JP4376908B2 - Improved olefin plant recovery system using a combination of catalytic distillation and fixed bed catalytic process - Google Patents

Description

  The present invention relates to a process for the production of olefins, and in particular to treating cracking heater effluent to more efficiently recover product and treat by-products.

In the production of ethylene and propylene via pyrolysis of various raw materials, several by-products and unsaturated diolefins and acetylene are produced. The final effluent from the pyrolysis heater (typically referred to as charge gas) requires processing to separate by-products from the primary olefin product and to remove diolefins and acetylenes. . Removal of C ₂ and heavier diolefins from the cracking gas is addressed through a combination of separation by distillation and reaction by hydrogenation. In particular, for acetylene, separation alone results in excess ethylene product loss because acetylene and ethylene have very similar volatility. Currently, distillation and hydrogenation are performed in several separate process steps designed to separate and hydrogenate C ₂ , C ₃ and C ₄ compounds independently. In order to achieve good control in hydrogenation, extended catalyst life, and increased performance, it is now necessary to separate various hydrocarbons prior to hydrogenation.

  One of the disadvantages of this widely practiced technique is that it generates energy from the cracked gas, first generating hydrogen and then the high pressures and temperatures required to separate the higher carbon molecules. The amount of consumption is large. In addition, the hydrogenation process for each group of hydrocarbons requires a separate reactor system consisting of several devices, which increases plant capital and complexity.

The invention described in US Pat. No. 5,679,241 proposes a one-step conversion process that converts all of C ₂ -C ₅ and heavier acetylenes and dienes without hydrogenating C ₂ or C ₃ olefins. Is. According to the claims, this process is carried out by one catalytic distillation unit which can handle hot, relatively low pressure charge gas before excessive compression and cryogenic cooling takes place. In addition, if necessary, this one-step process can hydrogenate C ₄ olefins again to paraffins without causing loss of C ₂ or C ₃ olefins. The patent, previously required cryogenic separation, C ₂ -C ₄ system by hydrogenation to acetylenes and dienes and C4 and heavier olefins paraffin can be removed more than 70% of hydrogen in the cracking gas About. Being able to remove more than 70% of the hydrogen can significantly reduce the energy consumption for separating C ₂ and heavier components, improving economy. By reducing the hydrogen partial pressure, the separation is performed at a lower pressure and the refrigeration power can be reduced. However, it has been demonstrated that extensive hydrogenation in such single-step systems does not occur substantially without loss of ethylene and propylene to paraffins due to hydrogenation.

The method disclosed in US Pat. No. 5,679,241 has significant limitations. First, C ₂ acetylene is removed particularly through hydrogenation because it is extremely difficult to remove C ₂ acetylene by distillation in operation at an ethylene plant and requires significant equipment and energy costs. That must be done. Since acetylene is a catalyst poison, it must be removed to a low level, often 1-2 ppm. Hydrogenating all of the C ₂ acetylene in a single catalyst distillation column without loss of ethylene or preferably increasing ethylene was not possible with reasonable catalyst volume and economically viable operating conditions. Second, while the carbon monoxide flow rate (which affects catalyst activity) and / or the diene / acetylene concentration in the feed varies, it is difficult to manage to maintain process performance and on a commercial scale Achieving is difficult. Third, there are limited methods for dealing with ultimate catalyst deactivation. These units must be operated for a long period of time between outages, so that the catalyst is in excess and the catalyst zones are separated in the tower (separated individually, The other zones remain operational and the catalyst is replaced) is the only option. When using larger amounts of catalyst, it is known that it is necessary to operate at lower temperatures to avoid overreaction while still maintaining the catalyst active. This adversely affects economics because some refrigeration is required to control operation at lower temperatures and / or excessive circulation of cold liquid in the column. Specifically, the pilot test showed the following:
a. When using a single catalytic distillation column to remove more than 95% of the C ₂ acetylene, the ethylene loss that occurs simultaneously is 1% by weight or more. This is economically undesirable.

b. When operating a single catalyst distillation column, significant hydrogen loss is expected with currently available catalysts if the C ₄ olefin hydrogenation is 20% or more.

  c. To operate a single catalyst distillation column and maintain a very high degree of conversion and hydrogen removal while at the same time achieving minimal ethylene and propylene losses, an excess of catalyst (proven by low productivity) and lower temperatures Operation (requiring freezing) is required.

  d. As the carbon monoxide varies in the feed, significant variations in catalyst activity occur. If such fluctuations are observed in the single-step method, the acetylene removal efficiency decreases, and as a result, the product does not satisfy the standard. This impact on performance due to reduced catalyst activity via CO poisoning is equivalent to the impact on performance due to catalyst aging.

  e. Significant fluctuations in the feed to the ethylene cracking heater result in fluctuations in both acetylene and dienes as well as the hydrogen flow rate. For changes in the hydrogen / reactant ratio, the single step method is limited in its ability to follow such changes. As a result, unless the system is equipped with substantial and expensive excess equipment used to overcome these process changes, acetylene emissions, or overreactions, leading to substandard ethylene products. Cause significant loss of ethylene and propylene.

