MXPA06008045A - Improved olefin plant recovery system employing a combination of catalytic distillation and fixed bed catalytic steps - Google Patents

Abstract

Presented is an improvement to a previous invention involving the catalytic hydrogenation of the C2 to C5 and heavier acetylenes and dienes in a thermally cracked feed stream without significantly hydrogenating the C2 and C3 olefins. The improvement involves the use of a fixed bed hydrogenation reactor system in combination with a modified version of the catalytic distillation unit used in the prior art. The modification to the catalytic distillation unit involves improvement of the liquid recycle scheme. The fixed bed reactors combined with the modified catalytic distillation allows for 100%conversion of acetylene and helps to maintain high conversion of the other dienes and acetylenes with no ethylene or propylene conversion under a variety of conditions. These condition variations include but are not limited to the feed diene and acetylene composition, the mol%carbon monoxide in the feed, and catalyst deactivation. With catalytic distillation alone, complete conversion of the acetylene as stated above can not be achieved without ethylene loss, nor would satisfactory operation and control be possible under the variety of conditions experienced during a commercial operation.

Description

IMPROVED OLEFIN PLANT RECOVERY SYSTEM THAT USES A COMBINATION OF CATALYTIC STATIONS OF FIXED BED AND CATALYTIC DISTILLATION Background of the Invention The present invention relates to a method for the production of olefins and particularly to process the effluent of catalytic disintegration heater in order to more effectively recover the product and process the by-products. In the production of ethylene and propylene through pyrolysis of a variety of feed materials, various by-products and unsaturated diolefins and acetylene are created. The net effluent from pyrolysis heaters, typically referred to as charge gas, requires processing for the separation of by-products and removal of diolefins and acetylenes from the primary olefin products. The removal of the heavier and acetylenic C2 and diolefins from the catalytic disintegration gas is handled through a combination of separation by distillation and reaction by hydrogenation. Specifically for acetylene, the separation alone will result in excessive loss of the ethylene product since acetylene and ethylene will have a similar relative volatility. Currently, the distillation and hydrogenation are carried out in several distinct process steps which are designed to separate and hydrogenate the compounds C2, C3 and C4 independently. The separation of the different hydrocarbons before hydrogenation is currently required to achieve better control over the hydrogenation, prolong the useful life of the catalyst and improve the performance.

