. Additionally, the reactant feed from the source 205 being supplied to the reactor stages 207-1 through 207-n can be heated before it is supplied to such reactor stages by well-known indirect heating means such as a coil in a sand bath or an infrared furnace (not shown) in order to have the appropriate temperature conditions in the catalyst bed inlet portions of such reactor stages.
The apparatus disclosed in
Using present invention, the time for scale-up of the catalytic process from discovery to commercial scale application can be significantly reduced. For example, in one particularly advantageous configuration, four multi-stage reactors of the type described above can be operated in parallel. In this embodiment, the stages of one of the reactors are loaded with crushed catalyst. This reactor provides Intrinsic Reaction Rate and selectivity data. The stages of the second reactor are loaded with commercial-size catalyst. The data from this second reactor can be used to define the degree of mass transfer limitation (Effectiveness Factor) based on a direct comparison of the relative residence times in the reactors containing the crushed catalyst in the commercial-size catalyst required to achieve a given amount of conversion. By obtaining conversion data at a series of residence times, it is possible to determine the Effectiveness Factor and hence the Effective Diffusivity with conversion or residence time. This data also provides information on the impact of mass transfer on selectivity. A third, probe reactor can be operated in parallel with the previous two reactors. This probe reactor can either be a shallow fixed bed reactor or a back-mixed reactor. Flow can be directed to the appropriate actor from any of the reactor beds in the previous two reactors. In addition, additional gases or liquids can be added to the probe reactor to determine the rates of adsorption or surface property changes on the catalyst. This information can provide valuable insight in modeling the fixed bed reactor. Finally, an adiabatic reactor can be operated in parallel to test the reactor model developed from the previous reactors. Operation of the series reactors in this parallel mode allows for much faster generation of the required scale-up data. In fact, all the required scale-up data, including deactivation and regeneration data, at one temperature can be obtained in one to two years, for a savings of several years of development time. A further improvement to the experimental design would be to operate several four reactor sets at the same time. These sets can be operated at different temperature, pressure, and feed compositions. The set producing the optimum economics can be used for the commercial design. The cost of operating several parallel sets of series reactors simultaneously is a small expense when compared to the potential savings associated with accelerating the scale-up of a new catalyst to a full-scale commercial operation. If the new catalyst results in a $1/barrel savings, a 100 thousand barrel/day plant will produce a savings of over $30 million per year. These savings would easily far more than offset the cost of operating the parallel sets of series reactors.
In an adiabatic reactor, it is possible to produce hot spots in the reactor, which may cause the adiabatic reactor to run away. Also, in an adiabatic reactor, because reaction parameters, such as temperature, kinetics parameters, etc., can change continuously, it is difficult to measure the reaction parameters by direct measurement. Dividing an adiabatic reactor into multistage series-connected reactor stages can help determine reaction parameters at different locations along a flow direction of the reactor, but it is difficult to keep continuities of the reaction parameters, especially temperature, between adjacent reactor stages.
Therefore, it is difficult to directly measure reaction parameters in an adiabatic reactor, and to exactly and securely determine reaction characteristics in the adiabatic reactor, such as kinetics, mass transfer, heat transfer etc.
In one embodiment, the reactor stages 61, 63 and 65 are isothermal reactor stages, which are used together to simulate an adiabatic reactor. Thus, temperature control devices 604, 605 and 606 are provided to control the temperature of the reactor stages 61, 63 and 65 respectively. A preheater (not shown) may be disposed between the source 60 and the first reactor stage 61 to preheat the feeds from the source 60 so that when the feeds flow into the first reactor stage 61, the feeds have already reached a desired temperature for the feeds. Alternatively, the preheater can also be disposed in the first reactor stage 61.