The present invention relates to an improved method for treating charge gas effluents from the pyrolysis of various feedstocks. The first objective is to achieve essentially complete hydrogenation of C ₂ acetylene without producing significant ethylene and propylene hydrogenation and hydrogenate C ₂ -C ₅ diolefins and acetylene in the feed. To remove most of the hydrogen in the effluent. The improved method achieves the above objectives even when there are variations in carbon monoxide concentration, changes in diene and diacetylene feed concentrations, and deactivation of the catalyst, as well as other anticipated processing condition disturbances. Is done. The present invention relates to catalytic distillation with improved liquid recycle combined with a fixed bed hydrogenation reactor system. Specifically, the operating conditions for catalytic distillation are such that the hydrogenation of C ₂ acetylene preferably produces ethylene without loss of ethylene and propylene, so that the highest hydrogenation of acetylene and dienes is achieved. Maintain or adjust. Maintaining a stable high rate of all conversions of C ₂ -C ₅ acetylene and dienes, with 100% conversion of C ₂ acetylene, under varying process conditions, is a fixed bed hydrogenation system (the rest of the system C ₂ acetylene is again fully hydrogenated without causing significant ethylene or propylene hydrogenation).

  For a better understanding of the present invention, the prior art represented by the method disclosed in US Pat. No. 5,679,241 is briefly described. FIG. 1 is essentially a copy of a drawing from a prior patent simplified to identify only the features relevant to the present invention. The charge gas 150 is compressed and supplied to the catalytic distillation tower 156. Similar to the present invention, this column performs catalytic reaction and distillation simultaneously. The column has a stripping section 158 below the feed position and a rectification / reaction section 160 containing the catalyst beds 166, 168 and 170 above the feed position. Through the intermediate condenser 180, the descending liquid is withdrawn as a side stream and reinjected into the column on the next lower catalyst bed. Part of the reaction heat is removed by an intermediate condenser. The liquid recycle 260 from the stripping section is recycled to the top. This liquid recycle 260 is the bottom portion 262 and / or the liquid portion 264 from the stripping section.

The overhead from column 156 enters condensers 186 and 188 and the partially condensed stream enters separation vessel 190. The product C ₂ -C ₅ vapor overhead 194 (containing ethylene and propylene) is then sent to a subsequent separation process, while the condensed hydrocarbon is utilized as reflux for the column. Since the object of the present invention is to completely remove acetylene from ethylene without causing loss of ethylene entering the column, this must be achieved in this single step (one catalytic distillation column) operation. Don't be. The overhead vapor stream enters an additional fractionator (not shown) where each carbon number fraction is isolated separately.

Hydrogenation in the catalytic distillation column 156 is performed in the liquid phase. Tower, liquid phase composition is mainly operated to consist C ₅ components. This minimizes the liquid phase concentration of ethylene and propylene, and thus minimizes their reaction. However, the concentration of these two expensive olefins in the liquid is not zero. C ₂ acetylene, C ₃ acetylene and diolefins are more reactive than their olefin counterparts. In the liquid phase, they are preferentially reacted to ethylene and propylene, respectively. However, more reactive C ₂ acetylene is no longer present at significant concentrations in the liquid phase because the reaction is nearly complete with the desired 100% C ₂ acetylene conversion. Under these conditions, ethylene can react to produce ethane (paraffin). This is the point that causes the loss of ethylene.

  There are several options for reducing the reactivity in the catalyst bed in this upper part of the catalytic distillation column. One option is to reduce the temperature. Lowering the temperature in one section is difficult in a single column mode because the distillation temperature is mainly controlled by the column pressure. To achieve this, a second column operating at a lower pressure is required. The second option is to utilize a different catalyst in this upper section. This catalyst is designed to more selectively hydrogenate acetylene and diolefins. Both of these options require a larger volume of catalyst in the column, which increases the size and cost of the column. A third option is to utilize an intercooler as shown for the prior art.

  Referring to FIG. 2, which illustrates one embodiment of the present invention, charge gas 10 is compressed at 12 to 150-250 psig and then fed to catalytic distillation column 14. The charge gas may or may not be preheated to match the tower temperature. The charge gas typically passes through one or more guard beds 15 to remove poisons such as lead (Pb), arsenic (As), and mercury (Hg). These are known catalyst poisons and the guard bed is used in a known manner to protect the catalyst for the catalytic distillation column. Upon entering the catalytic distillation column, 8-20% by weight of the diene and acetylene feed is hydrogenated in catalyst beds 16 and 18 located in the rectification section 20 of the column. The catalyst beds are of the same or different catalyst composition. The catalyst is a known hydrogenation catalyst mainly composed of one or more Group VIIIA metals (Ni, Pd, Pt) supported on a support. Ag or Au and / or alkali metals are typically used to control selectivity and activity. Specific examples of selective hydrogenation catalysts that are particularly suitable for this function are disclosed in US Pat. Nos. 6,417,136; 5,587,348; 5,698,752 and 6,127,588. The catalyst system used in the catalytic distillation column is arranged in a single catalyst, a catalyst with different metal additives to adjust the activity, arranged in different parts of the tower, or in different parts of the tower , Consisting of a mixture of different metal catalysts. Hydrogenation is carried out in the liquid phase by a catalytic distillation method. Although only two reaction catalyst beds 16 and 18 are shown in the figure, this is just one example and shows the catalytic activity through various special plant requirements or the use of more complex catalyst systems. Various numbers of floors can be used depending on the regulation requirements. The rectification section 20 is provided with fractional internals (which are trays or packing) 22 and 24. Additional fractional internals can also be placed between the catalyst beds 16 and 18.