A disadvantage of this widely practiced conventional technology is the high energy consumption necessary to generate the high pressures and cryogenic temperatures required to first separate the hydrogen from the catalytic disintegration gas and subsequently subsequently the higher carbon number molecules. Additionally, the hydrogenation stages for each of the hydrocarbon groups require an independent reactor system consisting of several pieces of equipment that increase the investment in capital and complexity of the plant. The invention set forth in the US patent. previous number 5,679,241 proposes the conversion in one stage of all C2 to C5 and heavier acetylenes and dienes, without integration of the C2 or C3 olefins. It is claimed that this is possible with a catalytic distillation unit capable of treating the relatively low pressure, hot charge gas, before excessive compression and cryogenic cooling is performed. Also, if desired, this same one-stage process claims that it is capable of hydrogenating the C4 olefin in paraffins, again without the loss of C2 or C3 olefins. The patent refers to a system that is described as capable of removing 70% and more of the hydrogen in the catalytic disintegration gas before the cryogenic separation required by the hydrogenation of the C2 to C4 acetylenes and dienes and the C4 and heavier olefins to paraffin The removal of 70% or more of the hydrogen improves the economy, through a significant reduction of the energy requirements for separation of the C2 and heavier components. By reducing the partial pressure of hydrogen, separation is achieved at lower pressures and with reduced cooling. However, it has been shown that this extensive hydrogenation in a single-stage system can not occur without substantial loss of ethylene and propylene in paraffins by hydrogenation. The process as described in USP 5,679,241 has significant limitations. First, in the operation of an ethylene plant, the C2 acetylene must specifically be removed by hydrogenation since its removal by distillation is extremely difficult requiring extensive equipment and energy costs. Since acetylene is a poison of polymerization catalysts, it must be removed at low levels, often at less than 1-2 ppm. The ability to hydrogenate all C2 acetylene at that level in a single catalytic distillation column, while no loss of ethylene is observed, or preferably a gain at reasonable catalyst volumes and commercially viable operating conditions was not possible. Second, maintaining the performance of the process during either a variation in carbon monoxide flow, which impacts catalyst activity and / or diene / acetylene concentration in the feed was difficult to manage and proves to be difficult to achieve. commercial scale. Third, methods to handle eventual catalyst deactivation were limited. Since these units must operate for long periods of time between interruptions, the only options where excess catalyst or separated catalyst zones in the reaction column can be isolated and the catalyst replaced while the other sections remain in operation. When large volumes of catalyst are used, it is known that it is necessary to operate at lower temperatures to avoid overreaction while the catalyst is still active. This negatively impacts the economy by requiring some refrigeration to control an operation at lower temperature and / or excessive recycling of the cooling liquid within the column. Specifically, pilot tests have shown that: a. When the simple catalytic distillation column is operated to remove more than 95% of the C2 acetylene, the concurrent loss of ethylene was about 1% weight. This is economically undesirable. b. When a single catalytic distillation column is operated, and the hydrogenation of C4 olefins is greater than 20%, significant ethylene loss will be expected with the catalysts currently available. c. In order to achieve minimal losses of ethylene and propylene while operating a single catalytic distillation column and maintaining extremely high conversions and hydrogen removals, the design requires excess catalyst as evidenced by low productivity and operation at colder temperatures than they require refrigeration. d. A significant variation in catalytic activity will occur with variations in carbon monoxide in the feed. These variations, if seen in a single-stage process, will result in loss of acetylene removal efficiency and subsequent products that do not meet the specification. This impact on performance due to the loss of catalytic activity by CO poisoning is equivalent to the impact on performance due to aging of the catalyst. and. A significant variation in the feed signal to the catalytic pyrolysis heaters of ethylene will result in substantial changes in both acetylenes and dienes as well as in the flow of hydrogen. As the ratio of hydrogen to reagents changes, a single-stage process has limited capacity to follow these changes. The results will already be an acetylene breakthrough that leads to the ethylene product outside the specifications or a high loss of valuable ethylene and propylene due to overreaction, unless the system has an excessively complex, substantial and expensive design. that could be used to overcome these process changes. SUMMARY OF THE INVENTION The present invention relates to an improved process for the processing of effluent cargo gas from the pyrolysis of a variety of feedstocks. The primary objective is still to remove a significant fraction of the hydrogen in the effluent by hydrogenation of the C2 to C5 diolefins and acetylenes in the feed while essentially total hydrogenation of the C2 acetylene is achieved without significant hydrogenation of the ethylene and propylene. In the improved process, this is achieved even with disturbances in the concentration of carbon monoxide, varying diethylene and acetylenic feed concentrations and deactivation of catalyst as well as other foreseeable processing disorders. The invention relates to catalytic distillation with improved liquid recycling in combination with fixed bed hydrogenation reactor systems. Specifically, the operating conditions of the catalytic distillation are maintained or adjusted to obtain the maximum hydrogenation of the acetylenes and dienes but without any loss of ethylene and propylene and preferably with a gain of ethylene by hydrogenation of C2 acetylene. Maintaining a high stable conversion of all C2 to C5 acetylenes and dienes with 100% conversion of C2 acetylene (even without hydrogenating ethylene and propylene) under varying process conditions, is made possible by the fixed bed hydrogenation system in where the remaining acetylene C2 is completely hydrogenated again without significant hydrogenation of ethylene or propylene. Brief Description of the Drawings Figure 1 is a flow diagram of the prior art involving catalytic distillation only with the recirculation of bottoms for temperature control. Figure 2 is a flow diagram illustrating the present invention. Figure 3 is a graph illustrating the gain or loss of ethylene against the level of diene output for the present invention compared to the prior art. Figure 4 is a flow chart similar to Figure 2 but illustrating another embodiment of the present invention. Figure 5 is a flow chart of an alternative embodiment of the present invention. Figure 6 is a flow diagram illustrating an alternative embodiment of the process of Figure 5. Figure 7 is a flow diagram similar to Figure 2 but illustrating an alternative embodiment of the present invention. DESCRIPTION OF THE PREFERRED EMBODIMENTS For a better understanding of the present invention, the prior art as represented by the process described in U.S. Pat. No. 5,679,241 will be briefly described. Figure 1 of the drawings of the present invention is essentially a copy of a drawing of that prior patent, simplified to identify only those features relevant to the present invention. The charge gas 150 is compressed and fed to the catalytic distillation column 156. This column, as in the present invention, simultaneously carries out a catalytic reaction and distillation. The column has an extraction section 158 below the feed and a rectification / reaction section 160 on the feed containing the catalyst beds 166, 168 and 170. Descending fluid is withdrawn as side streams through the intercapacitors 180 and the column is injected back onto the next lower catalyst bed. A portion of the heat of reaction is removed by these intercapacitors. A liquid recycle stream 260 from the extraction section is recycled to the evaporation product section of the column. This recycle 260 may be a portion 262 of the bottoms and / or a portion 264 of the interior of the extraction section. The evaporation products section of column 156 passes to condensers 186 and 188 and the partially condensed stream enters the separation vessel 190. The vapor product C2 to C5 of evaporation 194, contains ethylene and propylene, then passes to the subsequent separation while the condensed hydrocarbons are used as reflux 196 for the column. Since the object of the invention is to completely remove acetylene impurities from ethylene without loss of the ethylene entering the column, this must be achieved in this single stage operation (a catalytic distillation column). The vapor stream of evaporation products passes to additional fractionation (not shown) where the individual carbon number fractions are isolated. The hydrogenation in the catalytic distillation column 156 occurs in the liquid phase. The column is operated in such a way that the composition of the liquid phase are primarily C5 components. This reduces the concentration of liquid phase of ethylene and propylene and in this way reduces its reaction. However, the concentration of these two valuable olefins in the liquid will not be zero. C2 acetylenes and C3 acetylenes and olefins are more reactive than their olefin counterpart. In the liquid phase, preferably they are reacted with ethylene and propylene, respectively. However, as the reaction reaches completion near the desired conversion of 100% C2 acetylene, there is no more significant concentration of higher C2 acetylene reagent present in the liquid phase. Under these conditions, acetylene can react to form ethane (paraffin). This is the point where the loss of ethylene occurs. There are several options to reduce the reactivity in the catalyst beds of this upper portion of the catalytic distillation column. One option is to reduce the temperature. Reducing the temperature in a section is difficult in a single column concept because the distillation temperature is primarily controlled by column pressure. To do this, a second column that operates at a lower pressure will be required. A second option is to use a different catalyst in this upper section. This catalyst will be designed to more selectively hydrogenate acetylenes and diolefins. Both of these options, however, will require higher volumes of catalyst within the column and thus increase the size and cost of the column. A third option is to use an intercooler as illustrated in the prior art. Now with reference to Figure 2 illustrating one embodiment of the present invention, the charge gas 10 is compressed into 12 between 10.34 and 17.24 bars gauge (150 and 250 psig) and then fed to the catalytic distillation column 14. The gas load may or may not preheat to equalize column temperatures. The charge gas will typically pass through one or more protection beds 15 to remove poisons such as lead (Pb), arsenic (As) and mercury (Hg). These are poisons of known catalysts and the bed protectors will be employed in a known manner to protect the catalyst from catalytic distillation. Upon entering the catalytic distillation column, 8 to 20% by weight of diene and acetylenic feed is hydrogenated in catalyst beds 16 and 18 located in the rectifying section 20 of the column. Catalytic beds can be of the same or different catalytic composition. The catalysts are known hydrogenation catalysts consisting primarily of one or more metals of group VIIIA (Ni, Pd, Pt) on a support.

