Patent Application
20190127303
The present invention relates to the preparation of 2,3,3,3-tetrafluoropropene (HFO-1234yf). More particularly, the present invention relates to the fluorination of very high purity 1,1,1,2,3-pentachloropropane (HCC-240db) into very high purity 2,3,3,3-tetrafluoropropene (HFO-1234yf).
The protocol of Montreal for the protection of the ozone layer led to the end of the use of chlorofluorocarbons (CFCs). Less aggressive compounds for the ozone layer, such as the hydrofluorocarbons (HFCs) e.g. HFC-134a replaced chlorofluorocarbons. These latter compounds were indeed shown to provide greenhouse gases. There exists a need for the development of technologies, which present a low ODP (ozone depletion potential) and a low GWP (global warming potential). Although the hydrofluorocarbons (HFCs), which are compounds which do not affect the ozone layer, were identified as interesting candidates, they exhibit a relatively high GWP value. There still exists the need to find compounds which exhibit a low GWP value. Hydrofluoroolefins (HFO) were identified as being possible alternatives with very low ODP and GWP values.
New classes of environmentally friendly halocarbons are emerging and have been investigated, and in some cases, embraced in a number of applications, especially as refrigerants in the automotive and domestic fields. Examples of such compounds include 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf), 1,3,3,3-tetrafluoropropene (HFO-1234ze), 3,3,3-trifluoropropene (HFO-1243zf), and 2,3,3,3-tetrafluoropropene (HFO-1234yf), 1,2,3,3,3-pentafluoropropene (HFO-1225ye), 1-chloro-3,3,3-trifluoropropene (HCFO-1233zd), 3,3,4,4,4-pentafluorobutene (HFO-1345zf), 1,1,1,4,4,4-hexafluorobutene (HFO-1336mzz), 3,3,4,4,5,5,5-heptafluoropentene (HFO-1447fz), 2,4,4,4-tetrafluorobut-1-ene (HFO-1354mfy) and 1,1,1,4,4,5,5,5-octafluoropentene (HFO-1438mzz).
While these compounds are, relatively speaking, chemically non-complex, their synthesis on an industrial scale to the required levels of purity is challenging. Many synthetic routes proposed for such compounds increasingly use, as starting materials or intermediates, chlorinated alkanes or alkenes. Examples of such processes are disclosed in WO2012/098420, WO2013/015068 and US2014/171698. The conversion of the chlorinated alkane or alkene starting materials to the fluorinated target compounds is usually achieved using hydrogen fluoride and optionally transition metal catalysts, for example chromium-based catalysts.
An example of an optionally non-catalytic process for preparing fluoroalkenes is disclosed in WO2013/074324.
The issue of the formation of impurities during hydrofluorination reactions is considered in US2010/331583 and WO2013/119919, where thus the need for part fluorinated feedstock purity is described, and also in US2014/235903 regarding reactor impurities.
It has been recognized that when the chlorinated feedstock is obtained from a multi-step process, especially if such steps are linked and run continuously to achieve industrially acceptable product volumes, then the need to prevent cumulative side reactions from generating unacceptable impurities at each process step is very important.
The purity of the chlorinated starting materials will have a substantial effect on the success and viability of the processes (especially continuous processes) for preparing the desirable fluorinated products. The presence of certain impurities will result in side reactions, minimizing the yield of the target compound. Removal of these impurities through the use of distillation steps is also challenging. Additionally, the presence of certain impurities will compromise catalyst life, by, for example, acting as catalyst poisons.
Using high-purity chlorinated alkanes therefore is the crucial factor for the synthesis of the fluorinated compounds mentioned above.
Several processes for producing purified chlorinated compounds have been proposed in the art. For example, WO2013/086262 discloses a process for preparing 1,1,2,2,3-pentachloropropane from methylacetylene gas. As can be seen from the examples in that application, the bench scale syntheses disclosed therein resulted in a product having around 98.5% purity, despite being subjected to post-synthetic purification process steps, specifically distillation.
In WO2014/130445, a conventional process is discussed on page 2 of that publication, the first step of which involves the formation of 1,1,1,2,3-pentachloropropane from 1,1,3-trichloropropene. However, the purity profile of that intermediate product is not outlined, nor is any importance attached to the purity profile of that product. In Example 2 of WO2014/130445, a HCC-240db (1,1,1,2,3-pentachloropropane) rich material having a purity level of 96.5 to 98.5% is used.
WO2013/055894 discloses a process for producing tetrachloropropenes, particularly 1,1,2,3-tetrachloropropene and reports that the product obtained from the processes disclosed in that document have advantageously low levels of impurities that can be problematic in downstream processes for producing fluorocarbons. A discussion of the different types of impurities considered to be problematic by the authors of WO2013/055894 is set out in paragraphs [0016] and [0017] of that document.
US2012/157723 discloses a process for preparing chlorinated alkanes via a three step process. Seemingly high purity chloroalkanes appear to have been prepared according to the process disclosed in that document. However, the purity data presented in the examples of that application are only given to one decimal place.
From the provision of data presented in this way, it is apparent that the analytical equipment used to measure the impurity profile of the products obtained in the examples of US2012/157723 was insensitive; conventional analytical apparatus enables hydrocarbon levels to 1 ppm (i.e. to four decimal places). Given that one skilled in the art would need to know the impurity profile of chloroalkane feedstock to be used in industrial scale down to a ppm level, the data presented in US2012/157723 would not be of assistance.
The skilled person would also recognize that the process disclosed in US2012/157723 provides 1,1,1,2,3-pentachloropropane, which has relatively low selectivity; as can be seen, from paragraph [0146] of that document, selectivity towards the compound of interest was 95%.
Additional processes in which processes are streamlined by using crude intermediates in downstream stages are disclosed in WO2009/085862.
Despite these advances, problems can still arise through the use of chlorinated compounds obtained from the processes discussed above. Particularly, the presence of impurities especially those which are not easily separable from the compounds of interest (e.g. as a result of similar or higher boiling points) or which reduce the effectiveness or operating life of catalysts used in downstream processes can be problematic.
To minimize such drawbacks, a demand remains for very high purity chlorinated alkane compounds, and also for efficient, selective and reliable processes for preparing such compounds, especially enabling continuous industrial manufacture. Processes achieving these objects are described in this document further below.
Several processes for producing at least one of 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf) by fluorination of pentachloropropane have been proposed in the art.
WO 2012/052797 discloses for instance a two-steps process wherein pentachloropropane including 1,1,1,2,3-pentachloropropane (HCC-240db) and/or 1,1,2,2,3-pentachloropropane (HCC-240aa), is in a first step contacted with hydrogen fluoride (HF) to provide 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf), which is then converted into HFO-1234yf in a second step.
WO 2012/052798, for example, describes a process of catalytic fluorination in gas phase of product 1,1,1,2,3-pentachloropropane (HCC-240db) and/or 1,1,2,2,3-pentachloropropane (HCC-240aa) into product 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) in presence of a catalyst and oxygen.
WO 2012/098420 describes for instance a process of catalytic fluorination in gas phase of product 1,1,1,2,3-pentachloropropane (HCC-240db) and/or 1,1,2,2,3-pentachloropropane (HCC-240aa) into product 2,3,3,3-tetrafluoropropene (HFO-1234yf) in presence of a catalyst.
WO 2013/088195, for example, discloses a process for preparing 2,3,3,3-tetrafluoropropene (HFO-1234yf) from 1,1,1,2,3-pentachloropropane (HCC-240db) and/or 1,1,2,2,3-pentachloropropane (HCC-240aa), comprising the following steps: (a) catalytic reaction of 1,1,1,2,3-pentachloropropane (HCC-240db) and/or 1,1,2,2,3-pentachloropropane (HCC-240aa) with HF into a reaction mixture comprising HCl, 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf), 2,3,3,3-tetrafluoropropene (HFO-1234yf), unreacted HF, and optionally 1,1,1,2,2-pentafluoropropane (HFC-245cb); (b) separating the reaction mixture into a first stream comprising HCl and 2,3,3,3-tetrafluoropropene (HFO-1234yf) and a second stream comprising HF, 2-chloro-3,3,3-trifluoropropene and optionally 1,1,1,2,2-pentafluoropropane (HFC-245cb); (c) catalytic reaction of the second stream into a reaction mixture comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf), HCl, unreacted 2-chloro-3,3,3-trifluoropropene (HFO-1233xf), unreacted HF, and optionally 1,1,1,2,2-pentafluoropropane (HFC-245b), and (d) feeding the reaction mixture of step (c) directly without separation to step (a).
WO 2015/055927 discloses for instance a method for producing a fluorinated compound, comprising the provision of a gaseous flow comprising hydrofluoric acid; the provision of at least one liquid flow of a chlorinated compound and the vaporization thereof by mixing with said gaseous flow, the resulting mixture being a gaseous mixture; and the catalytic reaction of the chlorinated compound with hydrofluoric acid in a gaseous phase and the collection of a product flow.
In WO 2011/077192, a process for preparing 2,3,3,3-tetrafluoropropene (HFO-1234yf) is disclosed, with the process comprising (i) contacting 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) with hydrogen fluoride HF in gas phase in the presence of a fluorination catalyst under conditions sufficient to produce a reaction mixture; (ii) separating the reaction mixture into a first stream comprising HCl, 2,3,3,3-tetrafluoropropene (HFO-1234yf) and a second stream comprising HF, unreacted 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 1,1,1,2,2-pentafluorpropane (HFC-245cb); (iii) recycling at least a part of the second stream at least in part back to step (i). The document further describes a process for preparing 2,3,3,3-tetrafluoropropene (HFO-1234yf), comprising: (i) contacting 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) with hydrogen fluoride HF in gas phase in the presence of a fluorination catalyst under conditions sufficient to produce a reaction mixture; (ii) separating the reaction mixture into HCl and a stream containing the fluorinated products; (iii) separating said stream containing the fluorinated products into a first stream comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf) and a second stream comprising HF, unreacted 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 1,1,1,2,2-pentafluorpropane (HFC-245cb); (iv) recycling at least a part of the second stream at least in part back to step (i).
WO 2013/045791 discloses a method for the preparation of 2,3,3,3-tetrafluoropropene (HFO-1234yf) from halopropanes having the formula CX3CHClCH2X and halopropenes having the formulas CX3CCl═CH2/CClX2CCl═CH2 and CX2═CClCH2X, where X is independently a fluorine or chlorine atom. The method includes in particular at least one step during which 2-chloro-3,3,3-trifluoro-1-propene, optionally mixed with at least one halopropane having the formula CX3CHClCH2X and/or at least one halopropene having the formulas CClX2CCl═CH2 and CX2═CClCH2X, where X is independently a fluorine or chlorine atom, reacts with HF in the gaseous phase in the presence of a fluoridation catalyst at a temperature of between 320 and 420° C. with a molar ratio of oxygen to 2-chloro-3,3,3-trifluoro-1-propene of more than 1 but not more than 2.5, and a molar ratio of HF to the total amount of organic compounds to be reacted of between 5 and 40.
All the processes described above are sensitive to impurities contained in their starting materials and/or produce byproducts reducing the yield and/or the quality of the final product, i.e. 2,3,3,3-tetrafluoropropene (HFO-1234yf), particularly when they cannot be separated easily.
In consideration of the above it is therefore desirable to provide a process for preparing very high purity 2,3,3,3-tetrafluoropropene (HFO-1234yf) that is efficient, and/or selective, and/or reliable and/or that is practical for a continuous industrial production.
The invention is based on the finding that it is possible to prepare the compound HFO-1234yf of high purity starting from a very high purity initial product, i.e. very high purity 1,1,1,3-tetrachloropropane, or very high purity 1,1,3-trichloropropene, or very high purity 1,1,1,2,3-pentachloropropane.
The invention provides a process for preparing 2,3,3,3-tetrafluoropropene, comprising the following steps:
According to an embodiment, the process of the present invention may comprise steps 4-b) to 4-g) carried out subsequently to step 4-a):
According to an embodiment, the process of the present invention may optionally comprise a step whereby HF from the reaction mixture obtained in step 4-a) is separated, preferably by distillation or decantation, before step 4-b).
According to an embodiment, the process of the present invention may optionally comprise a step whereby HF, if present in the reaction mixture obtained in step 4-c), is separated, preferably by distillation or decantation, before step 4-d).
According to an embodiment the 1,1,3-trichloropropene feedstock is prepared by dehydrochlorination of 1,1,1,3-tetrachloropropane, this method step comprising:
According to an embodiment the 1,1,1,3-tetrachloropropane feedstock is prepared by telomerisation, this method step comprising:
Preferred embodiments have the process comprise all three steps (step 1-3) as explained above, namely step 1), a telomerisation reaction in which carbon tetrachloride is reacted with ethylene to produce 1,1,1,3-tetrachloropropane; step 2), a dehydrochlorination reaction, in which 1,1,1,3-tetrachloropropane is converted to 1,1,3-trichloropropene; and step 3) in which 1,1,3-trichloropropene is chlorinated to yield 1,1,1,2,3-pentachloropropane.
In the process, the reaction mixture obtained from each of the three steps 1 to 3 outlined above is controlled by degree of conversion and subjected to various treatment steps that will be discussed below in more detail. Global impurity profiles in the intermediates and final product are thus managed to produce high grade product 1,1,1,2,3-pentachloropropane. In embodiments of the production method, treatment steps 1-b), 2-b) and/or 3-b) may comprise one or more distillation steps. Additionally or alternatively, treatment steps 1-b), 2-b) and/or 3-b) may comprise contacting compositions comprising 1,1,1,3-tetrachloropropane (in the case of step 1-b), 1,1,3-trichloropropene (in the case of step 2-b), and/or 1,1,1,2,3-pentachloropropane (in the case of step 3-b) with an aqueous medium.
Embodiments have the feedstock produced in step 1-b) and used in step 2-a) comprising:
Embodiments have the feedstock produced in step 2-b) and used in step 3-a) comprising:
Embodiments have, in step 1-a), the concentration of the 1,1,1,3-tetrachloropropane in the reaction mixture in the principal alkylation zone maintained at a level such that the molar ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride in the reaction mixture does not exceed:
Embodiments have the reaction mixture produced in step 1-a) extracted from the principal alkylation zone and subjected to an aqueous treatment step in step 1-b), in which the reaction mixture is contacted with an aqueous medium in an aqueous treatment zone, a biphasic mixture is formed and an organic phase comprising catalyst is extracted from the biphasic mixture.
Embodiments have the catalyst used in step 1-a) being a metallic catalyst, optionally further comprising an organic ligand. The organic ligand may be an alkylphosphate, for example triethylphosphate and/or tributylphosphate.
Embodiments have the reaction mixture produced in step 1-a) being extracted from the primary alkylation zone and fed into the principal alkylation zone, wherein the ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride present in the reaction mixture extracted from the principal alkylation zone is greater than the ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride present in the reaction mixture taken from the primary alkylation zone.
Embodiments have the amount of unreacted ethylene in the reaction mixture leaving the principal alkylation zone being less than 0.6%, less than 0.3%, less than 0.2%, or less than 0.1%.
Embodiments have any unreacted gaseous ethylene being directly recycled back to the reaction zone's operating at elevated pressure, or recycled back to the reaction zone's operating at elevated pressure by absorbing ethylene into the cold liquid carbon tetrachloride feedstock.
According to embodiments, the concentration of the 1,1,3-trichloropropene in the reaction mixture produced in step 2-a) in the dehydrochlorination zone is controlled such that the molar ratio of 1,1,3-trichloropropene:1,1,1,3-tetrachloropropane is from 1:99 to 50:50.
Embodiments have step 2-b) comprising contacting a mixture comprising 1,1,3-trichloropropene, catalyst and 1,1,1,3-tetrachloropropane with an aqueous medium in an aqueous treatment zone, wherein a biphasic mixture may be formed in the aqueous treatment zone and, additionally or alternatively, an organic phase comprising 1,1,1,3-tetrachloropropane and 1,1,3-trichloropropene may extracted from the biphasic mixture.
Embodiments have the dehydrochlorination zone which comes into contact with the reaction mixture in step 2-a) have an iron content of about 20% or less, about 10% or less or about 5% or less, and/or are formed from non-metallic materials, for example enamel, glass, impregnated graphite (e.g. impregnated with phenolic resin), silicium carbide and/or plastics materials such as polytetrafluoroethylene, perfluoroalkoxy and/or polyvinylidene fluoride.
Embodiments have at least some parts of the dehydrochlorination zone which come into contact with the reaction mixture in step 2-a) formed of a nickel-based allow, such as Hastelloy.
In embodiments, in step 3-a), the molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene in the reaction mixture produced in step 3-a) does not exceed 95:5.
In embodiments, the reaction mixture produced in step 3-a) is extracted from the primary reaction zone and is then subjected to a principal conversion step in a principal reaction zone to produce a 1,1,1,2,3-pentachloropropane rich product, which is extracted from the principal reaction zone.
Embodiments have, in step 3-a), the principal conversion step comprise a reduced temperature conversion step in which the reaction mixture extracted from the primary reaction zone is fed into a principal reaction zone operated at a reduced temperature with the 1,1,1,2,3-pentachloropropane rich product being extracted from the principal reaction zone.
According to embodiments, the primary and/or the principal reaction zone is exposed to visible light and/or ultraviolet light.
Embodiments have the reaction mixture/1,1,1,2,3-pentachloropropane rich product produced in step 3-a) subjected to an aqueous treatment and/or hydrolysis step, wherein the aqueous treatment and/or hydrolysis step may comprise contacting the reaction mixture/1,1,1,2,3-pentachloropropane rich product with an aqueous medium in an aqueous treatment zone.
In embodiments, step 3-b) comprises one or more distillation steps, carried out on the reaction mixture and/or the chlorinated alkane rich product produced in step 3-a) and/or the organic phase extracted from the mixture formed in the aqueous treatment zone, as defined above.
According to embodiments HCl is separated from the reaction mixture obtained in step 4-a).
In embodiments, the reaction mixture obtained in 4-a) or the second stream obtained in 4-b) is reacted to obtain 2,3,3,3-tetrafluoropropene (HFO-1234yf).
Embodiments have at least a part of the reaction mixture obtained in step 4-a) recycled to step 4-a) or at least a part of the second stream 4-d) recycled to either step 4-a) or 4-c).
In embodiments, the second stream of step 4-b) or the second stream obtained in step 4-d) is cooled, optionally in the presence of an added amount of at least one compound (C1) chosen from chlorocarbons, hydrochlorocarbons, hydrochlorofluorocarbons, optionally fluorinated alcohols, optionally fluorinated ethers, ketones, esters, polyols and hydrofluorinated ethers in order to give an upper phase rich in HF and a lower organic phase rich in HFO-1234yf and optionally compound C1.
In embodiments, the fluorination reaction of step 4-a) is a gas phase fluorination reaction and/or a liquid phase fluorination reaction.
Embodiments have the fluorination reaction be a gas phase fluorination reaction.
In embodiments, step 4-a) is carried out with a molar ratio HF:HCC-240db from 3:1 to 150:1, preferably 4:1 to 125:1, more preferably 5:1 to 100:1.
In embodiments, step 4-a) is carried out at a pressure from atmospheric pressure to 20 bars, preferably 2 to 18 bars, more preferably 3 to 15 bars.
In embodiments, step 4-a) is carried out at a temperature of from 200 to 450° C., preferably from 250 to 400° C., more preferably from 280 to 380° C.
In embodiments, step 4-a) is carried out with a contact time from 3 to 100 sec, preferably from 4 to 75 sec, more preferably from 5 to 50 sec.
In embodiments, step 4-a) is carried out in the presence of O2 and/or Cl2.