Overhead 42 from the tower is cooled in an overhead condenser 44 by cooling water or, if necessary, by refrigeration and the resulting vapor and liquid are separated in a reflux drum 46. The liquid obtained from the reflux drum 46 is fed back to the column via the line 48 as reflux. As with the prior art, the overhead vapor 50 contains most of the C ₅ and higher light compounds, whereas the liquid phase 48 is used for reflux to the column. However, the overhead steam 50 does not enter the subsequent fractional distillation and enters a fixed bed reactor system consisting of one or more catalyst beds equipped with facilities for heating and / or cooling the steam feed. The overhead 50 is first heat exchanged with the final fixed bed reactor system effluent 74 to recover heat. The heater 66 is then entered, where the temperature of the steam entering the first fixed bed reactor 68 is controlled. In reactor 68, most of the C ₂ acetylene is hydrogenated, along with most of the C ₃ and heavier acetylenes and dienes that were not converted in the catalytic distillation column. The conditions and number of fixed bed reactors used were such that C ₂ acetylene was completely removed from effluent stream 74 without loss of ethylene and propylene throughout the system (catalytic distillation + fixed bed reactor). Selected. The addition of a fixed bed reactor system to a catalytic distillation column dramatically increases both the overall system performance and the ability of the system to accommodate process changes and catalyst deactivation.

The operating criteria for the rectification section of the catalytic distillation column is that the conditions under which unsaturated hydrocarbons are hydrogenated to the extent possible without causing hydrogenation of ethylene and propylene are achieved. this is,
1. 1. Operate the column such that liquid phase ethylene and propylene are maintained; Operating the catalytic distillation column so that overhead 50 still leaves unconverted C ₂ -C ₅ acetylene and diolefin;
Achieved by:

The catalytic distillation operation of the present invention, the distillation function, the distillation substantially distilled all C ₅ and from the lighter components as overhead, substantially components of all of the C ₆ and heavier as bottoms Intended and manipulated. Sorting the C ₄ is performed, substantially components of all of the C ₄ and lighter enters the overhead component of the C ₅ and heavier flows out as a bottom. In order to selectively hydrogenate C ₂ acetylene, C ₃ acetylene and dienes, and C ₄ and heavier acetylenes, dienes and olefins while at the same time leaving ethylene and propylene unhydrogenated, rectification section 20 is used. In the liquid phase where the majority of the hydrogenation reaction takes place, it is manipulated so that there is a substantial concentration gradient of C ₄ and C ₅ material relative to C ₂ and C ₃ material. This is controlled by changes in reboiler capacity and reflux rate to achieve the desired overhead and bottom composition.

  The choice of operating the catalytic distillation column as a depentane or debutane column depends on both the feed composition and the desired hydrogenation requirements for the product. Suitable operating conditions for the depentanizer are a pressure of 75-350 psig and a catalyst bed temperature of 50-150 ° C. Similarly, suitable operating conditions for the debutane tower are a pressure of 100-400 psig and a catalyst bed temperature of 30-130 ° C.

  In addition to controlling the entire fractionation, the temperature and composition profile of the entire reaction section is controlled by adjusting the heat removal rate of the entire column and by liquid recirculation within and / or around the catalyst bed. The As shown in FIG. 2, trays 30 and 31 collect the descending liquid and the liquid is withdrawn as sidestreams 32 and 34. These streams may or may not pass through the intermediate condensers 36 and 38 and are then reinjected into the column via the distribution header 40. Thereby, part of the heat of reaction is removed in the intercooler. By arranging the intercooler in this manner, the cooling medium is water, but the cooling in the overhead condenser needs to be provided at least in part by mechanical refrigeration. Thus, the use of an intercooler can clearly reduce the portion of reaction heat that needs to be removed by mechanical refrigeration.

Hydrogenation in column 14 occurs in the liquid phase. The extent of the reaction depends on the relative reactivity of each major component and the concentration of these components in the liquid phase at any particular point in the tower. C ₂ and C ₃ acetylenes and dienes are much more reactive than ethylene and propylene, and thus react first and rapidly. However, the relative reactivity of ethylene, propylene and C ₄ and heavier olefins, dienes and acetylenes is very close. Without causing significant loss of ethylene and propylene, olefin C ₄ and heavier in order to react a significant quantity of diene and acetylene, the concentration of ethylene and propylene in the liquid phase must be minimized, The concentration and temperature profile from the top to the bottom must be controlled. Since this stage of hydrogenation takes place in a fractionation column, this control is achieved by adjusting the overhead reflux generated by the overhead condenser 44 and the sidestream reflux from the intercoolers 36 and 38.

  In the catalytic distillation unit 14, the recirculation and pumparound circuits are modified from the prior art shown in FIG. The prior art shows a single intercooler 180 and an overall pump around line 260. In the present invention, the system has been modified within the rectification section 20 to provide the catalyst zones 16 and 18 with the flexibility to have both uncooled and cooled pump around. This improvement allows control of the desired temperature and composition with minimal disruption to the overall distillation. This removes the pump-around liquid directly under the catalyst bed from the removal points 53 and 31 as streams 52 and / or 54, respectively, via the pump 56 and heat exchanger 58 to the top of the same bed. Achieved by returning as 62. Alternatively, the liquid can be removed from the lowest catalyst bed and returned via stream 62 to the highest catalyst bed. If necessary, cooling at 58 can be used to provide overall compositional adjustment and intercooling between reaction beds. For example, while the intermediate cooling stream 34 is removed from point 31 and cooled in the exchanger 38 and returned to the distribution system 41, the liquid is cooled, but the composition is not changed. However, the removal of the same liquid from 31, the passage of this liquid by pump 56 through line 54 to exchanger 58, the cooling of the liquid, and the return to liquid distributor 40 on the catalyst bed is the composition in the column. Will change the profile. This design flexibility is used to maximize hydrogenation efficiency. In this manner, the cooling options for warmer cooling media used in the prior art are maintained in a modified pumparound / intercooler that reduces the expensive low level cooling required in overhead systems. Furthermore, the heat removed by these pump-around streams is utilized elsewhere in the ethylene plant to reduce energy consumption. Another advantage of the new pump-around system is that it allows a relatively high liquid flow rate without affecting the separation performance of the entire column due to heavy material in the overhead, as in the prior art. At high liquid flow rates, pump around can provide the required liquid load on the catalyst without the need for further reflux. This allows operation of the catalytic distillation column at a lower reflux rate than previously possible without the disadvantages in distillation efficiency observed in the prior art. A reflux rate in the range of 0.5-1.8 is sufficient to form the required catalyst liquid wetting (prior art required a value of the order of 5). In addition to a clear reduction in energy consumption, the present invention can utilize a higher hydrogen partial pressure due to a lower reflux rate, thereby reducing the required catalyst capacity.