Additives such as Ag or Au and / or alkali metals are typically used to control selectivity and activity. Specific examples of selective hydrogenation catalysts particularly suitable for this service are described in U.S. Pat. Numbers 6,417,136; 5,587,348; 5,698,752 and 6,127,588. The catalyst systems employed within the catalytic distillation column may consist of either a single catalyst, a catalyst with different metal charges to adjust the activity located in different portions of the column or catalyst mixtures of different metals located in different portions of the column. the spine. Hydrogenation occurs in the liquid phase in the form of catalytic distillation. Although only two reactive catalytic beds 16 and 18 are illustrated, this is only by way of example and can be any number of beds depending on the requirements of any particular plant or the desire to adjust the catalytic activity through the use of catalyst systems more complex. The internal fractionation components 22 and 24, which may be trays or packages, are provided in the rectification section 20.

Additional internal fractionation components can be located between the catalyst beds 16 and 18. The extraction section 26 contains internal fractionation components 28. The evaporation product section 42 of the column is cooled in the evaporation product section condenser 44 with cooling or cooling water as required and the resulting vapor or liquid are separated in the reflux drum 46. The resulting liquid from the reflux drum 46 is fed through line 48 back to the column as reflux. Similar to the prior art, the vapor of the evaporation product section 50 contains the majority of C5 and lighter components, while the liquid phase 48 is used for reflux of the column. The vapor of the evaporation product section 50, however, does not pass to the subsequent fractionation but to a fixed bed reactor system consisting of one or more catalyst beds supplied with heating and / or cooling of the steam feed. The evaporation product section 50 is first exchanged against the final fixed-bed reactor system effluent 74 to recover heat. It then passes to the heater 66 where the temperature of the steam entering the first fixed bed reactor 68 is controlled. In reactor 68, some portion of the C2 acetylene as well as some portion of C3 and heavier diethylenes and dienes that were not converted to the catalytic distillation column, are hydrogenated. The conditions and the number of fixed-bed reactors employed are such that C2 acetylene is completely removed from the effluent stream 74 without loss of ethylene and propylene over the entire system (catalytic distillation plus fixed-bed reactors). The addition of the fixed bed reactor system to the catalytic distillation column dramatically increases both the performance of the entire system and the ability for that system to respond to process variations and catalytic deactivation. The operating criteria for the rectification section of the catalytic distillation column is that conditions are created where the unsaturated hydrocarbons are hydrogenated to the extent possible without any hydrogenation of ethylene and propylene. This is achieved by: 1. Operation of the column in such a way that ethylene and propylene in the liquid phase are minimized, and 2. Operation of the catalytic distillation column in such a way that even without converting C2 to C5 acetylenes and diolefins remaining in the section of evaporation products 50 of the column. In the catalytic distillation operation of the present invention, the distillation function is designed and operated to distill essentially all of the C5 and lighter components as products of the evaporation section and essentially all of the C6 and heavier components, as background. Alternatively, the separation may be in the carbon number C where essentially all of the C4 and lighter components pass as evaporation products and the C5 and heavier components come out as bottoms. In order to selectively hydrogenate the C2 acetylenes, the C3 acetylenes and dienes, and the C and heavier acetylenes, dienes and olefins while leaving unhydrogenated ethylene and propylene, the rectification section 20 is operated in such a way that there is a gradient of Substantial concentration of materials C4 and C5 with respect to materials C2 and C3 in the liquid phase where the majority of the hydrogenation reaction occurs. This can be controlled by varying the performance of the kettle and the reflux expense to achieve the composition of evaporation products and bottoms desired. The operation selection of the catalytic distillation column either as a depentanizer or a debutanizer, will be a function of both the composition of the feed and the desired hydrogenation requirements for the products. The preferred operating conditions for a depentanizer will be a pressure of between 5.2 and 24.1 bars gauge (75 and 350 psig) and a catalyst bed temperature between 50 and 150 ° C. Similarly, the preferred operating conditions for a debutanizer column will be a pressure between 7 and 28 bars gauge (100 and 400 psig) and a catalytic bed temperature between 30 and 130 ° C. In addition to controlling total fractionation, the temperature and composition profiles on the reaction sections can be controlled by adjusting the heat removal ratios on the column and by recirculating liquid into and / or around the catalyst beds. As illustrated in Figure 2, the trays and 31 collect the descending liquid that is withdrawn as side streams 32 and 34. These streams may or may not pass through the intercoolers 36 and 38 and then be re-injected into the column through the distribution heads 40. This allows a portion of the heat of reaction to be removed in the intercoolers. By arranging the chillers in this manner, the cooling medium can be water, while cooling in the evaporation product condensers may require to be at least partially provided by mechanical cooling. Therefore, the use of intercoolers can significantly reduce the portion of the heat of reaction that needs to be removed by mechanical cooling.