In one embodiment, when using the isothermal reactor stages 61, 63 and 65 to simulate the characteristics of an adiabatic reactor, the temperature setting for each of the temperature control devices 604, 605 and 606 should be determined first. Generally, for a given catalytic process, based on data derived from operating the adiabatic reactor in practice, the temperature setting for the first temperature control device 604 and temperature variation in the first reactor stage 61 can be determined. Then, based on information from the first reactor stage 61, the temperature setting of the second temperature control device 605 can also be determined, and so on. Thus, after the temperature settings of each of the temperature control devices 604, 605 and 606 is determined, the reactor stages 61, 63 and 65 can be used to simulate the characteristics of the adiabatic reactor.
In this embodiment, the temperature of the temperature control devices 604, 605 and 606 are defined as T1, T2 and T3, which are different from each other. Different catalytic processes may have different T1, T2 and T3 settings. Alternatively, a common temperature control device (not shown) can be provided to control the temperatures of reactor stages 61, 63 and 65 together.
Thus, the isothermal reactor stages 61, 63 and 65 can respectively simulate successive catalyst bed slices of a catalyst bed of a larger adiabatic reactor. Thus, the characteristics of the catalyst bed, which is simulated by the catalyst beds 62, 64 and 66, are determined. Because it is relatively easy to operate the isothermal reactor stages, characteristics associated with the larger adiabatic reactor can be determined by simulating the adiabatic reactor using the isothermal reactor stages. In this embodiment, the first, second and third reactor stages 61, 63 and 65 can be arranged upright.
For a particular catalytic process between at least two successive reactors, for example a particular catalytic process in a multistage series-connected reactor stages, if an effluent fluid from one reactor stage is homogeneous, such as in a gas phase, transferring effluent fluid can be quite straightforward by using a properly sized and shaped tube connecting an outlet of one reactor stage to an inlet of a following reactor stage. In many catalytic processes, however, the effluent from a reactor stage may be in a multiphase state, meaning that it includes one or more gaseous fluids, which are fluids in gas phase (such as gases, vapors or mixtures of gases and vapors), and one or more liquid fluids, which are fluids in one or more liquid phases (such as water phase, oil phase, other immiscible phases and partial emulsion phases, etc.)
The multiphase fluid is often a multi-component fluid, each component being in its own state, which can be a single-phase state or multiphase state. If the multi-component fluid is in thermodynamic equilibrium, the fluid can be transferred directly by a tube connecting two successive reactor stages.
However, in certain catalytic processes, such as hydrodesulphurization etc., the multi-component fluid may not be in thermodynamic equilibrium. So, when the multi-component fluid is transferred directly through the tube connecting the outlet of one reactor stage to the inlet of the following reactor stage, the states of the components may vary during the transfer such that continuity or consistency of the fluid between adjacent two reactor stages may be broken. Thus, it is difficult to use the multistage series-connected reactor stages to model a plug reactor and to measure and optimize the corresponding catalytic processes.
Referring to
In one embodiment, ΔP is large enough so that it can drive the liquid fluid in the separator 703 to enter into the liquid conduit 76 and to flow into the second reactor stage 73 but is also small enough so that it can not affect reactions in the second reactor stage 73. The flow restrictor 705 can be a restricting valve, an orifice, or other restricting means etc. When properly sized and shaped, the gas conduit 75 can function as the flow restrictor 705. The flow resistance of the gaseous fluid can be adjusted by many ways, such as electrical, electromagnetic, pneumatic, mechanical or thermal ways etc., which are familiar to those ordinary skills in the art. The electromagnetic ways are preferred.
Additionally, the catalytic process development apparatus further includes a differential pressure sensor (not shown) disposed across the flow restrictor 705 or two ends of the gas conduit 75 to measure the ΔP. Combined ΔP and physical properties of the gaseous fluid, information about a mass flow rate of the gaseous fluid can be determined.