In a catalytic distillation column, it is important that the catalyst is always moistened to ensure that the reaction takes place in the liquid phase. The selectivity of the catalytic distillation system is partly performed in the liquid phase with certain components that are desired to be kept unreacted by the operator (for example, ethylene, which is maintained at the highest concentration in the vapor phase). Depends on the reaction Maintaining some liquid descending in the column is important to keep the catalyst wet. If the liquid feed is greater than 800 lb (liquid) / hour / (ft) ^{2 of} the tower cross section, the catalyst will be highly wetted and the selectivity of the reaction will be maintained.

  The secondary control variable is the change in reflux with a change in reboiler capacity. Thus, both the temperature and composition of the catalyst bed are altered to achieve the desired hydrogenation.

  In addition, a changeable feed position that allows for the feed of the main feed at a point below the stripping section 22 before reaching both the catalyst bed 16 and the side stream 52 for the first pump around. Provide some separation of the various heavy components present in the feed. In this way, the circulation of heavy and potential fouling components on the catalyst bed is eliminated. Furthermore, the feed point above the first catalyst bed allows for a lowering operation, thus avoiding problems with excess catalyst, and thus selectivity loss under these low flow conditions. Merged into The bottom 63 from the tower 14 is sent for further processing if desired.

As shown in FIG. 2, the present invention further comprises a fixed bed trim reactor system that provides further hydrogenation of stream 50. This system typically consists of two reactors with an intercooler, but may consist of a series of reactors in which there is an intercooler between successive reactors. The fixed bed reactor system offers four advantages:
1. It is not necessary to operate the catalytic distillation column at a high hydrogenation level, while maintaining the acetylene C ₂ standard, with high acetylene, methylacetylene, and propadiene conversion rates, maximum productivity for the catalyst, Manipulated for weight gain;
2. Changes in the catalytic distillation overhead concentration of acetylene compounds and dienes resulting from catalyst deactivation, increased carbon monoxide content, or feedstock changes are adjusted;
3. During temporary upset and / or catalyst deactivation or catalyst poisoning, hydrogen removal is sustained, thus stabilizing the performance of the downstream refrigeration system. As the amount of hydrogen from the system changes, the partial pressure of the downstream distillation system changes and the required refrigeration varies, which forms an upset of the process and is highly undesirable;
4). There is an opportunity to regenerate the catalyst by using a spare fixed bed reactor (longer on-stream operating life of the entire system).

  In addition to controlling the temperature and composition profile of the column, it is important to operate below the complete conversion of acetylene and diene throughout the catalytic distillation column. By doing so, an increase in ethylene and propylene is achieved. Furthermore, this operation requires a smaller amount of catalyst than prior art complete hydrogenation, thereby maximizing the productivity of the catalyst for catalytic distillation. This is possible in operation with a fixed bed reactor system following the column.

Tower, ethylene liquid concentration in the reaction bed when operating such that less than about 1%, hydrogenation over 95% of the diene C ₂ -C ₅ and heavier is achieved. As a result, diene and acetylene 5,000-7,500 ppm are present in the vapor stream 50 from the reflux drum 46, with a minimum ethylene loss of 1%. In order to set the conversion of acetylene to 100%, the loss of ethylene is higher. This operation is consistent with a hydrogen removal rate of about 30-35%, depending on the feed composition. However, the total conversion of the diene and acetylenic compounds of C ₂ -C ₅ and heavier is reduced to 80-95%, as a result, the C ₂ -C ₅ and heavier in the outlet stream 50 dienes and acetylenes Is 10,000-20,000 ppm, typically 15,000 ppm, and increased ethylene is achieved. FIG. 3 is a graph of diene (ppm) versus ethylene gain or loss (mass%) at the outlet for a catalytic distillation unit (CDU) alone and a CDU + fixed bed hydrogenation system. As can be seen from FIG. 3, by maintaining a certain amount of highly reactive acetylene and diene in the unreacted state in the overhead from the catalytic distillation column, the loss of ethylene and propylene can be eliminated, while still maintaining 100% conversion of the whole acetylene can be achieved.