The hydrogenation in column 14 occurs in the liquid phase. The extent of the reaction depends on the relative reactivity of the various components and the concentration of these components in the liquid phase at any particular point in the column. The C2 and C3 acetylenes and dienes are much more reactive than ethylene and propylene, so that they react first and quickly. However, the relative reactivities of ethylene propylene and the heavier olefins, dienes and acetylenes C4 are very close. In order to react a significant amount of C4 and heavier olefins, dienes and acetylenes, without significant loss of ethylene and propylene, the concentration of ethylene and propylene in the liquid phase must be minimized and the upper temperature and concentration profiles in the background they should be controlled. Since this stage of hydrogenation occurs in a fractionation column, this control can be achieved by adjusting the reflux of the section of evaporation products produced by the condenser of the section of evaporation products 44 and the reflux of the sidestream of the evaporators. intercoolers 36 and 38. The liquid compositions of ethylene and propylene can be kept low in the reactive zones through increases in reflow flow 48 and / or increased cooling between beds in 36 and 38. In the catalytic distillation unit 14 , the circuits of recycling and auxiliary reflux, have been modified from the prior art as illustrated in Figure 1. This prior art shows simple inter-chillers 180 and a total auxiliary reflux line 260. In the present invention, the system is modified to provide the flexibility to have both auxiliary fluids without cooling and cooling in the catalyst zones 16 and 18 within the sec rectification 20. This improvement allows the desired control of temperature and composition with minimal disturbance to the total distillation. This is achieved by removing auxiliary reflux liquid immediately below the catalyst beds as stream 52 and / or 54 from the extraction points 53 and 31 respectively and returning them through the pump 56 and the exchanger 58 to the top of the same bed as stream 60 and / or 62. Alternately, the liquid can be removed from the catalyst bed further downstream and returned to the higher bed by stream 62. Cooling at 58 can be used, if necessary to provide fine-tuning. composition and interenfriamiento combined between reactive beds. For example, while being removed from the intercooling stream 34 at point 31, cooling that stream in the exchanger 38 and returning the flow to the distribution system 41 will cool the liquid but will not change the composition. However, removing the same liquid 31 and passing it through line 54 through the pump 56 to the exchanger 58, cooling the liquid, and returning it to the liquid distributor 40 on the catalyst bed will change the composition profile within the column . This design flexibility can be used to maximize the efficiency of hydrogenation.

In this way, the option of cooling against a hotter cooling medium available in the prior art, it is maintained with the modified inter-cooler / auxiliary reflux reducing the low-level costly cooling required in the evaporation product system. In addition, the heat removed by these auxiliary reflux streams can be used elsewhere in the ethylene plant to reduce energy consumption. Another advantage of the new auxiliary reflux scheme is that it allows a relatively large liquid flow without affecting the separation performance of the total column due to the heavier ones in the section of evaporation products as in the prior art. With the large liquid flow the auxiliary reflux can provide the necessary liquid charge on the catalyst without need for further reflux. This allows operation of the catalytic distillation column at lower reflux ratios than previously possible without the penalty in the distillation efficiency observed with the prior art. The reflux ratios in the range of 0.5 to 1.8 by weight are satisfactory to produce the necessary liquid catalyst wet where values as high as 5 were required with the prior art. In addition to the obvious reduction in energy requirements, higher partial pressures of hydrogen due to lower reflux ratios will be available in the present invention resulting in lower volumes of catalyst required. In a catalytic distillation column, it is critical to keep the catalyst moist at all times to ensure that the reaction occurs in the liquid phase. The selectivity of a catalytic distillation system is partly based on the reaction being carried out in the liquid phase while certain components that the operator wishes to remain unreacted such as ethylene, remain at the highest concentration in the vapor phase . Maintaining some liquid traffic through the column is critical to keep the catalyst moist. If the liquid traffic is greater than .39 kg / hour / cm2 (800 pounds of liquid / hour / foot2) of column cross section, the catalyst will be highly wetted and the selectivity of the reaction will be maintained. A secondary control variable would be the variation in reflux with the associated variation in the performance of the kettle. In this way, both the composition and the temperature of the catalyst bed can be altered to achieve the desired hydrogenation. Additionally, a variable feed location that allows a main feed point below the extraction section 22, will provide some separation of any heavy components present in the feed before reaching both the catalyst bed 16 and the side stream 52 for the first auxiliary reflux. In this way, circulating the heavier, potentially incrusting components on the catalyst bed is eliminated. In addition, feed points on the first catalyst bed can be incorporated to allow rejection operation and this mache to avoid the problems of excess catalyst and loss of selectivity resulting under these conditions of lower flow. Funds 63 in column 14 are sent for further processing as desired. As illustrated in Figure 2, the present invention includes the addition of a fixed-bed reactor system that provides additional hydrogenation of stream 50. This system typically consists of two reactors with an intercooler, but may be a series of reactors with an inter- Coolers between successive reactors. The fixed-bed reactor system provides four advantages: 1. The catalytic distillation column does not require further operation at high levels of hydrogenation but can be operated for maximum productivity from the catalyst, a net gain of ethylene with high acetylene conversion, methyl acetylene and propadiene while maintaining the C2 acetylene specification. 2. Changes in the concentration of catalytic distillation evaporation products of acetylenics and dienes resulting from catalyst deactivation, increase in carbon monoxide content or change of feed material, can be tolerated. 3. Hydrogen removal can be maintained during temporary misalignments and / or catalyst deactivation or poisoning, thus stabilizing the performance of downstream cooling systems. If the amount of hydrogen in the system would vary, the partial pressures of the downstream distillation system will change and the required amount of refrigeration will vary. This would create disturbances in the process and it would be very undesirable. 4. Opportunity for catalyst regeneration by the use of replacement fixed bed reactors, thus extending the operating lifetime in current of the entire system. In addition to controlling the profile of temperature and composition in the column, it is important to operate with less than the complete conversion of the acetylenes and dienes in the catalytic distillation column. By doing so, the ethylene and propylene gains can be achieved. In addition, this operation requires less catalyst than the complete hydrogenation of the prior art, thereby maximizing the productivity of the catalytic distillation catalyst. The operation with a fixed bed reactor system following the column allows this to occur. If the column were to be operated in such a way that there was no more than about 1% concentration of liquid ethylene in the reactive beds, hydrogenation in excess of 95% C2 to C5 and heavier dienes can be achieved. This results in 5000-7500 ppm of dienes and acetylenics in the vapor stream 50 of the reflux drum 46 and a minimum ethylene loss of 1%. To achieve 100% conversion of acetylene, the ethylene losses would be even higher. This operation coincides with a hydrogen removal of approximately 30-35 percent depending on the feed composition. However, when the total conversions of the C2 to C5 dienes and acetylenics and heavier are reduced between 80 and 95% resulting in 10,000 to 20,000 and typically 15,000 ppm of dienes and acetylenes C2 to C5 in the output stream 50, it can be achieved ethylene gain. Figure 3 is a trace of the dienes at the outlet in ppm against, the gain or loss of ethylene in percent by weight for a single catalytic distillation unit (CDU) and for a CDU plus a fixed-bed hydrogenation system. As can be seen in Figure 3, allowing a certain amount of highly reactive acetylenes and dienes to remain unreacted in the evaporation products section of the catalytic distillation column, the ethylene and propylene losses can be eliminated while still obtaining 100% or total acetylene conversion. With fixed bed reactors located after the catalytic distillation column 14, the start of C2 acetylene with 10,000 to 55,000 and typically 20,000 ppm combined C3 dienes and heavier acetylenes can be tolerated from the catalytic distillation column. A typical system with two fixed-bed hydrogenation reactors with inter-cooling, has been shown to hydrogenate 100 percent of the C2 acetylenes entering the fixed-bed reactor system and approximately 75% of the C3 and the heavier dienes and acetylenes combined enter the fixed-bed reactor system. This results in 2500 to 14,000 and typically only 5000 ppm of starting dienes and acetylenics of the combined system. This represents approximately 97% hydrogenation of the total C2 and heavier acetylenes and diolefins in the feed. This operation allows total ethylene total gains of up to .5%, with 70% selectivity of total acetylene to ethylene at 100% conversion of acetylene. This is a substantial improvement over the previous technique.