In embodiments, the ratio of O2 and/or Cl2 with respect to HCC-240db is 0.005 to 15 mole %, preferably 0.5 to 10 mole %
In embodiments, step 4-a) and/or step 4-c) is carried out in the presence of a catalyst.
In embodiments, step 4-a) and/or step 4-c) is carried out in the presence of a catalyst which is a chromium catalyst, supported or unsupported, preferably unsupported.
In embodiments using a catalyst, said catalyst may be supported on a support selected from fluorinated alumina, fluorinated chromia, fluorinated activated carbon or graphite carbon.
In embodiments using a chromium catalyst, said catalyst further comprises a co-catalyst selected from Ni, Co, Zn, Mn or mixtures thereof, preferably zinc, and wherein said co-catalyst is preferably present in an amount from about 1-10 wt % of said fluorination catalyst.
In embodiments, step 4-a) and/or step 4-c) is carried out in the presence of a catalyst comprising Ni—Cr, preferably supported.
Embodiments have a stream comprising the 2,3,3,3-tetrafluoropropene stream undergo one or more further purification steps.
In embodiments, the 2,3,3,3-tetrafluoropropene stream is contacted with water and submitted to a drying step.
In embodiments, the 2,3,3,3-tetrafluoropropene stream to be purificated is fed to a distillation column to remove one or more light organic impurities.
In embodiments, the light organic impurity includes an organic compound having a boiling point of from −84° C. to −35° C.
In embodiments, the light organic impurity includes at least one selected from the group consisting of trifluoromethane (HFC-23), 1,1,1,2,2,2-hexafluoroethane (CFC-116), difluoromethane (HFC-32), 1,1,1,2,2-pentafluoroethane (HFC-125), 3,3,3-trifluoropropyne, 1,1,1-trifluoroethane (HFC-143a), 1-chloro-1,1,2,2,2-pentafluoroethane (CFC-115). In embodiments, the 2,3,3,3-tetrafluoropropene stream is subjected to extractive distillation.
Embodiments have the 2,3,3,3-tetrafluoropropene stream contacted with an adsorbent.
In embodiments using an adsorbent, said adsorbent is a molecular sieve, particularly of the X or A type. The molecular sieve may have a mean pore diameter of between 5 and 11 Å.
The invention further relates to a process for manufacturing at least one compound chosen from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf), which process uses as starting material a 1,1,1,2,3-pentachloropropane (HCC-240db) feedstock as prepared according a process comprising the following steps:
According to embodiments said process described above comprises all three production steps 1 to 3.
The invention further relates to a process for manufacturing at least one compound chosen from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf), using as starting material a composition comprising:
The invention further relates to a process for manufacturing at least one compound chosen from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf), using as starting material a composition comprising:
The invention further relates to a process for manufacturing at least one compound chosen from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf), using as starting material a composition comprising:
The invention further relates to a composition obtainable from a process comprising process step 3, and optionally process step 2, and optionally process step 1 as defined above as a feedstock for use in the synthesis of at least one compound chosen from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf), the synthesis comprising at least one fluorination step.
The invention further relates to the use of a composition comprising:
The invention allows the production of high purity 2,3,3,3-tetrafluoropropene (HFO-1234yf) and/or 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and/or 1,1,1,2,2-pentafluoropropane (HFC-245cb) based on a process for producing high purity 1,1,1,2,3-pentachloropropane (HCC-240db), which is preferably produced according a process comprising step 2 defined above using high purity 1,1,3-trichloropropene, the 1,1,3-trichloropropene being preferably produced according a process comprising step 1 defined above using high purity 1,1,1,3-tetrachloropropane.
Further features of the invention will be apparent to those skilled in the art when studying the following description of exemplary embodiments together with the claims and the enclosed figures. In the Figures, functionally and/or structurally equivalent components are, as far as possible, assigned identical or similar reference signs and numerals. It is noted that the invention is defined by the scope of the claims enclosed and is not limited to the configurations of the exemplary embodiments described herein. Other embodiments of the present invention may implement individual features in different combinations than the examples described below. In the following description of exemplary embodiments, reference is made to the enclosed Figures, of which
The invention is based on the finding that very high purity 2,3,3,3-tetrafluoropropene (HFO-1234yf) can be prepared by using very high purity 1,1,1,2,3-pentachloropropane (HCC-240db).
As is apparent, the process for the production of very high purity 2,3,3,3-tetrafluoropropene (HFO-1234yf) can be divided into two main Stages: STAGE 1: The production of very high purity 1,1,1,2,3-pentachloropropane (HCC-240db) (steps 1-3), and STAGE 2: the production of very high purity 2,3,3,3-tetrafluoropropene (HFO-1234yf) (step 4) using the very high purity 1,1,1,2,3-pentachloropropane (HCC-240db) of STAGE 1.
These and other process steps will now be discussed in more detail in the context of each of steps 1) to 4) specifically:
Stage 1: The Production of Very High Purity 1,1,1,2,3-pentachloropropane (HCC-240db) (steps 1-3)
Even though the preferred method for producing very high purity 1,1,1,2,3-pentachloropropane (HCC-240db) described in the following comprises all three steps 1-3, i.e. telomerisation to produce a 1,1,1,3-tetrachloropropane feedstock, dehydrochlorination of 1,1,1,3-tetrachloropropane to produce 1,1,3-trichloropropene, and chlorination of 1,1,3-trichloropropene to produce 1,1,1,2,3-pentachloropropane, it is noted that the telomerisation (step 1), and dehydrochlorination (step 1) are optional steps, only.
This step of the HCC-240db production involves a selective telomerisation reaction which takes place partially or completely in the principal alkylation zone. In that reaction, carbon tetrachloride is reacted with ethylene to produce a 1,1,1,3-tetrachloropropane. While such reactions are known in the art, one issue with such processes is the production of unwanted impurities.
It has been found that by controlling the degree of completion of the reaction, the production of unwanted impurities can be achieved. Thus, in embodiments, in step 1-a) the concentration of 1,1,1,3-tetrachloropropane in the reaction mixture in the principal alkylation zone is maintained at a level such that the molar ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride in the reaction mixture extracted from the principal alkylation zone does not exceed 95:5 where the principal alkylation zone is in continuous operation, or 99:1 where the principal alkylation zone is in batchwise operation.
In embodiments, the molar ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride in the reaction mixture is controlled in step 1-a) within certain numerically defined limits. As those skilled in the art will appreciate, in such embodiments, while control over the process is characterized herein in terms of the molar ratio between the carbon tetrachloride starting material and 1,1,1,3-tetrachloropropane, it can also considered as control over the conversion of starting material to product—thus a molar ratio of starting material:product of 95:5 equates to a conversion of 5%. The inventors have found that limiting the conversion of the starting material as outlined above minimizes the formation of undesirable impurities. Additionally, where reference is made to a molar ratio of the starting material:product being greater than a given value, this means a greater degree of conversion of the starting material to product, i.e. such that the proportion of the product is increased while the proportion of the starting material is decreased.
In step 1-a) of the process, the reaction mixture is formed by contacting the alkene and carbon tetrachloride. This may occur in the principal alkylation zone, e.g. by both the alkene and carbon tetrachloride being fed into that zone. Additionally or alternatively, the alkene may be contacted with carbon tetrachloride in a zone upstream of the principal alkylation zone and then fed into the principal alkylation zone.
In embodiments, in step 1-a), a primary alkylation zone may be employed, upstream of the principal alkylation zone. The reaction mixture may be formed by feeding carbon tetrachloride and ethylene into the primary alkylation zone to form the reaction mixture which is then fed into the principal alkylation zone. In such an embodiment, the partial conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane may occur in the primary alkylation zone such that that alkane is formed and comprised in the reaction mixture fed into the principal alkylation zone, along with carbon tetrachloride. In additional or alternative embodiments, the amount of ethylene fed into the primary alkylation zone may be limited to retard the conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane in the primary alkylation zone such that the reaction mixture fed into the principal alkylation zone therefrom comprises carbon tetrachloride and 1,1,1,3-tetrachloropropane, but low levels or substantially no ethylene.
The ethylene and carbon tetrachloride employed in step 1-a) may be contacted in a zone (for example, a primary alkylation zone or the principal alkylation zone) by being fed into that zone using any technique or equipment known to those skilled in the art, for example via dispersion devices such as dip tube/s, nozzle/s, ejectors, static mixing devices and/or sparger/s. In such embodiments, the feed of ethylene and/or carbon tetrachloride may be continuous or intermittent. The ethylene supplied as a feed into the zone in which the reaction mixture is formed may be in liquid and/or gaseous form. Likewise, the carbon tetrachloride may be in liquid and/or gaseous form.
In embodiments of the present invention, the reaction mixture (comprising carbon tetrachloride, 1,1,1,3-tetrachloropropane, catalyst and optionally unreacted ethylene) present in the principal alkylation zone (and/or any other alkylation zone that may be employed) may be homogenous, i.e. in a single phase, for example a liquid, or gaseous phase. This can be achieved even where one of the components of the reaction mixture is introduced into the system in a different phase to the other components. For example, in embodiments, gaseous ethylene may be contacted with liquid carbon tetrachloride, causing the ethylene to be dissolved, thus forming a liquid phase homogenous reaction mixture. Alternatively, the reaction mixture may be heterogeneous.
The carbon tetrachloride and ethylene starting materials employed in step 1-a) may have a high degree of purity, for example, either or both of those materials may be at least about 95% pure, at least about 97% pure, at least about 99% pure, at least about 99.5% pure, at least about 99.7% pure, or at least about 99.9% pure.
In embodiments, the carbon tetrachloride starting material comprises less than less than about 2000 ppm, less than about 1000 ppm, less than about 500 ppm, less than about 200 ppm, less than about 100 ppm, less than about 50 ppm or less than about 20 ppm bromides or brominated organic compounds.
Additionally or alternatively, the carbon tetrachloride starting material may have a moisture content of about 200 ppm or less, about 100 ppm or less, about 50 ppm or less or about 35 ppm or less.
The source of carbon tetrachloride may be located on the same site as the apparatus for operating the processes. In embodiments, the source of the carbon tetrachloride may be adjacent to a chlor alkali facility with, for example, a membrane electrolysis plant, from which high purity chlorine will be available to use in the production of the carbon tetrachloride. The site may also comprise plants for producing epichlorohydrin (for example from glycerol feedstock), glycidol, and/or epoxy resin, or oxychlorination plant (e.g. vinyl chloride monomer VCM plant, Perchloroethylene plant etc) or a site with HCl electrolysis plant, such that the hydrogen chloride gas, produced as a byproduct in any associated steps or processes, is effectively also utilized. Thus for best economic use of a chlor alkali facility, an integrated facility with plants for chlorine reactions and capture/re-use of hydrogen chloride is envisioned.
The reaction mixture formed in step 1-a) may be extracted from the principal alkylation zone (and/or, if employed, the primary alkylation zone). This may be conducted on a continuous or intermittent basis. For the avoidance of doubt, where reference is made in the context of step 1-a) of the process of the present invention to the continuous extraction of material from the zones employed in the process of the present invention, this should not be assigned a purely literal meaning. One skilled in the art would recognize that, in such embodiments, material may be removed on a substantially continuous basis while the zone in question is at operating conditions and, if its purpose is to set up a steady state reaction (e.g. an alkylation), once the reaction mixture therein has attained the required steady state.
One of the advantages is that the presence of certain impurities typically observed in commercially supplied ethylene (such as certain organic impurities, e.g. as alcohols, ethers, esters, and aldehydes) can be tolerated and/or removed using process steps outlined herein. The ethylene starting material may be derived from bioethanol, from ethanol or from crude oil. An additional advantage of the processes as described herein is that i) the continuous production of chlorinated alkane and ii) substantially full utilization of the ethylene starting material can be achieved with no escape of the ethylene into the off-gas system.
The amount of unreacted ethylene in the reaction mixture leaving the principal alkylation zone is less than 0.6%, less than 0.3%, less than 0.2%, or less than 0.1%. Any unreacted gaseous ethylene is directly recycled back to the reaction zone/s operating at elevated pressure. Alternatively, the unreacted gaseous ethylene is recycled back to the reaction zone/s operating at elevated pressure by absorbing ethylene into the cold liquid carbon tetrachloride feedstock. Advantageously, the gaseous reagent/s, if needed to be recycled, may be handled without using expensive compressor systems.
One of the advantages of the process of step 1-a) is that it permits the production of 1,1,1,3-tetrachloropropane with high isomeric selectivity. Thus, in embodiments, 1,1,1,3-tetrachloropropane is produced in step 1-a) with isomeric selectivity of at least about 95%, at least about 97%, at least about 98%, at least about 99%, at least about 99.5%, at least about 99.7%, at least about 99.8% or at least about 99.9%
The alkylation reaction conducted in step 1-a) of the process, to produce 1,1,1,3-tetrachloropropane is accelerated through the use of a catalyst. As used herein, the term catalyst is used to encompass not only the use of a single compound or material having catalytic effect, e.g. a solid metal or a metal salt, but a catalyst system which may additionally comprise a catalytic material and a co-catalyst or promoter such as a ligand.
Any catalyst known by those skilled in the art to find utility in the formation of 1,1,1,3-tetrachloropropane from carbon tetrachloride and ethylene may be employed.
In embodiments, the catalyst is metallic. Any metal which can function as a catalyst in the alkylation reaction of the present invention may be employed, including, but not limited to copper and/or iron. The metallic catalyst may be present in its solid form (e.g., in the case of copper or iron, in particulate form (e.g. powder or filings), wire and/or mesh or the like) and/or as a salt in which the metal may be in any oxidation state (e.g. cuprous salts such as cuprous chloride, cuprous bromide, cuprous cyanide, cuprous sulphate, cuprous phenyl and/or ferrous and/or ferric salts such as ferrous chloride and ferric chloride).
Where metallic salts are employed as catalysts in the processes of the present invention, these may be added to the alkylation zone/s and/or form in situ therein. In the latter case, solid metal may be added into the alkylation zone/s and, owing to the conditions therein, the salt may be formed. For example, if solid iron is added into a chlorination reaction mixture, the chlorine present may combine with the elemental iron to form ferric or ferrous chloride in situ. Where metallic salts are formed in situ, it may nevertheless be desirable to maintain a predetermined level of elemental metal catalyst in the reaction mixture (for example, an excess of elemental metal as compared to the level of metallic salt/s and/or ligand) and thus, additional elemental metal catalyst may be added as the reaction proceeds, either continuously or intermittently.
As mentioned above, in embodiments, the catalyst may also comprise a ligand, preferably an organic ligand, which may form a complex with the metallic catalyst. Suitable ligands include amines, nitrites, amides, phosphates and phosphites. In embodiments of the invention, the ligand employed is an alkylphosphate, such as trimethylphosphate, triethylphosphate, tributylphosphate, and triphenylphosphate.
Additional metallic catalysts and ligands are known to those skilled in the art and are disclosed in the prior art, for example, US6187978, the contents of which are incorporated by reference. Such catalysts may be employed in step 1-a).
The components of the catalyst system, where used, may be fed into the alkylation zone/s (e.g. the principal alkylation zone, and/or, if used, the primary alkylation zone) continuously or intermittently. Additionally or alternatively, they may be introduced into the alkylation zone/s (e.g. the principal alkylation zone, and/or, if used, the primary alkylation zone) prior to and/or during commencement of the alkylation reaction of step 1-a).
Additionally or alternatively, the catalyst (or components of the catalyst, for example the ligand) may be fed into the alkylation zone/s (e.g. the principal alkylation zone, or, if used, the primary alkylation zone) together with other components of the reaction mixture, for example in a feed of carbon tetrachloride and/or ethylene.
In embodiments in which the catalyst comprises a metallic catalyst and a promoter such as a ligand, the molar ratio of the promoter:metallic catalyst in the reaction mixture present in the principal alkylation zone, and/or, if used, the primary alkylation zone is maintained at a ratio of greater than 1:1, more preferably at a ratio of greater than 2:1, 5:1 or 10:1.
Where solid metal catalyst is added to the reaction mixture, this may be added into the primary alkylation zone, if used, and/or into the principal alkylation zone. In embodiments, solid metal catalyst is added into the primary alkylation zone, if used, and/or into the principal alkylation zone in amounts to maintain a level of about 0.1 to 4%, about 0.5 to 3% or about 1 to 2% by weight of the reaction mixture.
Additionally or alternatively, where metallic catalysts are employed, these are added to establish a dissolved metal content of about 0.1%, about 0.15% or about 0.2% to about 1.0, about 0.5 or about 0.3% by weight of the reaction mixture.
In embodiments in which the catalyst system employed comprises a metallic catalyst and promoter, the metallic catalyst and promoter can be added to the reaction mixture simultaneously and/or in the same part of the apparatus, for example in the primary alkylation zone (if used) and or the principal alkylation zone.
Alternatively, the metallic catalyst and promoter can be added at different locations in the apparatus, or sequentially or separately. For example, solid metal catalyst can be added to the primary alkylation zone with promoter being fed into that zone from a recycle loop to which additional, fresh promoter may also be added.
In embodiments, the primary and/or principal alkylation zones employed in step 1-a) are operated under atmospheric or superatmospheric pressure, i.e. at a pressure greater than about 100 kPa, greater than about 200 kPa, greater than about 300 kPa, greater than about 400 kPa, greater than about 500 kPa, greater than about 600 kPa, greater than about 700 kPa, or greater than about 800 kPa. Typically, the pressure in the primary and/or principal alkylation zones will be equal to or lower than about 2000 kPa, about 1700 kPa, about 1500 kPa, about 1300 kPa, about 1200 kPa or about 1000 kPa.
Additionally or alternatively, in embodiments, the primary and/or principal alkylation zones employed in step 1-a) are operated at elevated temperatures, i.e. temperatures equal to or greater than about 30° C., about 40° C., about 50° C., about 60° C., about 70° C., about 80° C., about 90° C. or about 100° C. Typically, the primary and/or principal alkylation zones will be operated at temperatures equal to or lower than about 200° C., about 180° C., about 160° C., about 140° C., about 130° C., about 120° C., or about 115° C.
The use of temperatures and pressures within these ranges combined with the other features of step 1-a) have been advantageously found to maximise yields and/or selectivity of 1,1,1,3-tetrachloropropane, while minimising the formation of problematic byproducts.
A plurality of alkylation zones may be employed in step 1-a). Any number of alkylation zones may be employed, for example 1, 2, 3, 4, 5, 6, 7, 8, 9 or 10 or more. In embodiments in which a plurality of primary and/or principal alkylation zones are employed, there may be any number (for example 1, 2, 3, 4, 5, 6, 7, 8, 9, or 10 or more) primary and/or principal alkylation zones present.
For the avoidance of doubt, where reference is made to the properties of an alkylation zone (primary and/or principal), e.g. its operating conditions, its method of operation, its properties, etc., insofar as embodiments disclosed herein are concerned which comprise a plurality of primary and/or principal alkylation zones, one, some or all of those zones may exhibit the property/ies in question. For example, if, for brevity, reference is made to a principal alkylation zone having a specified operating temperature, then, insofar as embodiments including a plurality of principal alkylation zones are concerned, this should be taken as a reference that one, some or all of those principal alkylation zones are operated at the specified temperature. In arrangements where a plurality of primary and/or principal alkylation zones are employed, those alkylation zones may be operated in parallel and/or in series.
In arrangements in which primary and principal alkylation zones are employed in step 1-a), the reaction between ethylene and carbon tetrachloride may be controlled to prevent it proceeding beyond a certain degree of completion in the primary alkylation zone, for example such that the molar ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride in the reaction mixture extracted from the primary alkylation zone and/or fed into the principal alkylation zone does not exceed 85:15, 90:10, 93:7 or 95:5 although this is not essential. Additionally or alternatively, the reaction may be permitted to run to a relatively advanced stage of completion, such that the molar ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride in the reaction mixture extracted from the primary alkylation zone and/or fed into the principal alkylation zone is greater than 50:50, 60:40, 70:30, 75:25 or 80:20.