The fixed bed reactor arranged behind the catalytic distillation column 14, the C ₂ acetylene is the C ₃ and heavier dienes and total 10,000-55,000 ppm and representative of acetylene with 20,000 ppm, from catalytic distillation column Breakthrough is allowed. A typical system with two fixed bed hydrogenation reactors with intercooling is 100% C ₂ acetylene entering the fixed bed reactor and C ₃ and heavier dienes and acetylenes entering the fixed bed reactor. About 75% of the total was shown to be hydrogenated. This results in breakthroughs of diene and acetylene 2,500-14,000 ppm and typically only 5,000 ppm from the merged system. This indicates that about 97% of the total acetylenes and dienes C ₂ and heavier in the feed is hydrogenated. By such an operation, when the conversion of acetylene is 100%, the selectivity of all acetylene to ethylene is 70%, and the total increase rate of ethylene is less than 0.5%. This is a substantial improvement over the prior art.

The intrinsic hydrogenation rate of ethylene is only slightly lower than that of propadiene. Thereby, cross-observation of C ₃ diene conversion provides a reliable indication of the stability of ethylene gain and is used as a control point for the system. In the catalytic distillation column alone, loss of ethylene is observed when the conversion of C ₃ diene is 40-60%, typically 45%. However, when operating under conditions where the C ₃ diene conversion in the catalytic distillation column is 10-35%, typically 20%, an increase in ethylene of 0.2-0.5% is possible. In the present invention, the conversion of propadiene can be substantially increased while maintaining the increased amount of ethylene.

  Under normal operation of the ethylene unit, a change in the carbon monoxide content of the charge gas 10 is observed. In addition, the feed quality and the severity of operations that affect the acetylene and diolefin content of the charge gas are varied. Due to the constant catalyst capacity in the catalytic distillation column, an increase in the carbon monoxide or diene and acetylene inlet concentration results in a lower conversion, thereby increasing the discharge of these unwanted products into the stream 50. Compensation for such anticipated disruptions would be difficult with the prior art alone shown in FIG. This would require an increase in operating pressure or temperature that affects the overall performance of the fractionation system. In an improved process involving a fixed bed reactor system, the temperature of the vapor 50 entering the fixed bed reactor system increases or decreases the reactivity of the reactor system, thereby following changes in catalytic distillation reaction activity, and Adjusted to maintain complete C2 acetylene removal and high hydrogen removal efficiency.

  Finally, the fixed bed hydrogenation reactor system is designed to include a reserve as well as an operating reactor. Catalyst deactivation occurs in both fixed bed and catalytic distillation systems. It is not possible to regenerate the catalyst for catalytic distillation without interrupting the process or installing parallel columns. Both of these options are expensive. However, a spare fixed bed vapor phase reactor is a relatively inexpensive option. By utilizing a fixed bed reactor system with reserve instead of the prior art single tower concept, the on-stream life of the process is substantially improved.

In a fixed bed hydrogenation system, overhead 50 from the catalytic distillation process passes through a cross-flow heat exchanger 64 and an inlet heater 66 and enters a fixed bed reactor 68. The effluent from the first reactor 68 passes through the intercooler 70 and enters the second fixed bed hydrogenation reactor 72. If required, a series of fixed beds with an intercooler behind can be used in the same manner to achieve the necessary heat transfer. The effluent from the last reactor 72 then passes through a cross flow heat exchanger 64 where heat is extracted and the feed 50 to the fixed bed reactor is heated. The inlet temperature to the fixed bed reactor is rapidly changed to increase or decrease the degree of hydrogenation in the fixed bed reactor. Such control is necessary to effectively cope with changes in carbon monoxide or diene and acetylene feed concentrations. Up to a maximum adiabatic temperature rise of 45 ° C. (80 ° F.) for both beds, stable fixed bed operation is possible without loss of ethylene. Under normal operation, a typical adiabatic rise is expected to be 20 ° C. (35 ° F.). An adiabatic increase of 39-45 ° C (70-80 ° F), typically 45 ° C (80 ° F), treats 35,000-58,000 ppm, typically 43,000 ppm, of acetylene and dienes from catalytic distillation. In the final product stream 74, 9,000 to 30,000 ppm, typically 10,000 ppm of C ₃ and heavier diene and acetylene compounds are produced, while the conversion of C ₂ acetylene primarily to ethylene is Maintained at 100%.

  In a similar manner, temperature control at the inlet to the fixed bed reactor can provide compensation for catalyst deactivation that provides typical initial and end temperature temperatures for the fixed bed system. In the prior art, this is done only by correcting the temperature in the catalytic distillation column. In this case, a pressure change in the column is necessary, which changes the fractionation conditions. When using both a catalytic distillation column and the fixed bed reactor system of the present invention, the catalytic distillation column is operated under constant fractional distillation conditions and the required temperature correction for the fixed bed system is lower. This improves system stability and allows for a longer catalyst life.

  FIG. 4 shows another embodiment of the present invention. In the catalytic distillation tower, the overhead stream 42 passes through the heat exchanger 44 and then enters the reflux drum 46, and in this embodiment, the overhead stream 42 is directly directed to the cross flow heat exchanger 64. And enter the fixed bed reactor system. Following the fixed bed system, the effluent is cooled at 65 and separated as condensate 69 at 67 to return reflux 48 for the column to the column.

  Since the stream entering the fixed bed reactor system still contains all of the reflux for the column, the operating temperature of the fixed bed reactor is somewhat higher to ensure complete vapor flow. This changes the expected catalyst activity and space velocity to ensure stable operation. The advantages of this approach are higher hydrocarbon mass flow (minimizing fixed bed temperature rise), reduced hydrogen partial pressure (improving selectivity) and higher space velocity. (To improve selectivity and reduce catalyst cost).