The specific hydrogenation reactivity of ethylene is only slightly less than the specific reactivity of propadienes. In this way, a close observation of the conversion of C3 diene allows a reliable indication of the stability of the ethylene gain and can be used as a control point for the system. For the catalytic distillation system alone, when the conversion of C3 diene is between 40 and 60% and typically 45%, ethylene losses are observed. However, when operating to conditions where the conversion of C3 diene into the catalytic distillation column is between 10 and 35 and typically 20%, ethylene gains are possible from 0.2 to 0.5%. With the present invention, the conversion of propadiene can be substantially increased while the ethylene gain is still maintained. During the normal operation of an ethylene unit, variations in the carbon monoxide content of the charge gas 10 are experienced. In addition, the quality of the feedstock and the severity of operation impacting the acetylene and diolefin content of the gas can be changed. charge gas. For a fixed catalyst volume in the catalytic distillation column, increases in carbon monoxide or diene and acetylenic concentrations inlet in the lower conversion and thus higher detachments of these unwanted products in stream 50. Compensation for these disturbances anticipated will be difficult with the previous technique alone as illustrated in Figure 1. It would require increases in operating pressure or temperature that impact the performance of the entire fractionation system. In the improved process including a fixed bed reactor system, the temperature of the vapor 50 entering the fixed echo reactor system can be adjusted either to increase or decrease the reactivity of the reactor system and thus follow changes in the activity of catalytic distillation reaction and maintain a complete C2 acetylene removal and high hydrogen removal efficiency. Finally, a fixed-bed hydrogenation reactor system is designed to include not only operating but also replacement reactors.

Catalyst deactivation will occur both in the fixed bed system and in the catalytic distillation system. It is not possible to regenerate the catalytic distillation catalyst without stopping the process or installing a column in parallel. Both options are expensive. However, a replacement fixed-bed steam phase reactor is a relatively inexpensive option. By using a fixed bed reactor system with a spare instead of the single column concept of the prior art, the operational life in the process stream can be substantially improved. In the fixed-bed hydrogenation system, the net evaporation products section 50 of the catalytic distillation passes through the cross-flow heat exchanger 64 and the inlet heater 66 to the first fixed-bed reactor 68. The effluent of the first reactor 68 passes through the intercooler 70 to the second fixed bed hydrogenation in the reactor 72. A series of fixed beds followed by intercoolers can be used in the same way to achieve the necessary heat transfer when required. The effluent from the last reactor 72 then passes back through the cross reflux heat exchanger 64 where the heat is removed and the feed 50 to the fixed bed reactors is heated. The inlet temperature of the fixed-bed reactors can be changed quickly either to increase or decrease the hydrogenation extension in the fixed-bed reactors. This control is necessary to successfully manage changes in the concentration of carbon dioxide or diene and acetylene feed. Up to a maximum adiabatic temperature increase of 26.7 ° C (80 ° F) in total for both beds, a stable fixed bed operation without loss of ethylene is possible. A typical adiabatic increase of 1.66 ° C (35 ° F) is expected for normal operation. With an adiabatic temperature increase of 21.1 to 26.7 ° C (70 to 80 ° F) and typically 26.7 ° C (30 ° F), the handling of 35,000 to 58,000 and typically 43,000 ppm of acetylenes and dienes from catalytic distillation results in 9,000 to 30,000 and typically 10,000 ppm of C3 and heavier dienes and acetylenics in the final product stream 74, while maintaining 100% of C2 acetylene conversion primarily in ethylene. In a similar way, the temperature control at the inlet of fixed-bed reactors can allow compensation for catalytic deactivation by providing the typical start-of-operation and end-of-operation operating temperatures to the fixed-bed system. In the prior art, this can only be done by a temperature correction in the catalytic distillation column. This requires a change of pressure in the column and in this way the fractionation conditions will be altered. With the system of both the catalytic distillation column and the fixed bed reactor of the present invention, the catalytic distillation column can operate at constant fractionation conditions and lower temperature corrections for the fixed bed system will be required. This improves the stability of the system and allows longer catalyst life. Figure 4 presents an alternate embodiment of the present invention. Instead of the stream of evaporation products from the catalytic distillation column 42 that passes to the exchanger 44 and then to the reflux drum 44, the stream of evaporation products 42 is passed directly to the cross-flow exchanger 64 and into the reactor system of fixed bed. Following the fixed-bed reactor system, the effluent is cooled to 65 and the reflux 48 for the column is separated at 67 as a condensed liquid 69 and returned to the column. Since the current entering the fixed-bed reactor system still contains all the reflux for the column, the operating temperature of the fixed-bed reactors will be somewhat higher to ensure full steam flow.