Control of the progress of the step 1-a) reaction in the primary alkylation zone may be achieved through the use of reaction conditions which do not favour the total conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane. Additionally, or alternatively, control of the progress of the alkylation reaction in the primary alkylation zones may be achieved through careful selection of the residence time of the reaction mixture in the primary alkylation zones, for example about 20 to 300 minutes, about 40 to 250 minutes, about 60 to about 200 minutes or about 90 to about 180 minutes. In embodiments of the invention, the molar ratio may be controlled by limiting the amount of ethylene fed into the primary and/or principal alkylation zones employed in step 1-a) of the invention. For example, the molar ratio of carbon tetrachloride:ethylene fed into the primary and/or principal alkylation zones may range from about 50:50 to about 55:45, about 60:40, about 65:35, about 70:30, about 75:25, about 80:20 about 85:15 or about 90:10.
In embodiments where primary and principal alkylation zones are employed in step 1-a), the bulk of 1,1,1,3-tetrachloropropane may be produced in the primary alkylation zone. In such embodiments, the proportion of 1,1,1,3-tetrachloropropane produced in the principal reaction zone may be significantly lower e.g. such that the molar ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride in the reaction mixture is increased by 1 to 10, 2 to 8 or 3 to 5.
For example, if the molar ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride in the reaction mixture extracted from the primary alkylation zone and fed into the principal alkylation zone is 90:10, that molar ratio may be increased by 2, 3 or 5 in the principal alkylation zone so that the molar ratio of 1,1,1,3-tetrachloropropane:carbon tetrachloride present in the mixture extracted from the principal alkylation zone may be 92:8, 93:7 or 95:5.
However, the viability of the processes of step 1-a) is not dependent on the major part of the conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane occurring in the primary reaction zone. Thus, in alternative embodiments, the degree of conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane may be balanced between the primary and principal alkylation zones, or may be greater in the principal alkylation zone than in the primary alkylation zone.
The reaction mixture may then be taken from the primary alkylation zone (continuously or intermittently) and fed into the principal alkylation zone in which a proportion of the remaining carbon tetrachloride present in the reaction mixture is converted to 1,1,1,3-tetrachloropropane. In such embodiments, any unreacted ethylene starting material present in the reaction mixture may advantageously be fully (or at least nearly fully) utilized.
In process step 1-a) of the process, where employed, the primary and principal alkylation zones may be operated under different conditions. The principal alkylation zone may be operated under a greater pressure than the primary alkylation zone/s, for example at a pressure which is at least about 10 kPa higher, about 20 kPa higher, about 50 kPa higher, about 100 kPa higher, about 150 kPa higher, about 200 kPa higher about 300 kPa or about 500 kPa higher.
According to embodiments, ethylene may not be fed into the principal alkylation zone; the only source of ethylene to those zone's may be in the reaction mixture fed into the principal alkylation zone.
Additionally, in embodiments in which the alkylation reaction between carbon tetrachloride and ethylene is catalyzed by a metallic catalyst (optionally including a ligand), metallic catalyst and/or ligand may not be fed in to the principal alkylation zone. In such embodiments, the sole source of catalyst may be the reaction mixture fed into the principal alkylation zone. Additionally or alternatively, the principal alkylation zone may be provided with a catalyst bed.
In step 1-a) of the process, where primary and principal alkylation zones are employed and solid metal catalyst is present in the reaction mixture in the primary alkylation zone (e.g. by being added directly thereto), when the reaction mixture is extracted from the primary alkylation zone in order to be fed into the principal alkylation zone, the extraction of the reaction mixture from the primary alkylation zone may be carried out such that very little, if any, solid metal catalyst is present in the reaction mixture, for example less than about 5 mg, about 2 mg, about 1 mg, about 0.5 mg, about 0.2 mg, about 0.1 mg of solid metal catalyst per liter of reaction mixture.
This may be achieved through the use of any technique and/or equipment known to those skilled in the art, for example a tube extending into the primary alkylation zone/s at an appropriate location, being provided with a filtering mesh and/or having an appropriate diameter.
Where employed, the primary and principal alkylation zones may be in the same or different reactors, which may be the same or different types of reactors. Further, in embodiments where a plurality of primary alkylation zones are employed, these may be in the same or different reactors. Likewise, in embodiments where a plurality of principal alkylation zones are employed, these may be in the same or different reactors.
Any type of reactor or reactors known to those skilled in the art may be employed in step 1-a of the process of the present invention. Specific examples of reactors that may be used to provide alkylation zones are column reactors (e.g. column gas-liquid reactors), tubular reactors, bubble column reactions, plug/flow reactors (e.g. tubular plug/flow reactors) and stirred tank reactors (e.g. continuously stirred tank reactors).
Arrangements in which the primary alkylation zone is present in a continuously stirred tank reactor (CSTR) and the principal alkylation zone is present in a plug/flow reactor have provided advantageous results.
One advantage of step 1-a) of the process of the present invention is that desirous results are obtained whether the alkylation zones (e.g. the primary alkylation zone and/or the principal alkylation zone) are operated in a continuous (steady state) or batchwise process. The terms ‘continuous process’ and ‘batchwise process’ will be understood by those skilled in the art. In embodiments, the primary alkylation zone, where employed, is operated in a continuous or batchwise process. Additionally or alternatively, the second alkylation zone/s, where employed, are operated in a continuous or batchwise process.
In embodiments of step 1-a), where the principal alkylation zone is in continuous operation, the content of 1,1,1,3-tetrachloropropane may be controlled such that the ratio of that compound:carbon tetrachloride in the reaction mixture extracted from the principal alkylation zone does not exceed about 94:6, about 92:8, or about 90:10.
In alternative embodiments of step 1-a) of the process where the principal alkylation zone is in batchwise operation, the content of 1,1,1,3-tetrachloropropane may be controlled such that the ratio of that compound:carbon tetrachloride in the reaction mixture extracted from the principal alkylation zone does not exceed about 97:3, about 95:5, or about 90:10.
Regardless of whether the principal alkylation zone is in continuous or batchwise process, the content of 1,1,1,3-tetrachloropropane may be controlled such that the ratio of that compound:carbon tetrachloride in the reaction mixture extracted from the principal alkylation zone is equal to or greater than about 70:30, about 80:20, about 85:15, or about 90:10.
It has surprisingly been found that by controlling the degree of conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane, and preventing the reaction from proceeding to completion, the formation of impurities is advantageously reduced. For example, in embodiments in which the ethylene feedstock employed in the processes is ethylene, the production of undesired byproducts such as pentanes (which would otherwise be formed) is minimized.
Thus, in embodiments, reaction mixture formed in step 1-a) and extracted from the principal reaction zone comprises serial reaction products, i.e. compounds comprising a greater number of carbon atoms than 1,1,1,3-tetrachloropropane, of less than about 5%, less than about 2%, less than about 1%, less than about 0.5%, less than about 0.2%, less than about 0.1%, less than about 0.05% or less than about 0.02%.
Control of the content of 1,1,1,3-tetrachloropropane may be achieved by retarding the progress of the alkylation process and/or by introducing additional carbon tetrachloride into the principal alkylation zone.
In embodiments of step 1-a) in which the content of 1,1,1,3-tetrachloropropane is controlled by retarding the alkylation process, this can be achieved through the use of reaction conditions which do not favor the total conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane. For example, this can be achieved through exposing the reaction mixture, or at least a portion thereof, to conditions which decelerate or halt the progress of the alkylation reaction. In such embodiments, the pressure that the reaction mixture is exposed to in the alkylation zone's (for example, the principal alkylation zone/s, where employed) could be reduced significantly, e.g. by at least about 500 kPA, by at least about 700 kPa, by at least about 1000 kPa.
Additionally or alternatively, the pressure to which the reaction mixture is exposed can be reduced to atmospheric or subatmospheric pressure. The reduction in pressure can occur in one or more alkylation zone (for example, one, some or all of the principal alkylation zones, if used). Additionally or alternatively, the reduction in pressure can occur following extraction of the reaction mixture from the alkylation zone/s.
Additionally or alternatively, in embodiments in which the content of 1,1,1,3-tetrachloropropane is controlled by retarding the alkylation process, this can be achieved through limiting the ethylene level present in the reaction mixture formed in step 1-a of the process of the present invention.
In embodiments, control of the progress of the alkylation reaction in the alkylation zone's may be achieved through careful selection of the residence time of the reaction mixture in the alkylation zone/s. For example, in embodiments in which one or more principal alkylation zones are employed, the residence time of the reaction mixture in those zone's may be, for example about 1 to 120 minutes, about 5 to 100 minutes, about 15 to about 60 minutes or about 20 to about 40 minutes.
In embodiments in which the content of 1,1,1,3-tetrachloropropane is controlled by retarding the alkylation process, this can additionally or alternatively be achieved by reducing the operating temperature of the principal alkylation zone, for example by about 5° C. or more, about 10° C. or more, about 20° C. or more, about 50° C. or more or by about 100° C. or more. Additionally or alternatively, the operating temperature of the principal conversion zone can be reduced to about 20° C., about 10° C. or about 0° C.
Additionally or alternatively, the alkylation process can be retarded by limiting the amount of catalyst present in the reaction mixture, or removing the catalyst bed (if present) from the principal alkylation zone.
The rate of agitation or stirring of the principal alkylation zone can also be reduced to retard the alkylation process.
As mentioned above, the reaction mixture extracted from the principal alkylation zone comprises carbon tetrachloride, catalyst and 1,1,1,3-tetrachloropropane. However, in embodiments, depending on the conditions and equipment employed, the reaction mixture extracted from the principal alkylation zone may additionally comprise unreacted ethylene starting material and/or impurities (e.g. chlorinated alkane impurities, chlorinated alkene impurities and/or oxygenated organic compounds).
Given that the presence of unreacted ethylene alongside 1,1,1,3-tetrachloropropane can be problematic in step 2) of the present invention, in embodiments, the reaction mixture extracted from the principal alkylation zone may be subjected to a dealkenation step (as part of step 1-b)) in which at least about 50% or more by weight of the ethylene present in the reaction mixture is extracted therefrom and at least about 50% of the extracted ethylene is fed back into the reaction mixture provided in the principal alkylation zone.
Such embodiments are particularly advantageous as they enable substantial if not total utilisation of the ethylene feed employed in the processes.
In embodiments, at least about 60%, at least about 70%, at least about 80%, at least about 90%, at least about 95%, at least about 97% or at least about 99% of the ethylene present in the reaction mixture extracted from the principal alkylation zone is removed during the dealkenation step.
The removal of unreacted ethylene from the reaction mixture can be achieved using any technique known to those skilled in the art. In embodiments, extraction of the ethylene from the reaction mixture can be achieved using distillation techniques which result in a stream rich in ethylene being obtained, for example flash evaporation, which may conveniently be deployed in embodiments where the boiling point of the ethylene is substantially lower than the boiling point of the other compounds present in the reaction mixture, as is the case with ethylene (−103.7° C.) vs carbon tetrachloride (76.6° C.) and 1,1,1,3-tetrachloropropane (159° C.).
Dealkenation of the reaction mixture in step 1-b) of the process of the present invention may be selective. In other words, the ethylene is selectively extracted, without the substantial removal of other compounds from the reaction mixture. In such embodiments, the ethylene extracted from the reaction mixture may comprise less than about 10%, less than about 5%, less than about 2% or less than about 1% of compounds other than the ethylene starting material.
In step 1-b) of the process, distillation of the reaction mixture can be achieved, using any techniques or equipment known to those skilled in the art. For example, conventional distillation apparatus (e.g. a distillation column) may be employed. Additionally or alternatively, in embodiments, where pressure in the principal alkylation zone from which the reaction mixture is extracted is superatmospheric, evaporation of ethylene from the reaction mixture may be achieved by maintaining the reaction mixture at a superatmospheric pressure following extraction from the principal alkylation zone and feeding it into an evaporation zone in which evaporation of the ethylene from the reaction mixture occurs.
In embodiments, evaporation of ethylene from the reaction mixture in the evaporation zone in step 1-b) can be achieved by depressurisation, for example, by significantly reducing the pressure that the reaction mixture is under, e.g. by at least about 500 kPA, by at least about 700 kPa, by at least about 1000 kPa, and/or to atmospheric or subatmospheric pressure. Conveniently, in embodiments in which depressurisation is used either partly or totally to decelerate or halt the conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane, and also to separate ethylene from the reaction mixture, these aims can be simultaneously achieved in a single depressurisation step.
The evaporation zone may be in any apparatus in which evaporation of the ethylene present in the reaction mixture can be achieved, for example, flash evaporation apparatus such as a flash drum.
The ethylene distilled off from the reaction mixture in step 1-b), for example by flash evaporation, is preferably extracted from the distillation apparatus in liquid or gaseous form. According to embodiments, at least about 50%, at least about 70%, at least about 80%, at least about 90%, at least about 95%, at least about 97% or at least about 99% by weight of the of the ethylene extracted from the evaporation zone is fed back (i.e. recycled) to the primary and/or principal alkylation zone.
For the avoidance of doubt, in embodiments, the distilled ethylene obtained in step 1-b) of the process, if in gaseous form, may or may not be converted back to a liquid, prior to being fed in to the reaction mixture provided in the principal alkylation zone. For example, conversion of the gaseous ethylene to liquid ethylene may be achieved by being passed through a condenser and/or being trapped in a stream of liquid (preferably cooled) carbon tetrachloride, which can then be fed into the alkylation zone/s. Gaseous ethylene may be trapped in a liquid stream of carbon tetrachloride using any techniques or equipment known to those skilled in the art, for example an absorption column. This arrangement is advantageous as it aids the full industrial utilization of the compounds employed in the alkylation process.
As mentioned above, reaction mixture formed in step 1-a), and extracted from the alkylation zone/s comprises catalyst. Given that the presence of catalyst may be problematic in step 2), it may be preferable to remove the catalyst from the reaction mixture. Step 1-b) may comprise such a removal step.
Additionally, for catalyst systems in which costly catalysts and/or promoters such as the alkylphosphate and alkylphosphite ligands mentioned above are employed, the recovery of reusable catalyst systems and/or components thereof is also preferable to minimise the quantities of fresh catalyst that must be used, thus reducing operational cost.
While the challenge of removing catalysts of the type employed in the processes of the present invention from reaction mixtures has been addressed in the past, the techniques and conditions employed to do so (typically involving distillation using aggressive conditions) can be damaging to the catalyst systems and can reduce their catalytic ability. This is especially the case where the catalyst system is temperature sensitive as is the case for systems including certain organic ligands as promoters, such as alkylphosphates and alkylphosphites.
Thus, in embodiments, step 1-b) may comprise the step of subjecting reaction mixture extracted from the alkylation zone/s to an aqueous treatment step in which the reaction mixture is contacted with an aqueous medium in an aqueous treatment zone, a biphasic mixture is formed and an organic phase comprising catalyst is extracted from the biphasic mixture.
In embodiments in which reaction mixture formed in step 1-a) is subjected to an aqueous treatment step in step 1-b), the reaction mixture may comprise unreacted carbon tetrachloride and 1,1,1,3-tetrachloropropane. Additionally, the reaction mixture comprises catalyst (for example a complex of the metallic catalyst and catalyst ligand, or the free catalyst ligand) and/or unreacted ethylene starting material.
Through the use of the aqueous treatment step in step 1-b of the process described, the damaging conditions described in the prior art (for example, high temperature, high catalyst concentration and/or the presence of iron compounds in anhydrous form) can be avoided, meaning that the recovered catalyst and/or components thereof (e.g. the ligand or promoter) can be re-used (for example, it can be recycled back to the reaction mixture provided in the alkylation zone/s) without any substantial loss in catalytic ability. In embodiments, the steam stripping of the biphasic aqueous treated mixture is preferred as the boiler temperatures in excess of 100° C. can be avoided and atmospheric pressure can be employed.
A further advantage of the aqueous treatment step in step 1-b) is that it results in the removal of impurities from the reaction mixture, for example, oxygenated organic products, if present. Advantageously, the levels of such materials in the reaction mixture are significantly reduced to acceptable levels, if not eliminated, by the aqueous treatment step.
In embodiments in a which an aqueous treatment step is performed in step 1-b), the reaction mixture provided in the aqueous treatment zone may comprise 1,1,1,3-tetrachloropropane (for example in amounts of about 50% or greater), catalyst, and optionally carbon tetrachloride and/or impurities, for example organic oxygenated compounds, chlorinated alkane compounds (other than 1,1,1,3-tetrachloropropane) and or chlorinated alkene compounds.
This catalytic recovery process in step 1-b) involves the reaction mixture being subjected to an aqueous treatment step in which the reaction mixture is contacted with an aqueous medium in an aqueous treatment zone. In embodiments, the aqueous medium is water (as a liquid and/or vapour). Additionally, the aqueous medium may additionally comprise other compounds, such as acids. Inorganic acids, such as hydrochloric acid, sulfuric acid and/or phosphoric acid may be employed.
Where the aqueous medium fed into the aqueous treatment zone is partially or totally in liquid form, a biphasic mixture will be formed upon the liquid aqueous medium contacting the reaction mixture.
Alternatively, where the aqueous medium is in gaseous form, e.g. steam, a biphasic mixture may not be formed immediately, but only once the gaseous aqueous medium condenses. The apparatus employed in aqueous treatment step may be configured such that condensation of the aqueous medium to form the biphasic mixture occurs within and/or remote from the aqueous treatment zone.
For example, 1,1,1,3-tetrachloropropane may be extracted from the mixture formed in the aqueous treatment zone. The majority (e.g. at least about 50%, at least about 60%, at least about 70%, at least about 80% or at least about 90%) of 1,1,1,3-tetrachloropropane present in the reaction mixture fed into the aqueous treatment zone may be extracted from the mixture formed in the aqueous treatment zone using any techniques or equipment known to those skilled in the art.
In embodiments, distillation is used to extract 1,1,1,3-tetrachloropropane from the mixture formed in the aqueous treatment zone. The distillation may result in a stream rich in 1,1,1,3-tetrachloropropane being obtained.
As used throughout this specification, the term ‘a stream rich in’ a specific compound (or corresponding language) is used to mean that the stream comprises at least about 90%, about 95%, about 97%, about 98% or about 99% of the specific compound. Further, the term ‘stream’ should not be interpreted narrowly, but encompasses compositions (including fractions) extracted from a mixture via any means.
For example, 1,1,1,3-tetrachloropropane may be distilled off, for example, from a gaseous mixture comprising that alkane and water vapour. 1,1,1,3-tetrachloropropane may be distilled off in a stream rich in 1,1,1,3-tetrachloropropane. This may be used as the feedstock for step 2-a). In embodiments in which the aqueous medium is partly or totally in liquid form, distillation of 1,1,1,3-tetrachloropropane may be achieved by boiling the mixture present to evaporate the 1,1,1,3-tetrachloropropane and produce the gaseous 1,1,1,3-tetrachloropropane/water vapour mixture from which 1,1,1,3-tetrachloropropane can be distilled, for example using steam distillation techniques.
Additionally or alternatively, where the aqueous medium is provided partly or totally in gaseous form, this evaporates 1,1,1,3-tetrachloropropane to form the gaseous mixture comprising that alkane and water vapour which can then optionally be subjected to distillation to remove 1,1,1,3-tetrachloropropane, for example steam distillation. 1,1,1,3-tetrachloropropane may be obtained in a stream rich in that compound.