  FIG. 5 shows another embodiment of the present invention with a pre-reactor. This arrangement is advantageous for bulk selective hydrogenation of feeds with high diene and acetylene contents. Following compression at 12 and possible processing in the guard bed (not shown), the vapor phase feedstock passes through the recycle liquid 76 from the pump 56 of the column 14 and the fixed bed reactor 78 in a cocurrent manner. Mix with the two-phase mixture. Hydrogenation occurs and the presence of liquid acts to control the temperature rise through vaporization. The hydrogenation reactor 78 is designed as an operational reactor plus a reserve to allow expansion of the on-stream operation of the system. Following the pre-reactor, the liquid / vapor mixture is sent as a mixed feed directly to the column or separated in a separation drum, and the liquid and vapor are sent separately to the column. The latter is preferred because the various oligomers formed in the initial hydrogenation are in liquid phase and are fed to the column below the catalyst bed, thereby reducing contamination.

  By performing fixed bed hydrogenation in front of the catalytic distillation column 14, the amount of highly active diene and acetylene compounds utilized in the hydrogenation that are preferentially absorbed is high, so that the former part of the hydrogenation has ethylene. High catalyst utilization is achieved without loss. With high catalyst utilization, the process can be economical with smaller catalyst capacity. A catalytic distillation unit is still required following the pre-reactor to reach the hydrogenation specification. In the pre-reactor, it is expected that up to 50%, typically 20% of the hydrogenation capacity will be achieved.

  Another advantage of the fixed bed hydrogenation reactor located in front of the catalytic distillation column 14 is that the reactor is used as a guard bed for the catalyst poison. The catalyst is nickel or palladium. For example, a nickel catalyst exhibits a catalytic action in a reaction in which a sulfur compound, thiophene, is reacted with butadiene to produce a heavy mercaptan. The mercaptan is then removed in the stripping section 22 of the column 14 and thus has no contact with the palladium catalyst. Yet another advantage is that the external pre-reactor system can have a reserve and thus can be regenerated without having to shut down the entire plant for catalyst replacement.

  Alternatively, as shown in FIG. 6, the liquid 76 from the pump 56 can flow downward through the fixed bed 78 and the vapor stream from the compressor 12 can flow upward. The liquid from the bottom of the fixed bed reactor 78 then flows to the lower part of the column 14 and the vapor flows to the higher inlet. The advantage of this countercurrent process arrangement is that the oligomers produced from the polymerization reaction of unsaturated hydrocarbons are removed from the catalyst bed as it is produced and do not pass through the rest of the catalyst bed. This liquor is also sent to column 26 at the lower inlet, thus minimizing potential contamination of catalyst in column 14.

  Oligomers that contaminate the catalyst for catalytic distillation are easily separated and do not rise in the column and do not contaminate the catalyst. Further, as in the co-current option, the pre-reactor catalyst bed can have a reserve, and thus can be regenerated while the rest of the system remains operational. The ability to easily regenerate online would increase the length of the system cycle as the catalyst zone at the feed inlet is expected to have the highest contamination rate.

  In order to minimize contamination in the fixed bed pre-reactor, the liquid flow rate should be sufficient to minimize local heat points due to the high heat of hydrogenation and to wash away the oligomers produced from the catalyst. is necessary. The operation of these beds is preferably a steam continuous zone. For cracked gases that exhibit extreme contamination tendencies, operation in the liquid continuous zone is also possible.

  FIG. 7 shows yet another embodiment of the present invention with a fixed bed reactor in the liquid cooler stream or intercooler stream withdrawn from the column 14. Fixed bed hydrogenation reactors 82 and 84 are arranged in a side stream from collection tray 30 and a side stream from collection tray 31, respectively. These fixed beds 82 and 84 are present in the hydrogenation sections 16 and 18 in the catalytic distillation column 14 in addition to the reactive hydrogenation sections 16 and 18. A mass transfer zone 85 in the form of a structured packing or tray is also added above the removal site and below the catalyst bed. This zone allows the hydrogen to be saturated in the liquid phase, thereby providing the hydrogen necessary for the hydrogenation of acetylene and dienes in the withdrawal liquid.

In the present invention, 30 to 40% of hydrogen can be removed from the charge gas before the chilling and condensing step, thereby reducing the energy consumption and reducing the equipment cost. It is not possible in the prior art to hydrogenate 100% acetylene regardless of carbon monoxide concentration without causing loss of C ₂ or C ₃ olefins. The combined fixed bed and catalytic distillation process can provide an excellent response to system upsets while maintaining stable diene / acetylene hydrogenation and hydrogen removal.

  In the following, several examples will be described, but these will explain specific examples of the present invention in comparison with the prior art. Table 1 below shows the composition of the feed used in all examples. Table 2 shows the results in each example.

Example 1

This example shows the prior art based on a one-step catalytic distillation column as outlined in US Pat. No. 5,679,241 (FIG. 1) operating at a reflux rate of 4.4. When using a typical front-end acetylene hydrogenation catalyst with a palladium level of 2,000 ppm or less and operating at a pressure of 195 psig and an average catalyst temperature of 110 ° C (230 ° F), the conversion of C ₂ acetylene is the loss of ethylene / Reach 84% with 0% acquisition. At the reactor outlet, C ₂ acetylene is 370 ppm and the total of diene and acetylene compounds is 19,070 ppm. The total acetylene / diene conversion is 79.5%. This example shows the case where a single column is operated for no loss of ethylene. As will be appreciated, there is a substantial breakthrough of C ₂ acetylene. This will produce a defective product.