This will change the design catalyst activity and the space velocity to ensure stable operation. The advantage of this approach will be an increased mass flow of hydrocarbons that will minimize the increase in temperature through the fixed beds, a reduced partial pressure of hydrogen that improves the selectivity and superior space velocity that both improve the selectivity and decrease the costs of catalyst. Figure 5 illustrates an alternative embodiment of the present invention that incorporates a pre-reactor. This arrangement is advantageous for selective hydrogenation in volume of feeds with high content of dienes and acetylenes. After compression at 12 and possible treatment in a guard bed or pre-bed (not shown), the vapor phase feed material is mixed with recirculating liquid 76 from pump 56 of column 14 and the mixture of two The phases pass concurrently through a fixed bed reactor 78. Hydrogenation occurs and the presence of liquid serves to control the increase of temperature through evaporation. The hydrogenation reactor 78 can be designed as an operation reactor plus a spare to allow extension of the operation in system current. Following the pre-reactor, the liquid / vapor mixture can already be sent to the column directly as a mixed or separated feed in a separating drum and the liquid and steam fed separately to the column. The latter is preferred since any oligomers formed in the initial hydrogenation will be in the liquid phase and can be fed to the column below the catalyst beds thereby reducing fouling. By carrying out the hydrogenations of the fixed bed before the catalytic distillation column 14, it will allow possibly higher catalyst utilization without experiencing loss of ethylene for that portion of the hydrogenation due to the large amount of preferentially absorbed higher reactivity dienes and acetylenes, available for hydrogenation . To a higher catalyst utilization, smaller volumes of catalyst will be necessary making the process more economical. A catalytic distillation unit is still required after a pre-reactor to reach the hydrogenation specifications. It is anticipated that a maximum of 50% and typically 20% of the hydrogenation yield can be achieved in the pre-reactor. Another advantage of a fixed bed hydrogenation reactor before the catalytic distillation column 14 is that the reactor can be used as a guard bed or pre-bed for catalyst poisons. The catalyst can already be nickel or palladium. A nickel catalyst for example will be able to catalyze the reaction of the sulfur compound thiophene with butadiene to form a heavy mercaptan. This mercaptan will then be removed in the extraction section 22 of the column 14 and in this way will never contact the palladium catalyst. An even further advantage is that the external pre-reactor system can have a spare and thus allow regeneration without requirement by interrupting or shutting down the entire plant for replacement of the catalyst. Alternately, as shown in Figure 6, the liquid 76 of the pump 56 can flow down through the fixed bed 78 and the vapor stream of the compressor 12 can flow upwards. Liquid from the bottom of the fixed bed reactor 78 then flows to a lower portion of the column 14 and the steam flows to a higher inlet point. The advantage of this countercurrent process sequence is that the oligomers resulting from the polymerization reactions of the unsaturated hydrocarbons are removed from the catalyst bed as they are formed and do not pass over the remaining portion of the catalyst bed.