In embodiments in which 1,1,1,3-tetrachloropropane is distilled from a gaseous mixture of 1,1,1,3-tetrachloropropane and water vapour, the distillation apparatus may be coupled to the aqueous treatment zone so that the gaseous chlorinated alkane/water vapour mixture can pass directly from the aqueous treatment zone to that apparatus. Alternatively, the distillation apparatus may be located remotely from the aqueous treatment zone such that the gaseous mixture is firstly extracted from the aqueous treatment zone and then conveyed to the distillation apparatus. In either arrangement, 1,1,1,3-tetrachloropropane may be obtained in a stream rich in that compound.
In alternative embodiments, where the aqueous medium and reaction mixture are in liquid form, 1,1,1,3-tetrachloropropane may be extracted from that liquid mixture using conventional distillation techniques known to those skilled in the art. 1,1,1,3-tetrachloropropane may be obtained in a stream rich in that compound. This stream may be used as the feedstock in step 2-a) of the process of the present invention.
The biphasic mixture may be formed in step 1-b) within the aqueous treatment zone or remotely therefrom. The biphasic mixture comprises an aqueous phase (as a result of the aqueous medium added to the aqueous treatment zone) and an organic phase (comprising 1,1,1,3-tetrachloropropane, optionally unreacted carbon tetrachloride, and importantly catalyst).
To maximize the volume of the organic phase and thus facilitate extraction of that phase from the biphasic mixture, a haloalkane extraction agent (e.g. carbon tetrachloride and/or 1,1,1,3-tetrachloropropane) may be added to the biphasic mixture (e.g. by being continuously or intermittently fed into the aqueous treatment zone) using techniques and equipment known to those skilled in the art.
The organic phase can be extracted from the biphasic residue using any technique known to those skilled in the art, e. g. decantation. For example, extraction of the organic phase can be performed by the sequential phase extraction from the aqueous treatment zone or the vessel in which it is contained. Alternatively, the biphasic mixture can be extracted from the aqueous treatment zone and subjected to a phase separation step remote from the aqueous treatment zone.
In embodiments, the biphasic mixture and/or the extracted organic phase can be filtered. In embodiments, this will result in a filter cake being obtained which can optionally be totally or partially employed as a source of iron.
Extraction of 1,1,1,3-tetrachloropropane from the mixture formed during the aqueous treatment step may be performed prior to extraction of the organic phase therefrom, and/or after the organic phase is extracted from that mixture. Some exemplary embodiments in which 1,1,1,3-tetrachloropropane is extracted from the mixture formed during the aqueous treatment step are outlined above.
As a further example, the biphasic mixture may be heated to form a gaseous mixture from which 1,1,1,3-tetrachloropropane can be extracted (optionally as a stream rich in 1,1,1,3-tetrachloropropane—which may be used as the feedstock in step 2-a)), e.g. via distillation. The organic phase, having a reduced proportion of 1,1,1,3-tetrachloropropane, may then be extracted from the biphasic mixture.
Additionally or alternatively, the organic phase may be extracted from the biphasic mixture as discussed above. 1,1,1,3-tetrachloropropane may then be extracted (optionally as a stream rich in 1,1,1,3-tetrachloropropane—which may be used as the feedstock in step 2-a)) from that phase, e.g. via distillation. In such embodiments, where the organic phase comprises catalyst, the distillation conditions selected to extract 1,1,1,3-tetrachloropropane are mild so as to minimise deactivation of the catalyst system, for example at a temperature of about 100° C. or lower, about 95° C. or lower, about 90° C. or lower, about 85° C. or lower or about 80° C. or lower, and/or at a pressure of about 1 to 10 kPa. Lower pressures can additionally or alternatively be used.
The extracted organic phase may comprise carbon tetrachloride and/or 1,1,1,3-tetrachloropropane. Additionally, the organic phase may comprise catalyst (for example the complex of a metallic catalyst and catalyst ligand or free ligand) and/or unreacted ethylene starting material. Once a stream rich in 1,1,1,3-tetrachloropropane (which may be used as the feedstock in step 2-a) of the process of the present invention) has been extracted from the mixture formed in the aqueous treatment step (either directly, or following extraction of the organic phase therefrom), the content of 1,1,1,3-tetrachloropropane of that phase will be lower than in the reaction mixture.
In arrangements, especially those in which the organic phase comprises carbon tetrachloride and/or catalyst, the organic phase may be fed back to the alkylation zone/s, for example in liquid form. In such arrangements, ethylene starting material (e.g. in gaseous form) may be trapped in the organic phase stream being fed into the alkylation zone/s.
In embodiments, one or more distillation steps in addition to those discussed above may be performed in step 1-b), optionally to obtain stream/s rich in specific products. For example, prior to an aqueous treatment step, if performed, the reaction mixture can be subjected to a distillation step. In embodiments in which the reaction mixture contains a temperature sensitive catalyst system, e.g. one including an organic ligand as a promoter, the distillation step is typically conducted under conditions to avoid deactivation of the catalyst, for example at a temperature of about 100° C. or lower, about 95° C. or lower, about 90° C. or lower, about 85° C. or lower or about 80° C. or lower, and/or at a pressure of about 1 to 10 kPa. Lower pressures can additionally or alternatively be used.
Additionally, it has been found that the inactivation of temperature sensitive catalyst systems can be avoided by not over-distilling the reaction mixture. Thus, in embodiments of the invention in which reaction mixture containing a catalyst system is distilled in step 1-b), distillation may not be permitted to result in the volume of the process liquid in the distillation apparatus being reduced such that the concentration of the catalyst system in that process liquid is about 2×, about 5× or about 10× higher than the level of that catalyst system present in the reaction mixture provided in the principal alkylation zone.
A distillation step conducted in step 1-b) prior to the aqueous treatment step (if performed) can be carried out using techniques and equipment known to those skilled in the art, for example, a distillation boiler (batch or continuous) in communication with a vacuum distillation column. In such an embodiment, the reaction mixture subjected to distillation may comprise greater than about 50% by weight of 1,1,1,3-tetrachloropropane, catalyst, and optionally carbon tetrachloride and/or impurities, for example organic oxygenated compounds, chlorinated alkane compounds (other than 1,1,1,3-tetrachloropropane) and or chlorinated ethylene compounds.
The distillation step typically results in the removal of chlorinated alkane distillate stream/s, for example stream/s of (and optionally rich in) unreacted carbon tetrachloride, 1,1,1,3-tetrachloropropane, and/or chlorinated organic impurities (i.e. chlorinated organic compounds other than 1,1,1,3-tetrachloropropane and carbon tetrachloride) from the reaction mixture. The carbon tetrachloride may be recycled back to the alkylation zone/s. The residue from such a step, which typically comprises quantities of 1,1,1,3-tetrachloropropane, carbon tetrachloride and/or catalyst, may be subjected to further treatment steps, e.g. an aqueous treatment step and/or further distillation step/s.
In embodiments, where the reaction mixture is subjected to a distillation step as part of step 1-b) prior to the aqueous treatment step (if performed), at least about 30%, at least about 50%, at least about 60% or at least about 70% to at most about 95%, at most about 90%, at most about 85% or at most about 80% by weight of 1,1,1,3-tetrachloropropane of interest is removed from the reaction mixture in that distillation step.
One or more distillation steps may additionally or alternatively be performed in step 1-b) following the aqueous treatment step (if performed). For example, the 1,1,1,3-tetrachloropropane extracted from the reaction mixture fed into the aqueous treatment zone may be present in the form of a mixture comprising, as the major constituent, the 1,1,1,3-tetrachloropropane, a haloalkane extraction agent, as well as chlorinated organic impurities (i.e. chlorinated organic compounds other than 1,1,1,3-tetrachloropropane and carbon tetrachloride). That mixture may be subjected to one or more distillation steps, to remove chlorinated organic impurities, to obtain a stream rich in 1,1,1,3-tetrachloropropane and/or to remove the haloalkane extraction agent. Again, any equipment or conditions known to those skilled in the art may be employed in such a distillation step, for example a distillation boiler (batch or continuous) in communication with a vacuum distillation column.
In such a distillation step, 1,1,1,3-tetrachloropropane extracted from the reaction mixture provided in the aqueous treatment zone may be subjected to distillation to separate 1,1,1,3-tetrachloropropane of interest from chloroalkane impurities. For example, a distillation step to purify 1,1,1,3-tetrachloropropane extracted from the reaction mixture provided in the aqueous treatment zone has been found to be particularly effective in removing chloropentane/chloropentene impurities.
Chlorinated organic impurities separated from mixtures comprising 1,1,1,3-tetrachloropropane in distillation steps performed at any stage in processes of the present invention may be retrieved and re-used in the production of carbon tetrachloride. This may be achieved by subjecting the chlorinated organic impurities to a high temperature chlorinolysis process. In such a process, any chlorinated organic compounds present are re-processed mainly back to pure tetrachloromethane in high yields. Thus the use of a chlorinolysis step in the processes of the present invention is useful to maximize the overall yield of the synthesis and purity of the target chloroalkane while minimizing waste production.
In embodiments of the invention, a residue of ‘heavies’ may be formed in a distillation boiler if used following the aqueous treatment step. The ‘heavies’ residue is typically extracted from the system and treated, for example, to a high temperature chlorinolysis process preferably leading to the production of chloromethanes.
Step 1) of the process described is particularly advantageous as it enables highly pure 1,1,1,3-tetrachloropropane feedstock to be produced, using simple and straightforward techniques and equipment with which one skilled in the art would be familiar.
As is apparent, step 1) of the process as outlined above can be employed to provide highly pure 1,1,1,3-tetrachloropropane feedstocks. In embodiments, the feedstock obtained in step 1-b) of the process of the present invention comprises:
This step involves the dehydrochlorination of 1,1,1,3-tetrachloropropane to produce 1,1,3-trichloropropene which is conducted in a dehydrochlorination zone.
It has unexpectedly been found that by controlling the level of 1,1,3-trichloropropene such that the molar ratio of that product to the 1,1,1,3-tetrachloropropane starting material does not exceed 50:50 advantageously prevents the formation of unwanted and problematic impurities, such as chlorinated oligomers which can adversely affect catalyst performance. Doing so also improves yield and catalyst activity. Advantageously, the processes described are also highly selective. Thus, in embodiments, in step 2-a) of the process, the concentration of 1,1,3-trichloropropene in the reaction mixture present in the dehydrochlorination zone may be controlled such that the molar ratio of 1,1,3-trichloropropene:1,1,1,3-tetrachloropropane is from 1:99 to 50:50.
The molar ratio of 1,1,3-trichloropropene:1,1,1,3-tetrachloropropane in the reaction mixture formed in step 2-a) is controlled within numerically defined limits. As those skilled in the art will appreciate, in such embodiments, while control over the process is characterized herein in terms of the molar ratio between 1,1,1,3-tetrachloropropane and 1,1,3-trichloropropene, it can also be considered as control over the conversion of starting material to product—thus a molar ratio of 1,1,3-trichloropropene:1,1,1,3-tetrachloropropane of 20:80 equates to a conversion of 20%. The inventors have found that limiting the conversion 1,1,1,3-tetrachloropropane as outlined above minimizes the formation of undesirable impurities and allows better catalyst lifetime. Additionally, where reference is made to a molar ratio of 1,1,3-trichlorpropene: 1,1,1,3-tetrachloropropane being greater than a given value, this means a greater degree of conversion of the 1,1,1,3-tetrachloropropane to 1,1,3-trichloropropene, i.e. such that the proportion of the 1,1,3-trichloropropene is increased while the proportion of 1,1,1,3-tetrachloropropane is decreased. Moreover, the inventors have surprisingly found out that the required molar ratio between the 1,1,1,3-tetrachloropropane product and the 1,1,3-trichloropropene starting material in the reaction mixture can be controlled not only by significantly limiting the conversion of 1,1,1,3-tetrachloropropane but, advantageously, also by efficient immediate extraction of produced 1,1,3-trichloropropene from such reaction mixture. According to embodiments the process in step 2-a) is continuous.
Step 2-a) of the process results in the formation of 1,1,3-trichloropropene. As those skilled in the art will recognize, 1,1,3-trichloropropane is reactive and the formation of oxygenated organic compounds, such as chlorinated alkanols, or chlorinated alkanoyl compounds in dehydrochlorination reactions of this type is possible. The importance of minimizing such compounds in steps 2-a) and 2-b) has been recognized by the inventors of the present process. While the exclusion of air from the apparatus can reduce the formation of oxygenated compounds, doing so is typically more technically and economically demanding, especially where subatmospheric pressure environments are used.
The in situ formation of such side products can be prevented through use of step 2) of the process, and this is especially advantageous in continuous processes. The reaction conditions described herein enable 1,1,3-trichloropropene to be produced selectively and be extracted from the reaction mixture, such that there is minimized risk of the production of undesired oxygenated compounds.
Additionally or alternatively, if oxygenated compounds are formed in the process of the present invention, e.g. alkanols or carbonyl compounds, then these can be removed through the use of an aqueous treatment step in step 2-b) of the process, discussed below in more detail.
Advantageous results have also been achieved when the content of 1,1,3-trichloropropene in the reaction mixture in step 2-a) is controlled such that the molar ratio of 1,1,3-trichloropropene: 1,1,1,3-tetrachloropropane in the reaction mixture does not exceed 40:60, 30:70, 25:75, 20:80 or 15:85. Additionally or alternatively, in embodiments, the molar ratio of 1,1,3-trichloropropene:1,1,1,3-tetrachloropropane in the reaction mixture may be equal to or greater than 2:98, 5:95 or 10:90.
Any technique or equipment may be used by those skilled in the art to determine the composition of the reaction mixture in step 2-a). For example, a direct determination of the composition can be made e.g. by providing the reaction zone with a port through which samples of the reaction mixture can be extracted for analysis and/or taking samples of reaction mixture upon extraction of that reaction mixture from the dehydrochlorination zone, e.g. via a port located at or in the vicinity of the outlet of the reaction zone. Additionally or alternatively, an indirect determination of the composition can be made e.g. by temperature control as temperature is a function of composition at constant pressure.
The level of 1,1,3-trichloropropene in the reaction mixture in step 2-a) may be controlled in one or more of the following ways: i) by removing 1,1,3-trichloropropene from the dehydrochlorination zone (either directly, or by firstly extracting reaction mixture from the dehydrochlorination zone and then extracting 1,1,3-trichloropropene therefrom), ii) control of the operating conditions in the dehydrochlorination zone (e.g. temperature, pressure, agitation speed, etc) which do not favour higher levels of 1,1,3-trichloropropene formation, and/or iii) by controlling the amount of 1,1,1,3-tetrachloropropane and/or catalyst present in the dehydrochlorination zone.
1,1,3-trichloropropene may be extracted from the reaction mixture on a continuous or batch-wise basis.
In step 2-b), 1,1,3-trichloropropene may be extracted from the reaction mixture formed in step 2-a) using any technique known to those in the art. In embodiments, step 2-b), 1,1,3-trichloropropene is extracted from the reaction mixture via distillation. Regardless of how extraction of 1,1,3-trichloropropene from the reaction mixture is carried out, 1,1,3-trichloropropene may be obtained as a stream rich in 1,1,3-trichloropropene. This stream can be used as the feedstock in step 3-a) of the process of the present invention.
As used throughout this specification, the term ‘a stream rich in’ a specific compound (or corresponding language) is used to mean that the stream comprises at least about 90%, about 95%, about 97%, about 98% or about 99% of the specific compound. Further, the term ‘stream’ should not be interpreted narrowly, but encompasses compositions (including fractions) extracted from a mixture via any means.
For the avoidance of doubt, where reference is made to ‘continuous extraction’ of the reaction mixture in the dehydrochlorination zone or to reaction mixture from the dehydrochlorination zone, a strict literal interpretation is not intended; one skilled in the art would recognise that the term is used to mean that extraction occurs on a substantially continuous basis, once the dehydrochlorination zone has attained the target operating conditions and the reaction mixture has attained a steady state.
1,1,3-trichloropropene can be extracted directly from the reaction mixture in the dehydrochlorination zone (e.g. via direct distillation as part of step 2-b)), or a portion of the reaction mixture formed in step 2-a) can be firstly extracted from the dehydrochlorination zone (on a continuous or batchwise basis) and 1,1,3-trichloropropene extracted from that mixture, remotely from the dehydrochlorination zone.
In embodiments, the reaction mixture may be subjected to additional treatment steps in step 2-b), for example one or more distillation steps and/or aqueous treatment steps (discussed below in more detail). Such additional treatment steps may be carried out before and/or after extraction of 1,1,3-trichloropropene from the reaction mixture. Those skilled in the art will recognize that where such additional treatment steps are conducted post-extraction of 1,1,3-trichloropropene, the 1,1,3-trichloropropene content of the mixture will be lower than that in the reaction mixture formed in the dehydrochlorination zone.
In step 2-b), 1,1,3-trichloropropene may be removed from the reaction mixture by distillation. Any technique and apparatus known to those skilled in the art may be employed to effect extraction of 1,1,3-trichloropropene from the reaction mixture in this way. In embodiments of the invention, a distillation column may be used, for example a rectification column. The reaction mixture may pass or be fed into the column bottom, with 1,1,3-trichloropropene being removed from the top of the column as a liquid distillate.
For example, in embodiments, in which the reaction mixture is totally or partially gaseous, for example due to the operating temperature in the dehydrochlorination zone, the apparatus may be configured such that the dehydrochlorination zone is in fluid communication with the apparatus for conducting the distillation. In such embodiments, the distillation apparatus may be coupled to the dehydrochlorination zone. Conveniently, this enables the gaseous 1,1,3-trichloropropene-containing mixture to pass (or be passed) directly from the dehydrochlorination zone in to the distillation apparatus. Alternatively, the distillation apparatus may be located remotely from the dehydrochlorination zone, meaning that the gaseous mixture must be extracted from the dehydrochlorination zone and passed to the distillation apparatus.
Additionally or alternatively, where the reaction mixture is present in the dehydrochlorination zone either partly or totally in liquid form, a portion of the liquid reaction mixture may be extracted from the dehydrochlorination zone and passed to distillation apparatus. In such embodiments, the reaction mixture may be subjected to one or more treatment steps in step 2-b) (e.g. an aqueous treatment step, discussed below) which may precede and/or follow distillation.
In embodiments where extraction of 1,1,3-trichloropropene from the reaction mixture in step 2-b) occurs in apparatus remote from the dehydrochlorination zone, the resulting mixture, comprising unreacted 1,1,1,3-tetrachloropropane starting material and depleted levels of 1,1,3-trichloropropene (if any) may be fed back into the dehydrochlorination zone.
In embodiments in which 1,1,3-trichloropropene is extracted from the reaction mixture formed in step 2-a), at least about 30%, at least about 40%, at least about 50%, at least about 60%, at least about 70%, at least about 80%, or at least about 90% by weight of 1,1,3-trichloropropene present in the reaction mixture is extracted from that mixture.
In step 2-b), distillation of 1,1,3-trichloropropene from the reaction mixture can be carried out continuously, semi-continuously or batch-wise.
An advantage of the method described is that the dehydrochlorination reaction produces highly pure gaseous hydrogen chloride from the reaction mixture that may be recovered using routine techniques, for example by condensation of distillation apparatus overhead vapours.
Thus, in embodiments in which hydrogen chloride is produced during the dehydrochlorination reaction (step 2-a)), the hydrogen chloride may be extracted. This can be achieved using any equipment and/or techniques for doing so known to those skilled in the art. For example, if the reaction mixture is subjected to distillation, the distillation apparatus may be provided with a condenser (e.g. a partial condenser), or a condenser (e.g. a partial condenser) may be provided downstream of the distillation apparatus, to enable the removal of hydrogen chloride gas.
Cooling apparatus (e.g. a second condenser) may additionally be employed, e.g. downstream of the first condenser. Arranging the apparatus in this way is advantageous as the first condenser can be used to condense the bulk of the 1,1,3-trichloropropene present, with the second condenser being used to purify the gas by condensing traces of 1,1,3-trichloropropene. The recovered 1,1,3-trichloropropene is highly pure (and may be used as a feedstock in step 3-a) of the process of the present invention) as is the hydrogen chloride.