In one embodiment, if ΔP is too small, the liquid fluid can not flow but accumulate in the separator 703. If ΔP is too large, the liquid fluid may keep flowing until all the liquid fluid in the separator 703 is transported to the second reactor stage 73. When the liquid fluid in the separator 703 is drawn out, the gaseous fluid may flow through the liquid conduit 76. Thus, ΔP is reduced due to an extra pathway for the gaseous fluid. Then, the liquid fluid begins to accumulate in the separator 703 and blocks the liquid conduit 76. Subsequently, the ΔP restores to a desired value little by little, and the liquid fluid starts to flow again. Thus, the flow rates of the gaseous and liquid fluids may fluctuate with respect to time because of fluctuation of the ΔP, which is disadvantageous to the second reactor stage.
In a preferred embodiment, the catalytic process development apparatus includes a liquid level sensor 706 disposed in the separator 703. The liquid lever sensor 706 monitors variation of a liquid level 704 in the separator 703. Liquid sensor signals from the liquid level sensor 706 are used to control the flow restrictor 705 to generate a suitable ΔP to drive the liquid fluid in such a manner that the liquid level 704 is maintained at a desired substantially constant level. Thus, the fluctuation of the fluids in the separator 703 can be eliminated. When the liquid fluid is transferred stably through the liquid conduit 76, the liquid mass flow rate information can also be obtained by using the measured ΔP in combination with physical properties of the liquid fluid.
In one embodiment, in certain low pressure reactions including low pressure FT synthesis etc., a small pressure drop ΔP may still be too big to tolerate, especially when the reactor stage is long or there are many reactor stages. Additionally, in the process of adjusting ΔP to maintain the liquid level 704 by the liquid level sensor 706 and the flow restrictor 705, the fluctuation of ΔP may also affect liquid flow in the first reactor stage 71.
In this embodiment, the liquid pump 707 includes a positive displacement pump or a centrifuge pump etc. Additionally, the liquid pump 707 can have metering capability, which can be used to obtain the liquid flow rate information directly. In order to cause the liquid fluid to be distributed uniformly in the second reactor stage 73, a sprayer or similar spraying devices (not shown) can be adopted inside the reactor stage 73. Alternatively, a check valve (not shown) may be disposed on the liquid conduit 76 and located behind the liquid pump 707 to prevent the liquid fluid in the liquid conduit 76 from reflux.
In the embodiments of the present invention, the gaseous fluid and the liquid fluid in the effluent of the first reactor stage 71 are separated in the separator 703, and then transported to the second reactor stage 73. Thus, possible interactions between the gaseous fluid and the liquid fluid in the effluent during transport can be minimized, and the potential of altering the states of the components in the effluent by fluid distribution and recombination processes can be reduced. The continuity or consistency of the components of the fluid can be maintained between the first and second reactor stages 71 and 73. Additionally, separation of the gaseous fluid and the liquid fluid also makes it easy for sampling the fluids for species analysis, whether continuously or intermittently, on-line or off-line.
As mentioned above, in certain catalytic processes, there are different types of liquid phases for the multiphase effluent fluid. In one example of the FT synthesis, its effluent may contain water phase liquid(s) and oil phase liquid(s). In order to transport such multiphase fluid uniformly, an agitation device (not shown) can be provided to cause homogenization of the multiphase fluid. The agitation device may include a mechanical stirring device, a magnetic stirring device or an ultrasonic stirring device etc. In one embodiment, the ultrasonic stirring device is provided, which can be installed near a bottom of the separator 703. The ultrasonic stirring device can provide sufficient homogenization of the liquid fluid, while having minimum interference to the performance of the liquid level sensor 706 and also without significantly increasing liquid temperature.
Referring to
Referring to
The composite multistage reactor 707 can include three or more series-connected reactor stages. The outlet of each of the reactor stages can connect to a separator. The separator and the reactor stage can be separate from or integrated with each other. All the reactor stages can also be arranged upright along a vertical line.
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/CN2007/071392 | 12/29/2007 | WO | 00 | 6/29/2009 |
| Number | Date | Country | |
|---|---|---|---|
| 60882868 | Dec 2006 | US |