Example 2

This example is also related to the prior art and is based on a single catalyst distillation column operated at a higher catalyst temperature and a slightly lower reflux rate of 4.1. The severity of hydrogenation in a single column is increased as low C ₂ acetylene levels are reached. This is accomplished by increasing the temperature or increasing the activity of the catalyst. The higher temperature operation is to reduce the acetylene content, thereby achieving a standard ethylene product. In this case, the conversions of diene and acetylene are both higher than in Example 1, but there is a loss of ethylene of 0.6%. At the reactor outlet, C ₂ acetylene is 240 ppm and the diene and acetylene compounds total 12,340 ppm. The total acetylene / diene conversion is 86.7%. As will be appreciated, an increase in the conversion of acetylene is accompanied by an increase in acetylene loss (which is economically undesirable). Furthermore, the ethylene product still does not meet the specification limits of 1-2 ppm.

Example 3

This example is related to the prior art and is based on the single catalytic distillation column of Example 1 using a higher carbon monoxide level feed. Carbon monoxide acts as a catalyst poison, thus substantially reducing the conversion of diene and acetylene compounds. If the carbon monoxide level in the feed is 0.1 mol% (carbon monoxide / hydrogen ratio 0.6%), the product has 460 ppm of acetylene and a total of 33,860 ppm of diene and acetylene compounds. The decrease in catalytic activity due to CO is reflected as a decrease in catalyst productivity (0.12-0.09 lbmol / hr ft ³ ) and a decrease in total acetylene / diene conversion (63.6% vs. 79.5% (for the base case)).

  In response to this reduction in activity, the temperature of the catalyst in the catalytic distillation column increases. For this reason, it is necessary to increase the pressure exceeding that practical in the operation unit. Options for compensating for such an increase in CO are limited to the prior art.

Example 4

This example represents an improved combined operation of the catalytic distillation column and fixed bed reactor shown in FIG. The combination operation, 100% conversion of the C ₂ acetylene, is necessary in order to achieve an increase in ethylene in conversion of 50-95% of all other dienes and acetylenic compounds. In the operation of a catalytic distillation column at a pressure of 195 psig and an average catalyst temperature of 110 ° C. (230 ° F.), diene and acetylene are 12,000 ppm (mass) in the catalytic distillation overhead, which is then fixed bed hydrogenation reactor. Supplied to the system.

By operating a fixed bed hydrogenation reactor at a gas hourly space velocity (GHSV) of 1,800 hours- ¹ and a bed inlet temperature of 53.9 ° C (230 ° F), 100% C ₂ acetylene can be converted, and moreover Hydrogenation can be provided, which achieves a total conversion of C ₃ diene (both catalyst and fixed bed hydrogenation) of 50%, with a total conversion of 96.1% of diene and acetylene compounds in the feed to the combined system. It was. This resulted in 0 ppm C ₂ acetylene at the outlet and 3640 ppm C ₃ and heavier diene and acetylene compounds. Standardized ethylene products are produced with very high total conversion for highly unsaturated species.

Example 5

This example represents an improved combined operation of a catalytic distillation column and a fixed bed reactor using a high carbon monoxide level feed. Under certain operating conditions with a carbon monoxide level of 0.1 mol% in the feed and a catalytic distillation column, it was only necessary to adjust the fixed bed inlet temperature in order to maintain product specifications. Specifically, when carbon monoxide is increased from 0.05 to 0.1%, the fixed bed inlet temperature is increased from 53.9 ° C. (129 ° F.) to 60 ° C. (140 ° F.) to maintain C ₂ acetylene conversion. ) Was sufficient. In addition, C ₃ and heavier dienes and acetylene compounds were further hydrogenated, so the total of diene and acetylene in the product was 7,740 ppm (mass).

Raise the improvements of the present invention, without hydrogenating the C ₂ and C ₃ olefins, stable, hydrogenated 90% of C ₃ -C ₅ and heavier acetylenic compounds in the feed stream +, C ₄ and C ₅ dienes hydrogenated 90% +, and C ₃ in diene conversion 50% +, can achieve 100% hydrogen of C ₂ acetylene. The hydrogen removal achieved by the present invention is stable at 30-40%, typically 30%, depending on the feed composition.

1 is a prior art flow diagram involving catalytic distillation alone with bottom recycle for temperature control. It is a flow diagram explaining the present invention. Figure 3 is a graph showing ethylene gain or loss versus diene exit level for the present invention compared to the prior art. FIG. 4 is a flow diagram similar to FIG. 2 but illustrating another example of the present invention. It is a flow diagram explaining the other specific example of this invention. 6 is a flow diagram of another example of the method of FIG. FIG. 4 is a flow diagram similar to FIG. 2 but illustrating another example of the present invention.

Claims (29)