Additionally or alternatively, an absorption column may be employed to absorb hydrogen chloride gas to produce hydrochloric acid solution.
In embodiments, in which hydrogen chloride gas is extracted from the dehydrochlorination zone or from reaction mixture extracted therefrom, this may be achieved through the use of deep cooling, i.e. by extracting the gas from the reaction mixture and then cooling it to a temperature of about 0° C. or lower, about −10° C. or lower or about −20° C. or lower. The resulting condensate may be recycled back to the dehydrochlorination zone or optionally used in other associated reaction zones, e.g. hydrochlorination of glycerol.
Advantageously, hydrogen chloride extracted in these ways is highly pure and thus can be used as a reactant in upstream or downstream reactions in the same industrial plant. An example of downstream use is for the hydrochlorination of glycerol to make monochlorohydrin or dichlorohydrin, and subsequently to lead to epichlorohydrin, glycidol and epoxies.
As mentioned above, in step 2-a), the rate of the reaction (and thus the molar ratio of 1,1,1,3-tetrachloropropane:1,1,3-trichloropropene) can be controlled by modification of the operating temperature in the dehydrochlorination zone. In embodiments, the dehydrochlorination reaction is carried out in the liquid phase, i.e. the reaction mixture is in the liquid form. In such embodiments, the dehydrochlorination zone may be operated at a temperature of about 50° C., about 60° C., about 70° C., about 80° C., about 100° C., about 120° C. or about 130° C. to about 160° C., about 170° C., about 200° C., about 250° C. or about 300° C.
In step 2-a), the reaction mixture is maintained in the dehydrochlorination zone for a period sufficient to enable the reaction (the conversion of 1,1,1,3-tetrachloropropane to 1,1,3-trichloropropene) to proceed to the required degree of completion. In embodiments, in which dehydrochlorination occurs in the liquid phase, the residence time of the reaction mixture in the dehydrochlorination zone may range from about 0.1, about 0.2, about 0.5, about 1, about 1.5, about 2, about 2.5 or about 3 to about 5 hours, about 7 hours, about 9 hours or about 10 hours. The dehydrochlorination zone may be operated in step 2-a) at subatmospheric pressure, atmospheric pressure or superatmospheric pressure. In embodiments, the dehydrochlorination zone is operated at atmospheric pressure or a pressure of about 10 kPa to about 400 kPa, about 40 kPa to about 200 kPa, or about 70 kPa to about 150 kPa.
Any catalyst which increases the rate of the dehydrochlorination reaction may be employed in step 2-a). In embodiments, the catalyst comprises a metal. In such embodiments, the metal may be present in solid form (e.g. where the catalyst is iron, it may be present as particulate iron (e.g. iron filings or iron powder) iron mesh, iron wire, packing (structured or random), fixed bed, fluid bed, dispersions in liquid, etc. or in alloys containing iron formed in any such way, e.g. carbon steel), and/or as a salt (e.g. where the catalyst is iron, it may be present as ferric chloride, ferrous chloride, etc). Additionally or alternatively, the apparatus in which the process is conducted may be provided with components formed either partially or totally of catalyst material, for example column internals.
In embodiments in which metal is present in the reaction mixture as a salt, it may be added to the reaction mixture in salt form and/or solid metal may be added to the reaction mixture, which then dissolves in the reaction mixture, forming the salt in situ. When present in the form of a salt, the catalyst may be added in amorphous form, crystalline form, anhydrous form and/or in hydrated form (e.g. ferric chloride hexahydrate). Liquid form catalysts may also be employed. In alternative embodiments, the dehydrochlorination reaction in step 2-a) is carried out in the vapour phase, i.e. both the 1,1,1,3-tetrachloropropane and the 1,1,3-trichloropropene are in gaseous form. In such embodiments, the dehydrochlorination zone may be operated at a temperature of about 300° C. to about 500° C., about 325° to about 425° C. or about 350° C. to about 400° C.
In embodiments in which the dehydrochlorination reaction occurs in the vapour phase, the residence time of the reaction mixture in the dehydrochlorination zone may range from about 0.5 to about 10 seconds.
It has been surprisingly found that, in embodiments in which the dehydrochlorination reaction in step 2-a) is carried out in the vapour phase, the reaction must be properly catalysed in order to attain high yield and selectivity. Therefore, in processes of the invention, a metallic catalyst may be used, for example one containing iron at levels of 50% by weight or greater.
Thus, in embodiments, there is provided a process for preparing 1,1,3-trichloropropene comprising in step 2-a), contacting 1,1,1,3-tetrachloropropane in the vapour phase with a catalyst having an iron content of 50% or greater in a dehydrochlorination zone to produce a vapour-phase reaction mixture comprising 1,1,1,3-tetrachloropropane and 1,1,3-trichloropropene.
Examples of catalysts which may be employed in step 2-a) include stainless steels, for example ferritic and/or austenic steels. Catalysts employed in processes of the present invention preferably have an iron content of at least about 50%, at least about 60%, at least about 70%, at least about 80%, at least about 90% or at least about 95% by weight. Pure iron may be employed as a catalyst.
Catalysts may be employed in step 2-a) in any form, for example fluid bed arrangements and/or fixed bed arrangements. Additionally or alternatively, components of the dehydrochlorination zone comprising the catalyst can be employed. For example, in embodiments in which the dehydrochlorination zone is in a tube reactor, the reactor tubes (or at least the surfaces of those tubes in contact with 1,1,1,3-tetrachloropropane) can be formed (partially or completely) of the catalyst, or be provided with catalytic zones formed of the catalyst.
During operation of the dehydrochlorination reaction (step 2-a)) in the vapour phase, the catalyst may become deactivated. Thus, in such embodiments, the processes include a catalyst recovery step. This step can be achieved using any techniques and/or equipment known to those skilled in the art, for example, by the injection of an oxidant such as oxygen-rich air and/or oxygen into the dehydrochlorination zone. Prior to such a step, the flow of reactants through the dehydrochlorination zone may be stopped and/or the dehydrochlorination zone may be purged (for example with nitrogen gas). If performed, once the catalyst recovery step is completed, the dehydrochlorination zone may again be purged (for example with nitrogen gas) and/or the flow of reactants into the dehydrochlorination zone can be re-started.
In embodiments in which the dehydrochlorination step (step 2-a)) is conducted in the vapour-phase, the reaction mixture extracted from the dehydrochlorination zone is typically in the vapour phase. Those hot product gases may be condensed using any technique and/or equipment known to those skilled in the art, to obtain chlorinated organic compounds in liquid form. For example, the hot reaction mixture can be cooled by indirect cooling methods, quenching (for example using spray nozzles), direct cooling methods, or the like.
Upon cooling the gases to condense the chlorinated organic compounds from the reaction mixture, hydrogen chloride gas may be extracted which can optionally be used in upstream or downstream processes. An example of downstream use is for the hydrochlorination of glycerol to make monochlorohydrin or dichlorohydrin, and subsequently to lead to epichlorohydrin and epoxies.
Regardless of whether the dehydrochlorination step 2-a) occurs in the gaseous or liquid phase, the mixture of chlorinated organics, including 1,1,3-trichloropropene and unreacted 1,1,1,3-tetrachloropropane, as well as impurities may then be subjected to one or more post dehydrochlorination treatment steps (2-b)) as discussed herein (including one or more distillation and/or aqueous treatment steps) to obtain pure 1,1,3-trichloropropene, which may be used as a feedstock in step 3-a) of the process of the present invention.
Any type of reactor known to those skilled in the art may be employed to provide a dehydrochlorination zone in step 2-a) of the process of the present invention. Specific examples of reactors that may be used to provide a dehydrochlorination zone are column reactors, tubular reactors, bubble column reactors, plug/flow reactors and continuously stirred tank reactors.
Step 2-a) of the process of the present invention may be carried out in a single dehydrochlorination zone or in a plurality of dehydrochlorination zones. Where a plurality of dehydrochlorination zones are employed, these may be operated in sequence (i.e. such that reaction mixture is passed along a number of dehydrochlorination zones) and/or in parallel.
In embodiments, where a plurality of dehydrochlorination zones are employed in step 2-a), optionally in cascade mode, these may be in the same or different reactors. For example, where a plurality of (e.g. 1, 2, 3, 4, 5 or more) dehydrochlorination zones are employed, these may be provided in a plurality (e.g. 1, 2, 3, 4, 5 or more) of reactors (e.g. continuously stirred tank reactors) which may each be optimised to have optimised operating conditions such as temperature, residence times,
In an embodiment, a plurality of dehydrochlorination zones may be present in a distillation column that may be employed in step 2-a) of the process of the present invention. In such embodiments, dehydrochlorination may be achieved by reactive distillation, for example where the dehydrochlorination reaction is carried out on trays in a distillation column and/or on packing provided in the column. In embodiments in which reactive distillation is carried out, the distillation column preferably comprises a stripping zone in which 1,1,3-trichloropropene is separated from 1,1,1,3-tetrachloropropane. The stripping zone may be located below the liquid feed.
It has been found that the components of the reaction mixture (e.g. 1,1,3-trichloropropene, hydrogen chloride and/or the starting material) obtainable from the dehydrochlorination reaction which is conducted in step 2-a), can unfavourably interact with certain materials. Thus, in embodiments of the invention, in step 2-a), those parts of the dehydrochlorination zone which, in use, come into contact with the reaction mixture may have an iron content of about 20% or less, about 10% or less or about 5% or less, and/or are formed from non-metallic materials, for example enamel, glass, impregnated graphite (e.g. impregnated with phenolic resin), silicium carbide and/or plastics materials such as polytetrafluoroethylene, perfluoroalkoxy and/or polyvinylidene fluoride. Additionally or alternatively, at least some parts of the dehydrochlorination zone which, in use, come into contact with the reaction mixture may be formed of a nickel-based alloy, such as Hastelloy.
In embodiments, the parts of all equipment employed with which 1,1,3-trichloropropene will contact are formed from suitable materials such as those identified above. One possible exception is where one or more regions of the surfaces of the apparatus employed in the processes are formed of metallic material which is selected to perform as a catalyst.
The inventors have also found that, under certain operating conditions, the exposure of the reactants used in the processes as well as the compounds formed in those processes to sources of oxygen and/or moisture, including air, water vapour and/or water can lead to the formation of unwanted impurities. Thus, in embodiments, dehydrochlorination and/or distillation may be conducted in an inert atmosphere, e.g. in the absence of oxygen.
In step 2-a), 1,1,1,3-tetrachloropropane may be fed into the dehydrochlorination zone using any technique known to those skilled in the art.
The 1,1,1,3-tetrachloropropane feedstock employed in step 2-a) preferably has a purity level of at least about 95%, at least about 97%, at least about 98%, at least about 98.5%, at least about 99%, or at least about 99.5%.
In embodiments, the 1,1,1,3-tetrachloropropane feedstock contains less than or equal to about 1000 ppm, less than or equal to about 500 ppm, less than or equal to 250 ppm or less than or equal to about 100 ppm of chlorinated alkane impurities, for example alkanes having a boiling point which is equal to or greater than the boiling point 1,1,1,3-tetrachloropropane and/or 1,1,3-trichloropropene and/or which, in the reaction conditions are dehydrochlorinated to produce a chlorinated alkene impurity, for example alkenes which have a boiling point within 10° C. of 1,1,3-trichloropropene, which have a boiling point equal to or greater than 1,1,1,3-tetrachloropropane, and/or which are isomers of 1,1,3-trichloropropene.
In additional or alternative embodiments, the 1,1,1,3-tetrachloropropane feedstock contains less than or equal to about 1000 ppm, less than or equal to about 500 ppm, less than or equal to 250 ppm or less than or equal to about 100 ppm of chlorinated alkene impurities, for example alkenes which have a boiling point within 10° C. of 1,1,3-trichloropropene, which have a boiling point equal to or greater than 1,1,1,3-tetrachloropropane, or 1,1,3-trichloropropene, and/or which are isomers of 1,1,3-trichloropropene.
Additionally or alternatively, the 1,1,1,3-tetrachloropropane feedstock comprises less than or equal to about 1000 ppm, less than or equal to about 500 ppm, less than or equal to about 200 ppm, less than or equal to about 100 ppm, less than or equal to about 50 ppm, less than or equal to about 20 ppm or less than or equal to about 10 ppm of tetrachloroethene, hexachloroethane and/or tetrachloropentanes.
One of the advantages of step 2-a) of the process of the present invention is that it permits the production of 1,1,3-trichloropropene with high isomeric selectivity. Thus, in embodiments of the invention, 1,1,3-trichloropropene is produced in step 2-a) with isomeric selectivity of at least about 95%, at least about 97%, at least about 98%, at least about 99%, at least about 99.5%, at least about 99.7%, at least about 99.8% or at least about 99.9%.
The feed of 1,1,1,3-tetrachloropropane and/or catalyst into the dehydrochlorination zone may be continuous or intermittent, as may extraction of the reaction mixture.
A further advantage of step 2-a) is that desirous results are obtained whether the dehydrochlorination zone is operated in a continuous or batch process. The terms ‘continuous process’ and ‘batch process’ will be understood by those skilled in the art.
A still further advantage of step 2-a) of the process is that it enables high purity 1,1,3-trichloropropene to be produced without the use of alkaline hydroxides. Thus, in embodiments, no alkaline hydroxide is added to the dehydrochlorination zone in step 2-a) and/or the reaction medium present in the dehydrochlorination zone in step 2-a) is free of alkaline hydroxide.
As mentioned above, in embodiments, reaction mixture comprising 1,1,1,3-tetrachloropropane, 1,1,3-trichloropropene and catalyst may be extracted from the dehydrochlorination zone. This may be subjected to further treatment steps in step 2-b).
In such embodiments, such a treatment step may be an aqueous washing step in which the extracted mixture is optionally filtered and then fed into an aqueous treatment zone. This step may be carried out before or after extraction of 1,1,3-trichloropropene from the mixture.
The mixture is contacted with an aqueous medium in the aqueous treatment zone which serves to deactivate the catalyst. The mixture may be contacted with acid in the aqueous treatment zone, for example inorganic acid such as sulphuric acid, phosphoric acid and/or hydrochloric acid. The acid may be pure, or may be dilute. Where dilute acid is used, this may provide the aqueous medium. The pH value of the aqueous medium should be sufficiently low to enable effective separation of the biphasic mixture.
The aqueous treatment step comprised in step 2-b) has the advantageous effect of removing certain classes of otherwise problematic impurities from the mixture, especially oxygenated impurities.
In such embodiments, catalyst deactivation can be achieved with only a short contact time, e.g. about 5, about 10, about 20 or about 30 minutes, with water at low temperature being required. For hydrolysis and extraction of chlorinated, oxygenated impurities, the contact time with the water may be longer, e.g. up to about 1 hour, about 2 hours, about 5 hours or about 10 hours and/or at a temperature of about 50° C. or less, about 40° C. or less or about 30° C. or less.
Thus, in embodiments, step 2-b) of the inventive process may comprise the step of removing oxygenated organic impurities from a mixture comprising 1,1,3-trichloropropene, oxygenated organic impurities and optionally a catalyst and/or 1,1,1,3-tetrachloropropane, comprising contacting the mixture with an aqueous medium to form a biphasic mixture and extracting the organic phase from that biphasic mixture. In embodiments, the mixture of this aspect of the invention is or comprises the mixture extracted from the dehydrochlorination zone employed in step 2-a).
Where a dilute acid is employed in such a step, this may additionally provide the aqueous medium with which the mixture is contacted. Additionally, or alternatively, the aqueous medium may comprise water (in any form, e.g. including steam) which may be added separately into the aqueous treatment zone.
In embodiments in which acid is added into the aqueous treatment zone, this preferably reduces the pH of the mixture present therein to about 6 or lower, about 5 or lower, about 4 or lower, about 2 or lower or about 1 or lower.
A proportion (e.g. at least about 30%, at least about 40%, at least about 50%, at least about 60%, at least about 70%, or at least about 80%) of the unreacted 1,1,1,3-tetrachloropropane and/or 1,1,3-trichloropropene may be extracted from the mixture formed in the aqueous treatment zone using any techniques or equipment known to those skilled in the art.
For example, in embodiments in which the mixture is partly or totally in gaseous form, for example due to the operating temperature in the aqueous treatment zone and/or through the addition of steam as the aqueous medium, the gaseous mixture may be subjected to distillation in step 2-b). In such embodiments, the distillation device may be in fluid communication with the aqueous treatment zone (optionally coupled to that zone) or may be remote from the aqueous treatment zone.
Additionally or alternatively, where the mixture is partly or totally in liquid form, that mixture may be extracted from the aqueous treatment zone and subjected to distillation in step 2-b). In embodiments where such a distillation step is conducted in step 2-b), a stream comprising (and optionally rich in) 1,1,1,3-tetrachloropropane and/or 1,1,3-trichloropropene may be obtained. The stream rich in 1,1,3-trichloropropene may be used as the feedstock in step 3-a) of the process of the present invention.
1,1,1,3-tetrachloropropane and/or 1,1,3-trichloropropene extracted from the mixture fed in to the aqueous treatment zone may be recycled back to the dehydrochlorination zone for use as a starting material.
A biphasic mixture, comprising an aqueous phase and an organic phase may be formed in the aqueous treatment zone (or in certain embodiments, remotely therefrom) in step 2-b), as a result of the presence of both the aqueous medium and also the predominantly organic mixture, In such embodiments where a biphasic mixture is formed in step 2-b) of the process of the present invention, the organic phase may be extracted from the biphasic mixture using phase separation techniques and/or equipment known to those skilled in the art. Where the biphasic mixture is formed in the aqueous treatment zone, the organic phase can be separated from the aqueous phase by the sequential extraction of the phases from the aqueous treatment zone. The aqueous phase, which contains impurities removed from the mixture can be further treated. To maximise phase separation efficiency and thus facilitate extraction of that phase from the biphasic mixture, a haloalkane extraction agent and/or phase separation intensifier (for example, 1,1,1,3-tetrachloropropane and/or various alcohols and/or ketones) may be added to the aqueous treatment zone, either intermittently or continuously, using techniques and/or equipment known to those skilled in the art. The use of 1,1,1,3-tetrachloropropane is preferred as this compound is part of the process and thus does not require removal using specific separation steps.
Optionally, phase separation intensifiers such as polar alcohols and/or ketones with boiling points sufficiently different to 1,1,3-trichloropropene and 1,1,1,3-tetrachloropropane may be employed. The difference in boiling points should be at least 20° C., at least about 30° C., at least about 40° C., at least about 50° C. or at least about 60° C. Examples of phase separation intensifiers that may be employed include aliphatic ketones e.g. acetone and aliphatic alcohols e.g. methanol, ethanol, propanol/s, butanol/s.
In embodiments, the extracted organic phase may then be subjected to a distillation step in step 2-b) in which streams of (and optionally rich in) 1,1,3-trichloropropene and/or unreacted 1,1,1,3-tetrachloropropane are distilled off. Such a step may be performed regardless of whether extraction of 1,1,3-trichloropropene from the reaction mixture was carried out prior to aqueous treatment or not. The stream of unreacted 1,1,1,3-tetrachloropropane may be recycled back to the dehydrochlorination zone. The stream rich in 1,1,3-trichloropropene may be used as the feedstock in step 3-a). A heavy ends residue may be extracted from the distillation apparatus, optionally filtered and incinerated and/or subjected to high temperature chlorinolysis.