Also, this liquid is sent to column 26 at a lower entry point, reducing any potential contamination of the catalyst in column 14. Oligomers that can embed the catalytic distillation catalyst are easily separated and do not rise in the column to contaminate the catalyst . further, as in the parallel flow option, the pre-reactor catalyst bed can have a spare, allowing regeneration while the rest of the system is operating. The ability to easily regenerate online will increase system cycle lengths since the catalyst zone in the feed inlet is expected to have the highest inlay rate. To minimize fouling in the fixed-bed pre-reactor, the liquid flow costs need to be sufficient to minimize local hot spots due to the high hydrogenation heat and wash away any oligomers that are formed from the catalyst. The operation of these beds is preferably in the continuous vapor zone. For pyrolized gas feeds that exhibit extreme incrustation tendencies, operation in the continuous zone of the liquid is also possible. Figure 7 illustrates a further embodiment of the present invention that incorporates fixed bed reactors within the auxiliary liquid reflux or intercooler streams that are removed from the column 14. Fixed-bed hydrogenation reactors 82 and 84 are placed in the lateral stream of the collection tray 30 and the lateral stream of the collection tray 31, respectively. These fixed beds 82 and 84 are in addition to the reactive hydrogenation sections 16 and 18 in the hydrogenation sections 16 and 18 in the catalytic distillation column 14. A mass transfer zone 82 in the form of structured packages or trays is also add on the extraction point and below the catalyst bed. This zone allows the hydrogen to saturate in the liquid phase and thereby provide the hydrogen required for the hydrogenation of the acetylenes and dienes in the extraction liquid. The ability of the present invention to remove 30 to 40 percent of the hydrogen from the charge gas before the cooling and condensing steps reduces energy consumption and capital cost. The ability to hydrogenate 100 percent of the acetylene independently of the carbon monoxide concentration without loss of C2 or C3 olefins was not possible with the prior art. The combined catalytic distillation and fixed bed stages provide superior handling of system disturbances while maintaining a stable hydrogen removal and acetylene / diene hydrogenation. Below are some examples that illustrate the present invention in its various modalities compared to the prior art. The following Table 1 establishes the feed composition used for all the examples. Table 2 cites the results for each of the examples. Table 1 Table 2 Example 1: This example represents the prior art established in the US patent. No. 5,679,241 (Figure 1) based on a one stage catalytic distillation column operating at a reflux ratio of 4.4. With a typical initial acetylene hydrogenation catalyst containing palladium levels below 2000 ppm and operating at a pressure of 13.44 bar man. (195 psig) and average catalyst temperature of 110 degrees C (230 degrees F), the acetylene conversion reached 84 percent with 0 percent ethylene loss / gain. At the reactor outlet, there was 370 ppm of C2 acetylene and a total of 19070 ppm of dienes and acetylenics. Total conversion of acetylene / diene is 79.5. This example represents the case where the single column is operated for no loss of ethylene. As can be seen, there is substantial loss of C2 acetylene. This will produce product outside of the specifications. Example 2 This example also represents the prior art and is based on the simple catalytic distillation column of Example 1 with higher catalyst temperature and slight lower reflow ratio of 4.1. The hydrogenation severity of a single column can be increased to achieve low C2 acetylene levels. This can be achieved by raising the temperature or increasing the catalyst activity. The operation at higher temperature is intended to reduce the acetylene content and in this way achieve ethylene product that meets the specifications. In this case, all diene and acetylenic conversions are superior compared to Example 1, however there is also a 0.6 percent loss of ethylene. At the reactor outlet there were 240 ppm of C2 acetylene and a total of 12340 ppm of dienes and acetylenics. A total conversion of acetylene / diene is 86.7 percent. As can be seen, the increased conversion of C2 acetylene is accompanied by an increased loss of ethylene which is economically undesirable. In addition, the ethylene product still does not meet the specification limits of 1-2 ppm. Example 3 This example represents the prior art and is based on the simple catalytic distillation column of Example 1, with higher levels of carbon monoxide in the feed. The carbon monoxide acts as a catalyst poison and therefore the diene and acetylenic conversions were substantially reduced. At carbon monoxide levels of 0.1 mol percent in the feed (0.6 percent ratio of carbon monoxide to hydrogen) the product had 460 ppm of acetylene and 33860 ppm of total dienes and acetylenics. The lower catalyst activity resulting from the CO reflects in the catalyst productivity loss .00153 - .00114 kg mol / hr-m3 (0.12 to 0.09 ibmol / hr-ft3) of catalyst structure and the lowest total conversion of acetylene / diene (63.6 vs. 79.5 percent for the base case). The response to this reduced activity would be to raise the temperature of the catalyst within the catalytic distillation column. This will require an increase in pressure beyond what is practical in a unit of operation. In this way, the options for compensating the increases in CO are limited for the prior art. Example 4 This example represents the improved combined operation of the catalytic distillation column and a fixed bed reactor described in Figure 2. This combined operation is necessary to achieve ethylene gains with 100 percent C2 acetylene conversion and 50 to 95 percent percent conversion of all the other compounds diene and acetylene. The operation of the catalytic distillation column at 13.44 bars gauge (195 psig) and an average catalyst temperature of 110 degrees C (230 degrees F) and pressure of 13.44 bars gauge (195 psig) resulted in 12,000 ppm by weight of dienes and acetylenes in the section of catalytic distillation evaporation products that are then fed to the fixed-bed hydrogenation reactor system.

The operation of the fixed bed hydrogenation reactors at a gas hourly space velocity (GHSV) of 1800 h "1 and at the bed entry temperature of 53.9 degrees C (129 degrees F), was successful in converting 100 percent of the C2 acetylene and providing sufficient additional hydrogenation resulting in 50 percent total conversion of C3 diene (both fixed and catalytic bed hydrogenation) as well as 96.1 percent total conversion of dienes and acetylenes in the combined system feed. in 0 ppm of C2 acetylene and 3640 ppm C3 and heavier dienes and acetylenes at the outlet The ethylene product of specification can be produced with very high conversion of total of the highly unsaturated species Example 5 This example represents the improved combined operation of the catalytic distillation column and a fixed bed reactor at high levels of carbon monoxide in the feed. rbono of 0.1 in mol percent in the feed with constant operating conditions for the catalytic distillation column, only the inlet temperature of the fixed bed needs to be adjusted to maintain product specifications. Specifically, with an increase in carbon monoxide of 0.05 to 0.1 mole percent, an increase in the temperature in the fixed bed at the inlet of 53.9 to 60 degrees C (129 to 140 degrees F) was sufficient to maintain the conversion of C2 acetylene. In addition, C3 and heavier dienes and acetylenes were more hydrogenated resulting in 7740 ppm in total weight, of dienes and acetylenes in the product. The improvement proposed in the present invention will make 100 percent C2 acetylene hydrogenation with stable hydrogenation 90 percent more C3 to C5 and heavier acetylenes, 90 percent more hydrogenation of C and C5 dienes, and 50 percent more conversion of C3 dienes in a feed stream without hydrogenating the C2 and C3 olefins. The resulting hydrogen removal with the present invention will remain stable at 30 to 40 and typically 30 percent depending on the composition of the feed.

Claims (29)