Hydrogen, ethylene, propylene, acetylene, processes the methyl acetylene, propadiene and other C _4, C _5, C ₆ and more pyrolysis feed stream containing unsaturated hydrocarbons heavier hydrogen ethylene and propylene Without conversion, essentially all of the acetylene is hydrogenated and converted at a high ratio to ethylene, and the methylacetylene, propadiene and other C ₄ , C ₅ , C ₆ and heavier non-volatiles. A process for treating a pyrolysis feed stream that hydrogenates at least a portion of a saturated hydrocarbon to an olefin, thereby consuming a portion of the hydrogen, comprising:
a. Introducing the feed stream into a catalytic distillation column containing at least one hydrogenation catalyst, and simultaneously (i) selectively hydrogenating a portion of the acetylene to produce ethylene, and the methylacetylene, propadiene And hydrogenating C ₄ , C ₅ , C ₆ and heavier unsaturated hydrocarbons, and controlling the hydrogenation conditions, thereby not hydrogenating the ethylene and propylene; and (ii) the feed Separating the product stream into light and heavy hydrocarbons by fractional distillation;
b. Removing substantially all of the hydrogen and the remainder of the light hydrocarbons from the catalytic distillation column as vapor phase overhead and substantially all of the heavy hydrocarbons as bottoms;
c. At least a portion of the vapor phase overhead is introduced into a vapor phase fixed bed reactor system containing a hydrogenation catalyst, and the remainder of the acetylene is hydrogenated to produce additional ethylene, and the methyl acetylene, propadiene and Hydrogenating additional portions of C ₄ , C ₅ , C ₆ and heavier unsaturated hydrocarbons and controlling the hydrogenation conditions to not hydrogenate said ethylene and propylene; and d. A process for treating a pyrolysis feed stream comprising the step of removing the mixed product from the fixed bed reactor system. Hydrocarbons lighter consists hydrocarbons C ₄ and lighter, heavier hydrocarbons are those consisting of hydrocarbons C ₅ and heavier, the process of claim 1. Hydrocarbons lighter consists hydrocarbons C ₅ and lighter, heavier hydrocarbons are those consisting of hydrocarbons C ₅ and heavier, the process of claim 1.   The method of claim 1 further comprising condensing a portion of the vapor phase overhead and returning the condensed portion to the catalytic distillation column as reflux.   The process of claim 1, wherein essentially all of the vapor phase overhead is introduced into the vapor phase fixed bed reactor system and a portion of the product is returned to the catalytic distillation column from the fixed bed reactor system as reflux. .   The method of claim 1, wherein controlling the conditions for selective hydrogenation in the catalytic distillation column and the vapor phase fixed bed reactor comprises controlling the temperature profile in the catalytic distillation column and the vapor phase fixed bed reactor. .   The step of controlling the temperature profile in the catalytic distillation column takes out a part of the liquid descending in the catalytic distillation column as a side flow at a selected site, cools this side flow, 7. The method of claim 6, comprising returning to the catalytic distillation column at a selected site or above it.   The method of claim 7, further comprising hydrogenating the side stream.   The step of controlling the temperature profile in the catalytic distillation column is to take a pump around stream from a site below the hydrogenation catalyst bed of the catalytic distillation column, cool the pump around stream, and cool the cooled pump around stream to the hydrogen 7. A process according to claim 6, comprising the step of returning to the catalytic distillation column at a site above the catalytic catalyst bed.   The vapor phase fixed bed reactor system comprises at least one reactor, and the step of controlling the temperature in the vapor phase fixed bed reactor system controls the temperature in the heat exchanger before the reactor. The method according to claim 1, comprising the step of:   The vapor phase fixed bed reactor system comprises two or more reactors in series, and the step of controlling the temperature in the vapor phase fixed bed reactor system is the heat before each of the reactors. The method of claim 1 including the step of controlling the temperature in the exchanger.   The method according to claim 1, wherein the step of selectively hydrogenating comprises operating the catalytic distillation column such that the concentration of ethylene in the liquid phase in the catalyst bed is less than 2% by mass. The step of selectively hydrogenating is a step of operating the catalytic distillation column so that the liquid flowing downward in the catalytic distillation column is at least 800 lb / hr / ft ² (cross-sectional area) in the catalyst bed zone. The method of claim 1, comprising:   The step of introducing a feed stream into the catalytic distillation column comprises mixing the feed stream with the recycle liquid from the catalytic distilling column, and the mixed feed stream and recycle liquid before the catalytic distilling column. Introducing into a fixed bed hydrogenation pre-reactor, thereby hydrogenating a portion of the highly unsaturated hydrocarbons and introducing a vapor and liquid stream into the catalytic distillation column, Item 2. The method according to Item 1.   The process according to claim 1, wherein the hydrogenation catalyst bed in the catalytic distillation column contains a catalyst comprising a Group VIIIA metal supported on a support.   The method according to claim 12, wherein the catalyst contains palladium supported on alumina.   The process according to claim 13, wherein the catalyst contains an additive selected from the group consisting of gold, silver and alkali metals.   The process of claim 1 wherein the catalysts having different amounts of palladium are placed in selected portions of the catalytic distillation column.   The process according to claim 15, wherein the different catalysts are arranged in different parts of the catalytic distillation column.   20. A process according to claim 19, wherein the different catalysts contain different metals.   20. A process according to claim 19, wherein the different catalysts have different metal contents.   The process according to claim 15, wherein the catalyst comprises nickel supported on a support.   The process according to claim 15, wherein the catalysts in different parts of the catalytic distillation column contain a combination of palladium supported on the support and nickel supported on the support.   The process according to claim 1, wherein the hydrogenation catalyst in the fixed bed reactor system comprises a Group VIIIA metal supported on a support.   25. The method of claim 24, wherein the hydrogenation catalyst in the fixed bed reactor system contains palladium supported on alumina.   25. The process of claim 24, wherein the hydrogenation catalyst in the fixed bed reactor system contains palladium supported on alumina with a co-catalyst comprising gold, silver, alkali metal or combinations thereof.   15. A process according to claim 14, wherein the fixed bed hydrogenation pre-reactor contains a nickel catalyst and the pre-reactor is responsible for reacting sulfur compounds for removal.   The method of claim 1 further comprising removing catalyst poisons from the feed stream prior to introduction into the catalytic distillation column.   29. The method of claim 28, wherein the catalyst poison is lead, arsenic and mercury.

Images