The organic phase comprising 1,1,1,3-tetrachloropropane and/or 1,1,3-trichloropropene as well as haloalkane extraction agent and/or phase separation intensifier may be fed back in to the dehydrochlorination zone. In such embodiments, a distillation step to remove the phase separation intensifier (if used) or other components of the organic phase may be conducted. Reducing the water content of the chlorinated alkene has been found to use such alkene in downstream applications such as chlorination. Thus, in embodiments of the present invention, the process conditions are controlled such that the obtained chlorinated alkene product's comprise less than about 500 ppm, about 200 ppm or less, about 100 ppm or less or about 50 ppm or less of water.
Step 2) of the process described above is advantageous as it enables highly pure 1,1,3-trichloropropene to be produced using simple and straightforward techniques and equipment with which one skilled in the art would be familiar.
Step 2) of the process results in the production of the 1,1,3-trichloropropene feedstock for use in step 3-a) of the process. That feedstock preferably comprises:
The process of this step of the present invention involves the chlorination of an already chlorinated alkene (1,1,3-trichloropropene) to produce 1,1,1,2,3-pentachloropropane with a high level of purity. The process is highly selective.
It has been found that controlling the degree of conversion of the 1,1,3-trichloropropene starting material to the 1,1,1,2,3-pentachloropropane product advantageously minimizes the formation of unwanted impurities. Thus, in embodiments, in step 3-a) of the process, the molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene in the reaction mixture extracted from the reaction zone does not exceed 95:5.
The molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene in the reaction mixture is controlled within numerically defined limits. As those skilled in the art will appreciate, in such embodiments, while control over the process is characterized herein in terms of the molar ratio between 1,1,1,2,3-pentachloropropane and 1,1,3-trichloropropene, it can also considered as control over the conversion of 1,1,3-trichloropropene to 1,1,1,2,3-pentachloropropane—thus a molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene of 95:5 equates to a conversion of 95%. The inventors have found that, in step 3-a) of the process, limiting the conversion of the starting material as outlined above minimizes the formation of undesirable impurities. Additionally, where reference is made to a molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene being greater than a given value, this means a greater degree of conversion of the 1,1,3-trichloropropene to 1,1,1,2,3-pentachloropropane, i.e. such that the proportion of 1,1,1,2,3-pentachloropropane is increased while the proportion of 1,1,3-trichloropropene is decreased.
In embodiments, the reaction zone may be a primary reaction zone.
One of the advantages of step 3-a) of the process is that it permits the production of 1,1,1,2,3-pentachloropropane with high isomeric selectivity. Thus, in embodiments, 1,1,1,2,3-pentachloropropane is produced in step 3-a) with isomeric selectivity of at least about 95%, at least about 97%, at least about 98%, at least about 99%, at least about 99.5%, at least about 99.7%, at least about 99.8% or at least about 99.9%.
It has been found that highly pure 1,1,1,2,3-pentachloropropane is less susceptible to degradation during storage and transport. It is believed that this is due to the absence (or presence in only trace amounts) of impurities which would otherwise trigger decomposition of 1,1,1,2,3-pentachloropropane. Accordingly, the use of stabilizing agents can advantageously be avoided.
A further advantage of step 3-a) is that, through control of the degree of conversion of the starting material to product, the formation of otherwise problematic serial products is minimized. Accordingly, in embodiments, reaction mixture extracted from the primary reaction zone, and/or 1,1,1,2,3-pentachloropropane rich material extracted from the principal reaction zone, comprises low levels of serial reaction products, i.e. compounds comprising a greater number of chlorine and/or carbon atoms 1,1,1,2,3-pentachloropropane, for example in amounts of less than about 5%, less than about 2%, less than about 1%, less than about 0.5%, less than about 0.2%, less than about 0.1%, less than about 0.05% or less than about 0.02%. In embodiments, the process may be continuous.
It has unexpectedly been found that through the careful control of the level of 1,1,1,2,3-pentachloropropane in the reaction mixture formed in the primary reaction zone in step 3-a), the production of impurities is minimized, and/or high selectivity for 1,1,1,2,3-pentachloropropane, is achieved. The level of 1,1,1,2,3-pentachloropropane in the reaction mixture may be controlled by, for example, i) removing 1,1,1,2,3-pentachloropropane (either specifically, or by extracting reaction mixture) from the primary reaction zone/s, ii) by controlling the reaction conditions in the primary reaction zone (e.g. temperature, exposure to light, and/or pressure), and/or iii) by controlling the amount of 1,1,3-trichloropropene and/or chlorine present in the primary reaction zone.
For example, the amount of chlorine present in the reaction mixture formed in step 3-a) can be controlled such that there is no molar excess of chlorine present in the reaction mixture in the primary and/or principal reaction zone/s.
Any conditions which result in the formation of 1,1,1,2,3-pentachloropropane may be employed in the primary reaction zone used in step 3-a). However, in embodiments, the operating temperature in the primary reaction zone is maintained at a relatively low level, for example about 100° C. or lower, about 90° C. or lower or about 80° C. or lower. The operating temperature of the primary reaction zone may be about −30° C., about −20° C., about −10° C. or about 0° C. to about 20° C., about 40° C., or about 75° C. The use of such temperatures in the primary reaction zone has been found unexpectedly to be advantageous as this results in a reduction in the formation of isomers of 1,1,1,2,3-pentachloropropane and over-chlorinated compounds, yet gives the required product selectively in high yield. To increase the reaction rate at these temperatures, light (visible and/or ultra violet) may optionally be used to promote the addition of chlorine at these low temperatures.
In step 3-a), the operating temperature in the primary reaction zone may be controlled by any temperature control means known to those skilled in the art, for example heating/cooling jackets, heating/cooling loops either internal or external to the reactor, heat exchangers and the like. Additionally or alternatively, the temperature may be controlled by controlling the temperature of material's added into the reaction mixture, thus, controlling the temperature of the reaction mixture. The reaction mixture is maintained in the primary reaction zone for a time and under conditions sufficient to achieve the required level of 1,1,1,2,3-pentachloropropane in the reaction mixture.
In embodiments, the primary reaction zone employed in step 3-a) may be exposed to light, for example visible light and/or ultra violet light. Exposure of the reaction mixture to light promotes the reaction when operated at low temperatures which is advantageous where the use of higher temperatures is to be avoided.
For the avoidance of doubt, in embodiments, the primary conversion step in step 3-a) may be carried out in a plurality of primary reaction zones (e.g. 1, 2, 3, 4, 5, 6, 7, 8, 9, 10 or more primary reaction zones), which may be operated at the same or different pressures, temperatures and/or light conditions.
In step 3-a), the residence time of the reaction mixture in the primary reaction zone may range from about 30 to 300 minutes, from about 40 to about 120 minutes or from about 60 to about 90 minutes.
Optimal results have been observed when the level of 1,1,1,2,3-pentachloropropane in the reaction mixture present in the primary reaction zone is maintained at a level such that the molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene in reaction mixture extracted from the primary reaction zone does not exceed 50:50. In embodiments, the level of 1,1,1,2,3-pentachloropropane present in the reaction mixture in the primary reaction zone may be maintained at lower levels, for example such that the molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene in reaction mixture extracted from the primary reaction zone does not exceed 75:25, 50:50, 40:60 or 30:70. Additionally or alternatively, the level of 1,1,1,2,3-pentachloropropane in the reaction mixture present in the primary reaction zone's is maintained at a level such that the molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene in reaction mixture extracted from the primary reaction zone is at least 5:95, 10:90, 15:85, 20:80, 30:70, 40:60 or 50:50.
The composition of reaction mixture, enabling a determination of the molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene, may be determined as soon as is practicable following extraction of the reaction mixture from the primary reaction zone. For example, a sample of reaction mixture may be extracted at a point adjacent to or slightly downstream of the outlet of the primary reaction zone. In embodiments, the outlet may be located at the upper end of the primary reaction zone.
Reaction mixture comprising 1,1,3-trichloropropene and 1,1,1,2,3-pentachloropropane formed in step 3-a) may be extracted from the primary and/or principal reaction zone. This may be done either continuously or intermittently.
One skilled in the art would recognize that, in embodiments where reaction mixture/1,1,1,2,3-pentachloropropane rich product is extracted from the respective reaction zone, that material may be removed on a substantially continuous basis while the zone in question is at operating conditions and, if its purpose is to set up a steady state reaction (e.g. a chlorination), once the reaction mixture therein has attained the required steady state.
In embodiments, the reaction in step 3-a) conducted in the primary reaction zone is in the liquid phase, i.e., the reaction mixture present therein is predominantly or totally liquid. The reaction mixture may be analysed using any techniques known to those skilled in the art e.g. chromatography.
The 1,1,3-trichloropropene feedstock used in step 3-a) preferably has a high degree of purity. In embodiments, the 1,1,3-trichloropropene feedstock has a purity level of at least about 95%, at least about 97%, at least about 99%, or at least about 99.5%.
Additionally or alternatively, the 1,1,3-trichloropropene feedstock used in step 3-a) may include less than about 2%, less than about 1%, less than about 0.1%, less than about 0.01% or less than about 0.001% by weight of chlorinated alkene and/or chlorinated alkane impurities. For example, the 1,1,3-trichloropropene feedstock may comprise less than about 2%, less than about 1%, less than about 0.1%, less than about 0.01% or less than about 0.001% by weight of chlorinated alkene impurities such as perchlorethylene, tetrachloroethylene, hexachloroethylene, isomeric trichloropropene, tetrachloropropenes and/or chlorinated alkane impurities such as 1,1,1,3-tetrachloropropane.
The feed of chlorine and/or 1,1,3-trichloropropene into the primary reaction zone/s and/or principal reaction zone/s employed in step 3-a) may be continuous or intermittent.
Chlorine may be fed into reaction zone/s employed in step 3-a) of the process of the present invention in liquid and/or gaseous form, either continuously or intermittently. For example, the primary reaction zone may be fed with one or more chlorine feeds. Additionally or alternatively, reaction zone/s downstream of the primary reaction zone (e.g. the principal conversion zone) may be fed with one or more chlorine feeds. In embodiments of the invention, the only reaction zone supplied with chlorine is the primary reaction zone.
Where the reaction mixture in the reaction zone/s is liquid, the chlorine may be fed into the reaction zone/s as gas and dissolved in the reaction zone. In embodiments, the chlorine is fed into reaction zone/s via dispersing devices, for example, nozzles, porous plates, tubes, ejectors, etc. The chlorine, in embodiments, may be fed directly into the liquid reaction mixture. Additionally or alternatively, the chlorine may be fed into liquid feeds of other reactants upstream of the reaction zone/s.
Additional vigorous stirring may be used to ensure good mixing and/or dissolution of the chlorine into the liquid reaction mixture.
The chlorine used as a starting material in step 3-a) is preferably highly pure. In embodiments, the chlorine fed into the reaction zone's employed at any stage in the present invention preferably has a purity of at least about 95%, at least about 97%, at least about 99%, at least about 99.5%, or at least about 99.9%.
Additionally or alternatively, the chlorine used in step 3-a) may comprise bromine or bromide in an amount of about 200 ppm or less, about 100 ppm or less, about 50 ppm or less, about 20 ppm or less or about 10 ppm or less.
The use of chlorine gas comprising low amounts of oxygen (e.g. about 200 ppm or less, about 100 ppm or less, about 50 ppm or less, about 20 ppm or less or about 10 ppm or less) is also envisaged. However, in embodiments, lower grade chlorine (including higher oxygen levels, e.g. of 1000 ppm or higher) can advantageously be employed in step 3-a) without the product of the processes comprising unacceptably high levels of oxygenated impurities.
As mentioned above, it is envisaged that in embodiments, the reaction mixture produced in step 3-a) in the primary reaction zone will be liquid. However, alternative embodiments are envisaged in which the reaction mixture is gaseous. In such embodiments, the primary reaction zone may be operated at temperatures of about 150° C. to about 200° C. Gas phase reactors, for example, one or more tubular gas phase reactors, may be employed in such embodiments.
The term ‘highly pure’ as used in the context of step 3) means about 95% or higher purity, about 99.5% or higher purity, about 99.7% purity, about 99.8% or higher purity, about 99.9% or higher purity, or about 99.95% or higher purity. Unless otherwise specified, values presented herein as percentages are by weight.
Extraction of the reaction mixture from the primary reaction zone can be achieved using any technique known to those skilled in the art. Typically, reaction mixture extracted from the primary reaction zone will comprise unreacted 1,1,3-trichloropropene, unreacted chlorine and 1,1,1,2,3-pentachloropropane. Alternatively, where control of the formation of 1,1,1,2,3-pentachloropropane is achieved by controlling (i.e. limiting) the amount of chlorine fed into the primary reaction zone, the reaction mixture extracted from the primary reaction zone may comprise very low levels of chlorine, for example about 1% or less, about 0.5% or less, about 0.1% or less, about 0.05% or less or about 0.01% or less.
In embodiments, where reaction mixture comprising unreacted 1,1,3-trichloropropene is extracted from the primary reaction zone, a principal conversion step may be performed in step 3-a) in which majority significant proportion, but not all, of the unreacted 1,1,3-trichloropropene present in the reaction mixture extracted from the primary reaction zone is converted to 1,1,1,2,3-pentachloropropane, thus producing a 1,1,1,2,3-pentachloropropane rich product, which is then extracted from the principal reaction zone. The 1,1,1,2,3-pentachloropropane rich product may comprise unreacted 1,1,3-trichloropropene starting material and 1,1,1,2,3-pentachloropropane product.
In such embodiments, the reaction mixture may additionally comprise chlorine. Additionally or alternatively, chlorine may be fed into the principal reaction zone to enable the chlorination reaction to proceed.
The degree of conversion 1,1,3-trichloropropene to 1,1,1,2,3-pentachloropropane is for example controlled such that the molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene present in the 1,1,1,2,3-pentachloropropane rich product extracted from the principal reaction zone does not exceed about 95:5, about 93:7, about 91:9, about 90:10 or about 87.5:12.5.
Additionally or alternatively, the degree of conversion of 1,1,3-trichloropropene to 1,1,1,2,3-pentachloropropane is controlled such that the molar ratio of 1,1,1,2,3-pentachloropropane: 1,1,3-trichloropropene present in the 1,1,1,2,3-pentachloropropane rich product extracted from the principal reaction zone is greater than about 70:30, about 75:25, about 80:20 or about 85:15.
In certain embodiments of step 3-a) in which a principal reaction step is carried out, the molar ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene present in the 1,1,1,2,3-pentachloropropane rich product extracted from the principal reaction zone is greater than that for reaction mixture extracted from the primary reaction zone. In other words, the degree of conversion of the starting material to product is higher for the product extracted from the principal reaction zone than for the reaction mixture extracted from the primary reaction zone. In step 3-a), where a 1,1,1,2,3-pentachloropropane rich product is employed or produced, it may have the 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene ratios outlined above.
It has unexpectedly been found that through the careful control of the degree of conversion of 1,1,3-trichloropropene in the principal reaction zone, the production of impurities is minimized in step 3-a). The level of 1,1,1,2,3-pentachloropropane in the reaction mixture may be controlled by, for example, i) removing the 1,1,1,2,3-pentachloropropane (either specifically, or by extracting 1,1,1,2,3-pentachloropropane rich product) from the principal reaction zone, ii) by controlling the reaction conditions in the principal reaction zone (e.g. temperature, exposure to light, and/or pressure), and/or iii) by controlling the amount of 1,1,3-trichloropropene and/or chlorine present in the principal reaction zone.
In embodiments in which the degree of conversion of 1,1,3-trichloropropene to 1,1,1,2,3-pentachloropropane is controlled (i.e. limited) in step 3-a) by controlling the amount of chlorine present in the principal reaction zone (e.g. supplied directly thereto and/or present as a component of the reaction mixture), the chlorine content in the obtained 1,1,1,2,3-pentachloropropane rich product may be very low, for example about 1% or less, about 0.5% or less, about 0.1% or less, about 0.05% or less or about 0.01% or less.
This principal conversion step will typically take place in one or more principal reaction zones downstream of the primary reaction zone. Any number of principal reaction zones may be employed, for example 1, 2, 3, 4, 5, 6, 7, 8, 9, 10 or more principal reaction zones.
Any conditions which result in the conversion of 1,1,3-trichloropropene to 1,1,1,2,3-pentachloropropane may be employed in the principal conversion step in step 3-a). In embodiments, the principal conversion step may comprise a reduced temperature conversion step. When such a step is performed, the reduction in temperature of the extracted reaction mixture is preferably achieved by feeding the reaction mixture into a principal reaction zone operated at a reduced temperature (for example about −30 to about 30° C., about −25 to about 10° C., or more preferably about −20 to about −10° C.) and extracting a 1,1,1,2,3-pentachloropropane rich product from the principal conversion zone.
It has been unexpectedly found that in step 3-a), maintaining, at low temperature, a reaction mixture comprising 1,1,3-trichloropropene, chlorine and 1,1,1,2,3-pentachloropropane, results in the conversion of 1,1,3-trichloropropene to 1,1,1,2,3-pentachloropropane while minimizing the production of unwanted impurities, improving selectivity and/or the yield.
Thus, in step 3-a), a reduced temperature conversion step may be performed in which a reaction mixture comprising 1,1,3-trichloropropene, and 1,1,1,2,3-pentachloropropane is fed into a principal reaction zone, operated at a temperature of about −30° C. to about 30° C., about −25° C. to about 10° C., or more preferably about −20° C. to about −10° C., and a 1,1,1,2,3-pentachloropropane rich product may then be extracted from the principal reaction zone.
For certain embodiments of step 3-a), exposure of the reaction mixture in the principal reaction zone to light (for example ultra violet light) is useful in conducting the reaction successfully at low temperatures.
In step 3-a), the ratio of 1,1,1,2,3-pentachloropropane:1,1,3-trichloropropene present in the reaction mixture fed in to the principal reaction zone may be 70:30 or lower, 60:40 or lower, 50:50 or lower, 40:60 or lower or 30:70 or lower and/or 5:95 or higher, 10:90 or higher, 20:80 or higher or 40:60 or higher.
In embodiments, in step 3-a), the operating temperature of the principal reaction zone may be achieved in a single cooling action, or a series of cooling actions in which the principal reaction zone's are operated at successively lower temperatures. Operating the principal reaction zone's at reduced temperature can be achieved using any technique known to those skilled in the art.
The reduced temperature conversion step in step 3-a) preferably takes place in one or more principal reaction zones downstream of the primary reaction zone. For example, where the reduced temperature conversion step requires a single cooling action, it may occur in a single principal reaction zone. Where the reduced temperature conversion step requires a series of cooling actions, this may be achieved in a single principal reaction zone, or in a plurality of principal reaction zones.
In embodiments, in step 3-a), the reaction mixture is maintained in the principal reaction zone for a time and under conditions sufficient to achieve the required level of 1,1,1,2,3-pentachloropropane in the reaction mixture.
The principal reaction zone's may be operated under subatmospheric, atmospheric or superatmospheric pressure. Additionally or alternatively, the primary and/or the principal reaction zone's may be exposed to light, for example visible light and/or ultra violet light.
In embodiments, in step 3-a), the residence time of the reaction mixture in the principal reaction zone may range from about 30 to 300 minutes, from about 40 to about 120 minutes or from about 60 to about 90 minutes.
In embodiments, the reaction conducted in the principal reaction zone is in the liquid phase, i.e, the reaction mixture present therein is predominantly or totally liquid.
In embodiments, in step 3-a) reaction mixture extracted from the primary reaction zone is subjected directly to the principal conversion step. In alternative embodiments, the extracted reaction mixture is subjected to one or more pre-treatment steps prior to being subjected to the principal conversion step.