1. CLAIMS 1. A method to process a thermal catalytic disintegration feed stream containing hydrogen, ethylene, propylene, acetylene, methyl acetylene, propadiene and other C, C5, C6 and heavier unsaturated hydrocarbons to hydrogenate and more essentially convert all acetylene with high proportion in ethylene and hydrogenating at least a portion of the methyl acetylene, propadiene and other C4, C5, C6 and heavier unsaturated hydrocarbons in olefins and thus consuming a portion of the hydrogen without hydrogenating ethylene and propylene, characterized in that it comprises the caps from: a. introducing the feed stream into a catalytic distillation column containing at least one bed of hydrogenation catalyst and concurrently: (i) selectively hydrogenating a portion of the acetylene to form ethylene and hydrogenating portions of the methyl acetylene, propadiene and C, C5, C6 and heavier unsaturated hydrocarbons and will control hydrogenation conditions whereby ethylene and propylene are not hydrogenated; and (ii) fractionally separating the feed stream into lighter hydrocarbons and heavier hydrocarbons; b. substantially removing all remaining portions of the hydrogen and lighter hydrocarbons as vapor phase evaporation products and substantially all of the heavier hydrocarbons as bottoms of the catalytic distillation column; c. introducing at least a portion of the vapor phase evaporation products into a fixed vapor phase reactor system containing a hydrogenation catalyst and hydrogenating the remaining portion of the acetylene to further form ethylene and hydrogenating larger portions of methyl acetylene, propadiene and C, C5, C6 and heavier unsaturated hydrocarbons and control the hydrogenation conditions whereby ethylene and propylene are not hydrogenated; and d. Remove the mixed product from the fixed bed reactor system.
2. 2. A method according to claim 1, characterized in that the lighter hydrocarbon consists of C4 and lighter hydrocarbons and the heavier hydrocarbon consists of C and heavier hydrocarbons.
3. 3. A method according to claim 1, characterized in that the lighter hydrocarbon consists of C5 and lighter hydrocarbons and the heavier hydrocarbon consists of C6 and heavier hydrocarbons.
4. 4. A method according to claim 1, characterized in that it further comprises the steps of condensing a portion of the evaporation products in the vapor phase and returning the condensed portion to the catalytic distillation column, as reflux.
5. 5. A method according to claim 1, characterized in that essentially all the vapor phase evaporation products are introduced to the fixed-phase steam-bed reactor system and further comprises the step of returning a product portion of the bed reactor system. fixed to the catalytic distillation column as reflux.
6. 6. A method according to claim 1, characterized in that the step of controlling the step of selective hydrogenation in the column of catalytic or distillation and in the fixed-bed reactor in the vapor phase, includes the steps of controlling the temperature profile.
7. A method according to claim 6, characterized in that the step of controlling the temperature profile in the catalytic distillation column includes the step of removing a portion of liquid descending from the column at a selected point as a side stream. , cool the lateral current and inject the cooled side stream back to the column at or above the selected point.
8. 8. A method according to claim 7, characterized in that it also includes the step of hydrogenating the side stream.
9. 9. A method according to claim 6, characterized in that the step of controlling the temperature profile in the catalytic distillation column, includes the step of removing an auxiliary reflux stream from a point in the column below the hydrogenation catalyst beds, cooling the auxiliary reflux stream and injecting the cooled auxiliary reflux stream back into the column on the beds of hydrogenation catalyst.
10. A method according to claim 1, characterized in that the fixed-bed reactor system in vapor phase comprises at least one reactor and wherein the step of controlling the temperature comprises the step of controlling the temperature in the exchangers before the reactor.
11. 11. A method according to claim 1, characterized in that the fixed-bed steam reactor system comprises two or more reactors in series and wherein the step of controlling the temperature comprises the steps of controlling the temperature in thermo exchangers before each of the reactors.
12. 12. A method according to claim 1, characterized in that the step of selective hydrogenation consists of the step of operating the catalytic distillation column in such a way that the concentration of ethylene in the liquid phase in the catalyst bed is less than 2 times. cent in weight.
13. 13. A method according to claim 1, characterized in that the step of selective hydrogenation comprises the step of operating the catalytic distillation column in such a way that the liquid flow through the column is greater than .39 kg / hr / m2 ( 800 Ib / hr / ft2) of cross-sectional area in the catalyst bed area.
14. 14. A method according to claim 1, characterized in that the step of introducing the feed stream into the catalytic distillation column includes the step of mixing the feed stream with a recycle fluid from the catalytic distillation column and introducing the mixed feed stream and the recycled liquid in a fixed-bed hydrogenation pre-reactor before the catalytic distillation column whereby a portion of the highly unsaturated hydrocarbons is hydrogenated and the vapor and liquid streams are introduced into the column of catalytic distillation.
15. 15. A method according to claim 1, characterized in that the hydrogenation catalyst bed in the catalytic distillation column contains a catalyst of a group VIIIA metal in a support.
16. 16. A method according to claim 12, characterized in that the catalyst comprises palladium in alumina.
17. 17. A method according to claim 13, characterized in that the catalyst also includes an additive selected from the group consisting of gold, silver and alkali metals.
18. 18. A method according to claim 14, characterized in that the catalysts having different amounts of palladium are located in select portions of the catalytic distillation column.
19. 19. A method according to claim 15, characterized in that different catalysts are located in different portions of the catalytic distillation column.
20. 20. A method according to claim 19, characterized in that the different catalysts contain different metals.
21. 21. A method according to claim 19, characterized in that the different catalysts have different metal charges.
22. 22. A method according to claim 15, characterized in that the catalyst comprises nickel in a support.
23. 23. A method according to claim 15, characterized in that the catalyst comprises a combination of palladium in a support and nickel in a support in different portions of the distillation or catalytic column.
24. 24. A method according to claim 1, characterized in that the hydrogenation catalyst in the fixed-bed reactor system comprises a metal of the VINA group on a support.
25. 25. A method according to claim 24, characterized in that the hydrogenation catalyst in the fixed-bed reactor system comprises palladium in alumina.
26. 26. A method according to claim 24, characterized in that the hydrogenation catalyst in the fixed-bed reactor system comprises palladium in alumina with a promoter consisting of gold, silver, and alkali metal or combinations thereof.
27. 27. A method according to claim 14, characterized in that the fixed-bed hydrogenation pre-reactor contains a nickel catalyst and the pre-reactor causes the reaction of sulfur compounds for removal.
28. 28. A method according to claim 1, characterized in that it also comprises the step of removing catalyst poisons from the feed stream before introducing it into the catalytic distillation column.
29. 29. A method according to claim 28, characterized in that the catalyst poisons are lead, arsenic and mercury.