In embodiments, to attain the desired level of 1,1,1,2,3-pentachloropropane in the 1,1,1,2,3-pentachloropropane rich product, the principal conversion step may involve heating the 1,1,1,2,3-pentachloropropane rich product to elevated temperatures, for example to about 20° C. or higher, about 30° C. or higher, about 40° C. or higher, about 50° C. or higher or about 60° C. or higher.
Heating the 1,1,1,2,3-pentachloropropane rich product in this way may be achieved in a single heating step. Alternatively, the 1,1,1,2,3-pentachloropropane rich product may be subjected to a series of heating steps at successively higher temperatures.
As mentioned above, in step 3-a), different reaction zones may be operated at different temperatures, pressure and/or to the exposure to differing types and/or intensity of light. For example, reaction mixture extracted from the primary reaction zone/s could be passed into a first principal reaction zone in which a reduced temperature conversion step is carried out. The obtained 1,1,1,2,3-pentachloropropane rich product could then be passed into a second principal reaction zone downstream of the first principal reaction zone in which a heat treatment or UV exposure step is performed, to convert the bulk of the remaining unreacted 1,1,3-trichloropropene present to 1,1,1,2,3-pentachloropropane. Alternatively, the reduced-temperature conversion step and heating and/or UV exposure steps could all take place in the principal reaction zone.
Thus, in step 3-a), a plurality of principal reaction zones may be employed in sequence. For ease of comprehension, these may be characterized as upstream principal reaction zones and downstream principal reaction zones, the upstream principal reaction zones being upstream of the downstream principal reaction zones when those zones are operated in sequence.
In such embodiments, there may be any number of upstream principal reaction zones and/or downstream principal reaction zones, for example 1, 2, 3, 4, 5, 6, 7, 8, 9 or 10 or more upstream principal reaction zones and/or downstream principal reaction zones.
Where such arrangements are employed, heat treatment and/or light (e.g. ultraviolet light) exposure may be conducted in some or all of the upstream and/or downstream principal reaction zones. The intensity of the light exposure may be higher in the downstream principal reaction zones. Additionally or alternatively, the wavelength of the light to which the reaction mixture is exposed in the downstream principal reaction zones may be lower than that in the upstream principal reaction zones.
In certain embodiments, heat treatment and/or light exposure steps may only be conducted in the downstream principal reaction zones.
One advantage of step 3-a) is that desirous results are obtained whether the primary and/or principal reaction zones are operated in a continuous or batch process. The terms ‘continuous process’ and ‘batch process’ will be understood by those skilled in the art.
Any type of reactor known to those skilled in the art may be employed in step 3-a). Specific examples of reactors that may be used to provide primary reaction zone/s and/or principal reaction zone/s are column reactors (e.g. column gas-liquid reactors), tubular reactors (e.g. tubular gas phase reactors), bubble column reactions, plug/flow reactors and stirred tank reactors, for example continuously stirred tank reactors.
Reactors used in step 3-a) may be divided into different zones each having different flow patterns and/or different operating temperatures/pressures. For example, the principal conversion step may be performed in a reactor including a plurality of principal reaction zones. Those zones may be operated at different temperatures and/or pressures. For example, in embodiments where the principal conversion step is a reduced temperature conversion step, the principal reaction zones may be operated at successively lower temperatures.
Additionally or alternatively, reactors used in step 3-a) may be provided with external circulation loops. The external circulation loops may optionally be provided with cooling and/or heating means.
As those skilled in the art will recognize, in step 3-1), reaction zones can be maintained at differing temperatures through use of cooling/heating elements such as cooling tubes, cooling jackets, cooling spirals, heat exchangers, heating fans, heating jackets or the like.
Some or all of the primary and/or principal reaction zones used in step 3-a) may be exposed to visible light (natural or artificially generated), ultra violet light and/or be operated in darkness. Chlorine, either in liquid, in solution, and/or gaseous form, may be fed into the principal reaction zone/s. 1,1,3-trichloropropene may also or alternatively be fed into the principal reaction zone/s, if required.
Those skilled in art will recognize that, in certain embodiments, the reaction zones utilized at any stage in the processes of the present invention may require agitation means, e.g. stirrers, followers, flow channeling means or the like and the use of such means in the primary and/or principal reaction zones in the processes of the present invention is envisaged. The primary and/or principal reaction zones may be operated with differing flow types of reaction mixture. The primary and/or principal reaction zones employed in step 3-a) may be located within a single or multiple reactors. Thus, for example, in embodiments of the invention, all of the primary reaction zones could be different reaction zones in a single reactor, for example, a column liquid-gas reactor.
Alternatively, the primary reaction zones could be in different reactors (e.g. a series of continuously stirred tank reactors) or even different types of reactors (e.g. one or more primary reaction zones could be in a continuously stirred tank reactor and additional primary reaction zone/s could be in a tube reactor).
The reaction zones employed in step 3-a) may be operated at differing pressures and/or temperatures and/or have differing flows (e.g. flows of differing intensity/direction) of reaction mixture therein.
The reaction zones employed in step 3-a) may be operated in sequence (e.g. where reaction mixture is passed from an initial upstream reaction zone to a terminal downstream reaction zone, optionally via intermediate reaction zones) and/or in parallel.
In embodiments where the reaction zones are operated in step 3-a) in sequence and at differing temperatures and/or pressures, the temperature and/or pressure in some or all of the reaction zones may increase or decrease successively.
One, some or all of the reaction zones employed in step 3-a) may be operated at subatmospheric, atmospheric or superatmospheric pressure.
It has unexpectedly been found that the formation of chlorinated alkane degradation products can be minimized if the apparatus employed to operate step 3) (or at least those parts of it which come into contact with the reaction mixture and/or product streams) does not comprise certain materials.
Thus, in step 3), the apparatus for conducting the step is configured such that those parts of the apparatus which come into contact with 1,1,1,2,3-pentachloropropane and/or 1,1,3-trichloropropene, in use of the apparatus, comprise less than about 20%, about 10%, about 5%, about 2% or about 1% of iron.
In such embodiments of step 3), the apparatus for conducting the process is configured such that those parts of the apparatus which come into contact with 1,1,1,2,3-pentachloropropane and/or 1,1,3-trichloropropene are produced from fluoropolymers, fluorochloropolymers, glass, enamel, phenolic resin impregnated graphite, silicium carbide and/or fluoropolymer impregnated graphite. The combination of glass, PVDF, ETFE and Hastelloy, may be used for achieving a combination of effects, for example to provide the necessary conditions for visible or ultraviolet light to be provided to the reaction mixture while also ensuring that other problems such as corrosion and temperature are controlled.
In step 3-a), the principal reaction zone is for example in a plug/flow reactor. An advantage of the use of such apparatus is that the reactor can be configured to minimize or prevent back flow mixing.
The process steps outlined above minimize the formation of impurities, especially those impurities which are difficult to remove from 1,1,1,2,3-pentachloropropane.
To maximize the purity of the reaction mixture extracted from the primary reaction zone or the 1,1,1,2,3-pentachloropropane rich product obtained from the principal reaction zone, additional purification steps may be carried out in step 3-b). For example, one or more distillation steps may be conducted. Such distillation steps may be conducted under low temperature/reduced pressure conditions.
Additionally or alternatively, one or more hydrolysis steps may be performed in step 3-b). In embodiments in which the reaction mixture/1,1,1,2,3-pentachloropropane rich product (either typically being a mixture comprising 1,1,3-trichloropropene, 1,1,1,2,3-pentachloropropane and impurities including oxygenated organic compounds) is subjected to a hydrolysis step, this typically involves contacting the reaction mixture extracted from the primary reaction zone/1,1,1,2,3-pentachloropropane rich product with an aqueous medium in a hydrolysis zone. Examples of aqueous media which may be employed in the hydrolysis step include water, steam and aqueous acid.
Hydrolysis is conducted at appropriate conditions to allow hydrolysis reaction(s), if any, to proceed.
Performance of a hydrolysis step in step 3-b) is preferable as this reduces the content of oxygenated organic compounds present in the reaction mixture/1,1,1,2,3-pentachloropropane rich product. Examples of oxygenated organic compounds include chlorinated alkanols, chlorinated acid chlorides, chlorinated acids, or chlorinated ketones.
In embodiments in which a hydrolysis step is performed, the reaction mixture/1,1,1,2,3-pentachloropropane rich product subjected to such a step may have an oxygenated organic compound content of about 500 ppm or less, about 200 ppm or less, about 100 ppm or less, about 50 ppm or less, or about 10 ppm or less.
Thus, in embodiments, step 3-b) comprises removing oxygenated organic compounds from a mixture (obtainable from any upstream process) comprising 1,1,1,2,3-pentachloropropane, 1,1,3-trichloropropene and oxygenated organic compounds, comprising feeding the 1,1,1,2,3-pentachloropropane rich product into an aqueous treatment zone, contacting the 1,1,1,2,3-pentachloropropane rich product with an aqueous medium to produce a mixture and extracting i) an organic phase from that mixture or ii) a 1,1,1,2,3-pentachloropropane stream from that mixture, the organic phase/1,1,1,2,3-pentachloropropane stream comprising reduced levels of oxygenated organic compounds as compared to the 1,1,1,2,3-pentachloropropane rich product fed into the aqueous treatment zone.
In processes in which a hydrolysis step is performed in step 3-b), the reaction mixture/1,1,1,2,3-pentachloropropane rich product fed into the aqueous treatment zone may have a low chlorine content, for example about 0.8% or less, about 0.5% or less, about 0.1% or less, about 0.05% or less or about 0.01% or less. For the avoidance of doubt, where reference is made in this context to chlorine, this encompasses free chlorine, unreacted chlorine, and dissolved chlorine. Chlorine which is bonded to atoms other than chlorine should not be considered.
In embodiments, the hydrolysis zone is in a washing tank. In such embodiments, the reaction mixture/1,1,1,2,3-pentachloropropane rich product may be washed with water and/or steam. In step 3-b), once the reaction mixture/1,1,1,2,3-pentachloropropane rich product has been contacted with the aqueous medium to form a mixture in the hydrolysis zone, that mixture may be subjected to one or more treatment steps. For example, components of reaction mixture/1,1,1,2,3-pentachloropropane rich product (e.g. 1,1,1,2,3-pentachloropropane and/or unreacted 1,1,3-trichloropropene) can be extracted from the mixture formed in the aqueous treatment zone, for example via distillation preferably under reduced pressure and/or low temperature. Such a step can be achieved while the mixture is present in the aqueous treatment zone. Additionally or alternatively, the mixture may firstly be extracted from the aqueous treatment zone and subjected to the extraction step remotely from that zone.
Additionally or alternatively, in embodiments, a biphasic mixture may be formed in the aqueous treatment zone in step 3-b). In such embodiments, a phase separation step may be performed in which the organic phase comprising at least 1,1,1,2,3-pentachloropropane from the reaction mixture/1,1,1,2,3-pentachloropropane rich product is separated from the aqueous waste phase. This may be achieved by the sequential extraction of the phases from the aqueous treatment zone. Alternatively, the biphasic mixture could be extracted from the aqueous treatment zone and subjected to a phase separation step remote from the aqueous treatment zone to extract the organic phase.
The organic phase may, after optional filtering, be subjected to distillation to obtain streams comprising purified 1,1,1,2,3-pentachloropropane and/or 1,1,3-trichloropropene. 1,1,3-trichloropropene may be recycled to the primary and/or principal reaction zone's. The purified 1,1,1,2,3-pentachloropropane may be the highly pure 1,1,1,2,3-pentachloropropane product.
Additionally or alternatively, the organic phase can be subjected to additional hydrolysis steps as outlined above in step 3-b) of the process. The hydrolysis steps can be repeated if required, for example, one, two, three or more times.
In embodiments, mixtures comprising 1,1,1,2,3-pentachloropropane (e.g. the reaction mixture obtained from the primary reaction zone, the 1,1,1,2,3-pentachloropropane rich product obtained from the principal reaction zone, the mixture formed in the aqueous treatment zone and/or the organic phase extracted from the biphasic mixture) can be subjected to a distillation step in step 3-b), preferably conducted at a temperature of about 100° C. or lower, about 90° C. or lower or about 80° C. or lower.
Such a distillation step may be conducted under vacuum. Where vacuum distillation is carried out, the vacuum conditions may be selected such that the distillation may be conducted at a low temperature and/or to facilitate the extraction of higher molecular weight chlorinated alkanes. In embodiments, in step 3-b), any distillation steps conducted in the process may result in streams comprising at least about 50%, at least about 80%, at least about 90%, at least about 95%, at least about 97%, at least about 98%, at least about 99%, at least about 99.5%, at least about 99.7%, at least about 99.8%, or at least about 99.9% of i) 1,1,3-trichloropropene and/or ii) 1,1,1,2,3-pentachloropropane. As used herein, the term ‘streams’ should be construed broadly to encompass a composition obtained from any distillation step, regardless of the apparatus used or the form of the composition obtained. Streams of highly pure 1,1,1,2,3-pentachloropropane may be the highly pure 1,1,1,2,3-pentachloropropane product of step 3-b). Any distillation equipment known to those skilled in the art can be employed in step 3-b), for example a distillation boiler/column arrangement. However, it has unexpectedly been found that the formation of chlorinated alkane degradation products can be minimized if distillation apparatus formed of certain materials are avoided.
Thus, in embodiments, step 3-b) comprises the step of distilling a 1,1,1,2,3-pentachloropropane rich product (regardless of the process from which it was obtained), in which distillation apparatus is employed, the distillation apparatus being free of components which, in use of the distillation apparatus, would come into contact with the process fluids (including the liquid or distillate) and comprise about 20% or more, about 10% or more, about 5% or more, about 2% or more or about 1% or more of iron.
In embodiments in which distillation step/s are carried out in step 3-b), the distillation apparatus may be configured such that all of its components which, in use of the distillation apparatus, would come into contact with the distillate or process fluid, are produced from fluoropolymers, fluorochloropolymers, glass, enamel, phenolic resin impregnated graphite, silicium carbide and/or fluoropolymer impregnated graphite.
Where distillation steps are performed as part of step 3-b), streams obtained in such steps which comprise 1,1,3-trichloropropene may be recycled and fed into the primary and/or principal reaction zone/s.
The processes described above are particularly advantageous as they enable highly pure 1,1,1,2,3-pentachloropropane to be produced using simple and straightforward techniques and equipment with which one skilled in the art would be familiar.
In embodiments of the invention, the process of step 3) of the present invention can be used to produce high purity 1,1,1,2,3-pentachloropropane which comprises:
For the avoidance of doubt, the term ‘inorganic compounds of chlorine’ encompasses non-organic compounds containing chlorine, including chlorine (Cl2), hydrogen chloride and phosgene.
In embodiments, the composition may comprise less than about 1000 ppm, less than about 500 ppm, less than about 200 ppm, or less than about 100 ppm of organic compounds other than 1,1,1,2,3-pentachloropropane. Additionally or alternatively, the composition may collectively comprise less than about 0.5%, less than about 0.3%, less than about 0.1% of organic compounds other than 1,1,1,2,3-pentachloropropane.
As can be seen from the disclosure provided herein, the processes of steps 1), 2) and 3) described above can be operated in an integrated process in a fully continuous mode, optionally in combination with other processes. The process steps of the present invention may employ starting compounds which are converted to highly pure intermediates which are themselves further processed to the required target chlorinated compounds. Those compounds have the requisite purity to be employed as feedstocks in a range of downstream processes, for example for hydrofluorination conversions.
Additionally, compositions having the purity profiles corresponding to the products of step 3) are especially well suited to be used as starting materials in the synthesis of fluoroalkanes or fluoroalkenes and/or chlorofluorinated alkenes. Thus, according to a further aspect, there is provided the use of the high purity 1,1,1,2,3-pentachloropropane compositions outlined herein as feedstocks in the synthesis of 2,3,3,3-tetrafluoropropene (HFO-1234yf) and/or 2-chloro-3,3,3-trifluoropropene (HFO-1233xf) and/or 1,1,1,2,2-pentafluoropropane (HFC-245cb).
Main advantages of preferred embodiments of the process for producing 1,1,1,2,3-pentachloropropane (HCC-240db) can be listed as:
According to the process of the invention the 1,1,1,2,3-pentachloropropane (HCC-240db) feedstock obtained at the end of step 3 is reacted with HF in the presence or absence of a catalyst to produce a reaction mixture comprising HCl, HF, and at least one compound chosen from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf).
The method of preparation of the HFO-1234yf is a fluorination reaction of the 1,1,1,2,3-pentachloropropane (HCC-240db) obtained at the previous stage in 2,3,3,3-tetrafluoro-1-propene, the final product.
Embodiments of this step are the following:
In the following, the embodiments indicated above are described in more detail and, furthermore, particular embodiments are described by way of example.
Fluorination Reaction without Using a Catalyst in Gas Phase
According to embodiments, the fluorination reaction of step 4-a) may be conducted without the use of a catalyst. The temperatures, pressures and molar ratio HF:HCC-240db are easily determined by the skilled worker. Typical conditions are given below.
Typically, this step is carried out with a molar ratio HF:HCC-240db from 3:1 to 150:1.
Typically, this step is carried out at a pressure from 1 to 20 bars.
Typically, this step is carried out at a temperature of from 200 to 450° C., preferably from 300 to 430° C.
According to embodiments, wherein a catalyst is used in a gas phase process, the catalyst is for example a catalyst based on a metal including a transition metal oxide or a derivative or halide or oxyhalide of such a metal. Catalysts are e.g. FeCl3, chromium oxyfluoride, chromium oxides (that can optionally be subject to fluorination treatments), chromium fluorides, and mixtures thereof. Other possible catalysts are the catalysts supported on carbon catalysts based on antimony, catalysts based on aluminum (as AlF3 and Al2O3 and oxyfluoride of alumina and aluminum fluoride). Generally speaking, catalysts that can be used are chromium oxyfluoride, aluminium fluoride and oxyfluoride, and supported or unsupported catalyst containing a metal such as Cr, Ni, Zn, Ti, V, Zr, Mo, Ge, Sn, Pb, Mg. Reference can also be made to the disclosures of WO-A-2007/079431, at page 7, lines 1-5 and 28-32, EP-A-939071, at paragraph [0022], WO2008/054781 at page 9 line to page 10 line 34, WO2008/040969 in claim 1, all incorporated herein by reference.
According to embodiments, a particular catalyst is used, which is a chromium based catalyst and more preferably a mixed catalyst containing both chromium and nickel. The molar ratio Cr:Ni, with respect to the metallic element is generally between 0.5 and 5, for example between 0.7 and 2, including close to 1. The catalyst may contain in weight from 0.5 to 20% chromium and 0.5 to 20% nickel, preferably between 2 and 10% of each metal.
The metal may be present in metallic form or as derivatives, including oxide, halide or oxyhalide. These derivatives, including halide and halide oxides, are obtained by activation of the catalytic metal. Although the activation of the metal is not necessary, it is preferred.
The support is preferably made from aluminum. There are several possible carriers such as alumina, activated alumina or aluminum derivatives. These derivatives include aluminum halides and halide oxides of aluminum, for example described in U.S. Pat. No. 4,902,838, or obtained by the activation process described below.
The catalyst may include chromium and nickel in a non-activated or activated form, on a support that has been subjected to activation or not.
Reference can be made to WO2009/118628, and especially to the disclosure of the catalyst from page 4, line 30 to page 7, line 16, which is incorporated herein by reference.
The catalyst can also be a high surface area Cr based catalyst, which is preferably unsupported. The catalyst can optionally contain a low level of one or more co-catalyst such as Co, Zn, Mn, Mg and Ni salt. A preferred co-catalyst is nickel, zinc or magnesium, particularly preferred zinc. Another preferred co-catalyst is nickel. Another preferred co-catalyst is Mg. A disclosure of the high surface area Cr based catalyst can be found in WO2009/158321, pages 4 and 6). According to preferred embodiments, said co-catalyst is preferably present in an amount from about 1-10 wt % of said fluorination catalyst.
Prior to its use, the catalyst may be subjected to activation with air, oxygen or chlorine and/or with HF. The catalyst may be subjected to activation, typically with HF, under suitable conditions. The catalyst is, in embodiments of the process, subjected to an activation process with oxygen or air and HF at a temperature of 100-500° C., preferably from 250-500° C., and more preferably from 300-400° C. The period of activation is preferably from 1 to 200 h and more preferably form 1 to 50 h.
This activation can be followed by a final fluorination activation step in the presence of an oxidizing agent, HF and organics, for example the starting materials of the reaction to be carried out with the activated catalyst. The molar ratio of HF/organics is preferably from 2 to 40 and the molar ratio of oxidizing agent/organics is preferably from 0.04 to 25. The temperature of final activation is preferably from 300 to 400° C.; and more preferably for about 6 to 100 h.
According to embodiments, the liquid phase fluorination reaction is catalyzed. The catalysts may be catalysts known by the person skilled in the art of fluorinations in liquid phase.
One can use an acid of Lewis, a catalyst containing a metal halide, in particular containing halide of antimony, tin, tantalum, titanium, metals of transition such as molybdenum, niobium, iron halides, cesium, oxides of metals of transition, halides of metals of the IVb group, halides of metals of the Vb group, a fluorinated chromium halide, a fluorinated chromium oxide or a mixture of both. One can advantageously use metal chlorides and fluorides. Examples of such catalysts include: SbCl5, SbCl3, TiCl4, SnCl4, TaCl5, NbCl5, TiCl4, FeCl3, MoCl6, CsCl, and their corresponding fluorinated derivatives. Pentavalent metal halides are suitable.
Advantageously one will use a catalyst containing an ionic liquid. These ionic liquids are particularly interesting for fluorination by HF in liquid phase. One will be able to mention the ionic liquids described in patent applications WO2008/149011 (in particular from page 4, line 1 to page 6 line 15, included by reference) and WO01/81353 in the name of the applicant, as well as the reference “liquid-phase HF Fluorination”, Multiphase Homogeneous Catalysis, Ed. Wiley-VCH, (2002), 535.
The fluorination reaction can be implemented in the gas phase. Accordingly, the fluorination process involves contacting the HCC-240db obtained in step 3 with HF in the reaction zone in a gas phase, under conditions sufficient to convert the HCC-240db to the desired fluorination products.
Such conditions, particularly molar ratio HF:HCC-240db, pressure, temperature, contact times conditions are exemplified below in the context of a single step or a two steps fluorination process.
In embodiments, 1,1,1,2,3-pentachloropropane is fluorinated in a single stage process, this process being particularly a gas phase process.
This single stage process is preferably carried out in one reactor, more preferably in one catalytic bed.
Embodiments of the single stage process the following:
According to embodiments, Step 4-a of the process may be a two stage process, comprising the following steps: Step 4-a1 of reacting 1,1,1,2,3-pentachloropropane with HF into product 2-chloro-3,3,3-trifluoropropene; and step 4-a2 of reacting the thus-obtained 2-chloro-3,3,3-trifluoropropene into 2,3,3,3-tetrafluoropropene, wherein step 4-a1 is carried out in a liquid-phase, or in gas phase. The two stages can be implemented continuously or in a discontinuous way, with intermediate storage and/or purification of the HCFO-1233xf.
Step (4-a1): Fluorination of HCC-240db with HF
The reaction can be implemented in two steps, wherein the first step of reacting 1,1,1,2,3-pentachloropropane with HF may be a fluorination reaction in a liquid solvent medium, or a fluorination reaction in the gas phase.
Step (4-a1): Fluorination of HCC-240db with HF in Liquid Phase
In embodiments, the liquid phase process is carried out in an organic phase. Using an organic phase rather than an HF phase favors the reaction into HCFO-1233xf, i.e. there exist conditions that allow fluorination into HCFO-1233xf. Notably, when the reaction is carried out in an organic phase (comprised of the HCC-240db starting material and/or solvent), then HCFO-1233xf can be formed. When HF is added to an initial medium, it will not remain in the medium since it reacts and the amount of HF (or concentration) will be very low, compared to the other products.
The term “organic phase” can thus be defined as referring to a reaction phase comprising the catalyst and the starting material and possibly a solvent if used, but substantially free of HF. Especially the process carried out in an “organic phase” refers to the process in which the initial load does not comprise any HF, in contrast with the prior art.
Because of particular operating conditions, gaseous HCFO-1233xf can be removed from the reactor under gaseous phase, keeping polymerization reactions at a low level.
The liquid phase fluorination of HCC-240db into HCFO-1233xf is carried out in the presence of a catalyst, for example in the presence of a catalyst as exemplified above.
The reaction can be implemented in a liquid solvent medium, the reaction zone being either loaded at the beginning with a starting amount of organic (the starting material) and/or the necessary quantity of solvent, or fed continuously with this quantity of solvent (possibly preliminary mixed with the raw material). When carried out with solvent, it is preferred that the solvent be loaded at the beginning; injections with a view of adjusting the quantity of solvent may however be carried out if necessary.
The reaction conditions (notably pressure) are such that the reactants are liquid. According to an embodiment the reactants are liquid while the reaction product is gaseous. The fact that the reaction products are gaseous allows their recovery in a gaseous phase at the exit of the reaction zone. In embodiments, this stage is in particular implemented under a pressure higher than 2 bar. Advantageously, the pressure lies between 4 and 50 bar, in particular between 5 and 25 bar.
For example, the reaction may be implemented at a temperature ranging between 30° C. and 200° C., preferably between 40° C. and 170° C., advantageously between 50° C. and 150° C.
The molar ratio HF:Organics lies generally between 0.5:1 and 50:1, preferably between 3:1 and 20:1. Values of about 5:1 can be used with advantage. The amount of HF added will correspond to the stoichiometry of the reaction (here 3), to which one will add the amount of HF that is present in the exiting streams (HF and organics) which are usually azeotropic mixtures.
The other reaction conditions, notably flow rates, can be determined by the skilled person according to common general knowledge, depending on the temperature, pressure, catalyst, reactant ratios, and the like.
The solvent, if used, is an inert organic solvent under the reaction conditions. Such a solvent will be generally saturated, advantageously in C2 to C6, in order to avoid the reactions of addition. Such solvents can for example be those mentioned in patent application FR2733227. Such solvents have a boiling point (measured at atmospheric pressure), for example higher than 40° C., advantageously higher than 50° C., in particular higher than 60° C. Higher reaction temperatures will imply higher pressures, so that the boiling point of the solvent under the conditions of reaction is higher than the temperature of implementation of the reaction.
One can in particular mention as a solvent the saturated compounds of ethane, propane or butane, substituted by at least two atoms of halogen, chosen among chlorine and fluorine, or a mixture thereof. As an example one can mention 1,2-dichloroethane, 1,2,3-trichloropropane, 1-chloro-1-fluoroethane, 1,1-difluoroethane, 1,1-dichloroethane and 1,3-dichloro-1-fluorobutane, tetrachlorofluoropropane isomers, trichlorodifluoropropane isomers and dichlorotrifluoropropane isomers, 1,1,1,3,3-pentafluorobutane and 1,1,2-trichloro-2,2-difluoroethane, or a mixture thereof. Nitrated solvents like nitromethane or nitrobenzene and sulfones like tetramethylene sulfone (also known as sulfolane) or dimethyl sulfone may also be used. A preferred solvent is the 1,1,2-trichloro-2,2-difluoroethane (HCFC-122). One can also use possibly reactive solvents, in so far as the product of their reaction is a nonreactive solvent.
The solvent can be present in a quantity for a dilution ratio from at least 20%, preferably between 20% and 80%, advantageously between 40% and 60%.
One can operate with variable ratios catalyst/organic (including solvent if used), but in general one will prefer that this molar ratio lies between 2 mol % and 90 mol %, preferably between 4 mol % and 80 mol % and more preferably between 6 mol % and 75 mol %.
It is also possible that the product of the reaction be stripped using a light gas allowing its drive by mechanical entrainment. Removing gaseous HCFO-1233xf from the liquid phase reactor keep polymerization reactions at a low level (since polymerizable material is in a low amount in the medium) as well as side-reactions (such as addition onto the double bond of the HCFO-1233xf). The addition of a gaseous compound can be advantageous for the reaction, which can be favored for example by the improvement of agitation (bubbling).
This gas can be inert as the nitrogen or helium or the gas can be preferably HCl. When HCl is used, the reaction performs despite the addition into the medium of HCl, which is a reaction product.
Advantageously, this added gas is anhydrous hydrochloric acid. The flow of the stripping gas is determined according to the operating conditions. For example, the flow of HCl, compared to the flow of starting product is such that the molar ratio HCl:starting product lies between 0.5:1 and 5:1, advantageously, between 1:1 and 3:1.
The fluorination process in liquid phase can be implemented continuously or semi-continuously.
According to embodiments, the process is continuous.
Step (4-a1): Fluorination of HCC-240db with HF in Gas Phase
Step (4-a1) of the two stage process may be a gas phase reaction.
In embodiments, HCC-240db is catalytically fluorinated in gas phase into HCFO-1233xf. The catalyst used may be for example the same type of catalyst as described above.
The present fluorination process involves contacting HCC-240db with HF in the reaction zone in a gas phase, under conditions sufficient to convert the HCC-240db to fluorination products comprising mainly HCFO-1233xf.
Typically, this process is carried out with a molar ratio HF:HCC-240db from 3:1 to 150:1, preferably 4:1 to 70:1, more preferably 5:1 to 50:1.
Typically, this process is carried out at a pressure from 1 to 20 bars, preferably 3 to 15 bars, more preferably 5 to 10 bars.
Typically, this process is carried out at a temperature of from 200 to 450° C., preferably from 300 to 430° C., more preferably from 320 to 420° C. The temperature of the bed can be substantially uniform in the reactor or can be adjusted along the path of the stream, decreasing or increasing along the direction of flow.
The temperature of this step 4-a1) is usually less than the temperature of step 4-a2), preferably by at least 30° C.
Contact times (catalyst volume divided by the total flow rate of reactants and co-feeds, adjusted to the operating pressure and temperature) are typically from 6 to 100 sec, preferably from 10 to 80 sec, more preferably from 15 to 50 sec.
Step (4-a2): Reaction of HCFO-1233xf into HFO-1234yf.
In embodiments, the second stage of this method of preparation of the HFO-1234yf is a fluorination reaction of the 2-chloro-3,3,3-trifluoro-1-propene (HCFO-1233xf) obtained at the previous stage in 2,3,3,3-tetrafluoro-1-propene, the final product.
In embodiments, this second stage can comprise direct fluorination in the presence of HF, on a catalyst, in gas phase.
This gas phase reaction is carried out in the presence of a fluorination catalyst. The reaction is carried out in a single gas-phase reactor. The temperatures, pressures and contact times are easily determined by the skilled worker. Typical conditions are given below.
The level of the conversion and selectivity of the final product can vary according to the processing conditions. The catalyst can be present in any suitable form, such as fixed or fluidized bed, preferably in a fixed bed. The direction of flow may be downward or upward.
This catalyst is for example a catalyst as exemplified above, supported or unsupported, activated, where appropriate. Furthermore, a co-catalyst may be used. Suitable supports and and co-catalyst are mentioned above, and may also be used in connection with this specific embodiment. An activation step can be performed as described above.
This step of the process, as well as the entire process, is preferably run continuously.
In embodiments, the HCFO-1233xf fluorination process involves contacting HCFO-1233xf with HF in the reaction zone in a gas phase, under conditions sufficient to convert the HCFO-1233xf to fluorination products comprising, HFO-1234yf and optionally HFC-245cb. Such conditions are given below.
Typically, this step is carried out with a molar ratio HF:HCFO-1233xf from 3:1 to 150:1, preferably 4:1 to 70:1, more preferably 5:1 to 50:1.
Typically, this step is carried out at a pressure from 1 to 20 bars, preferably 5 to 15 bars, more preferably 7 to 10 bars.
Typically, this step is carried out at a temperature of from 200 to 450° C., preferably from 300 to 430° C., more preferably from 320 to 420° C. The temperature of the bed can be substantially uniform in the reactor or can be adjusted along the path of the stream, decreasing or increasing along the direction of flow.
Contact times (catalyst volume divided by the total flow rate of reactants and co-feeds, adjusted to the operating pressure and temperature) are typically from 6 to 100 sec, preferably from 10 to 80 sec, more preferably from 15 to 50 sec.
The entire process step 4-a, as well as the steps 4-a1 and or 4-a2 mentioned above may be implemented continuously or in a discontinuous way, preferably the process and partial steps thereof run continuously, which from an industrial point of view is highly desirable.
According to embodiments, the reactants (starting product and HF) and other compounds used in the reaction (chlorine, oxygen) can be fed in the reactor at the same location, at different locations, or using staged feeding at staged locations along the reactor. A preferred feeding system is to vaporize the reactants in the reactor. The reactants may in this case by heated by the recycle stream further fed in the reactor.
Reactions are implemented in a dedicated reactor for reactions involving halogens. Such reactors are known to those skilled in the art and can include linings based e.g. Hastelloy™, Inconel™, Monel™ or fluoropolymers. The reactor may also include means of heat exchange, if necessary.
The catalyst used for the fluorination reaction can be regenerated. The regeneration may be carried out by contacting the spent catalyst with an oxidizing agent-containing gas flow. The oxidizing agent used is oxygen or air or an oxygen/nitrogen mixture or chlorine. When the regeneration is carried out with air or an oxygen/nitrogen mixture, the proportion of oxygen can range from 20 to about 100 mol. % relative to the mixture of oxygen plus nitrogen.
In another embodiment, the regeneration can be carried out with oxygen or air or an oxygen/nitrogen mixture or chlorine and HF. The proportion of oxygen can range from about 2 to about 98 mol. % relative to the mixture of oxygen plus HF, and from about 20 to about 100 mol. % relative to the mixture of oxygen plus nitrogen.
The temperature during the regeneration may range from 250 to 500° C., preferably from 300 to 450° C., more preferably from 350 to 400° C.; with a contact time of from 1 to 200 s, preferably from 1 to 150 s, more preferably from 5 to 100 s; and for a time of from 1 to about 1500 hours, preferably from 2 to 1000 hours, more preferably from 4 to 500 hours, most preferably from 10 to 200 hours, in particular from 15 to 150 hours. The regeneration can be carried out at a pressure from atmospheric pressure to 20 bars. In a preferred embodiment, the temperature during the regeneration can range from about 250 to 500° C., with a contact time of from about 1 to 200 s, for a time of from 10 to 200 hours and at a pressure from atmospheric pressure to 20 bars.
In a particular embodiment, the fluorination of HCC-240db, obtained preferably according to the present process, into HCFO-1233xf or HFO-1234yf may be carried out alternatively with a regeneration step of the catalyst used for said fluorination of HCC-240db, preferably according to the present process, into HCFO-1233xf or HFO-1234yf. In a preferred embodiment, a first and a second reactor can be used. Indeed, the first reactor can be used to carry out the fluorination reaction (step 4-a) or step 4-c)) while in the second reactor the regeneration of the spent catalyst is carried out. When the fluorination reaction is ended in the first reactor, the regeneration is carried out therein while the fluorination reaction is carried out in the second reactor with a regenerated catalyst.
The final product is readily recovered by any means known in the art, such as by scrubbing, washing, extraction, decantation and preferably distillation. It can also be further purified by distillation techniques.
A polymerization inhibitor can be used for example to extend the catalyst life, typically in a concentration of from about 50-1000 ppm, more preferably between 100-500 ppm. The polymerization inhibitor can be p-methoxyphenol, t-amylphenol, limonene, d,1-limonene, quinones, hydroquinones, epoxides, amines and mixtures thereof. The preferred polymerization inhibitor is p-methoxyphenol or t-amylphenol. The co-feeding of a low level of a polymerization inhibitor can control such polymerization of chloroolefins and extend the life of the catalyst as described in U.S. Pat. No. 5,714,651, incorporated herein by reference.
Oxygen and/or Chlorine Co-Feed
In embodiments, an oxygen and/or chlorine co-feed may be used to extend the catalyst lifetime, typically in an amount of from 0.05 to 15 mole %, preferably 0.5 to 10 mole % of oxygen or chlorine per pentachloropropane molecule. The oxygen can be introduced as an oxygen-containing gas such as air, pure oxygen, or an oxygen/nitrogen mixture.
Step 4-b: Separating the reaction mixture obtained in step 4-a) into a first stream comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf) and/or HCl, and a second stream comprising HF, and 1,1,1,2,2-pentafluoropropane (HFC-245cb) and/or 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf)
Removal of 2,3,3,3-Tetrafluoropropene (HFO-1234yf) and/or HCl
According to embodiments, the process described herein may comprise a step of separating 2,3,3,3-tetrafluoropropene (HFO-1234yf) and/or HCl from the product mixture obtained in step 4-a) above.
The process carried out according to embodiments of separation step 4-b can be implemented as depicted in
The process carried out according to embodiments of separation step 4-b can be implemented as depicted in
In other embodiments, the reaction stream exiting the gas-phase reactor can be recycled in part to the reactor, before it is subjected to the separation into a first, light, stream and a second, heavy stream. The recycling ratio can be as high as 0.7. This recycling allows dilution of HCFO-1233xf which is very reactive and avoids polymerization.
In step 4-a of the process described herein, HCFO-1233xf may be produced along with HFO-1234yf, and HCFO-1233xf and HFC-245cb are separated and recycled into the gas phase reactor according to embodiments of the invention.
Hence, the process may also comprise further steps of: separating the reaction mixture into a first stream comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf) and a second stream comprising 2-chloro-3,3,3-trifluoro-1-propene (HCFO-1233xf); and recycling at least a part of the second stream at least in part back to step (4-a).
According to embodiments, the reaction mixture comprising HCl, HF, and at least one compound chosen from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf), which is obtained in step 4-a, and/or the second stream obtained in step 4-b) is converted to 2,3,3,3-tetrafluoropropene (HFO-1234yf).
In embodiments, the reaction mixture obtained in step 4-a, i.e. comprising at least one compound chosen from 1,1,1,2,3-pentachloropropane (HCC-240db), 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) is converted to 2,3,3,3-tetrafluoropropene (HFO-1234yf). According to embodiments, the reaction mixture obtained in step 4-a can be for conversion at least in part recycled to process step 4-a. Alternatively or additionally, also the second stream obtained in step 4-b), i.e. the reaction mixture after separation of HFO-1234yf and/or HCl may be converted to 2,3,3,3-tetrafluoropropene (HFO-1234yf), particularly with recycling of the second stream to process step 4-a.
Step (4-c) may be a fluorination reaction, preferably in gas phase, of the second stream of step (4-b) with HF in the presence of a catalyst and it may comprise mainly fluorination of 2-chloro-3,3,3-trifluoro-1-propene obtained in step (4-a) in 2,3,3,3-tetrafluoro-1-propene, the final product.
Step 4-d: Separating Product Stream 4-c) into a First Stream Comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf) and/or HCl, and a Second Stream Comprising HF, and 1,1,1,2,2-pentafluoropropane (HFC-245cb) and/or 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf);
As explained above in connection with process step 4-b, also the product stream obtained in process step 4-c, i.e. a product stream comprising HCl, HF, and at least one compound chosen from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf) may be separated according to embodiments into a first stream comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf) and/or HCl, and a second stream comprising HF, and 1,1,1,2,2-pentafluoropropane (HFC-245cb) and/or 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf).
Optional method step 4-d has been added to step 4 of the present process to express without any doubt that a separation step may be conducted after the fluorination of HCC-240db (step 4-a) as well as after a further reaction of the reaction mixture obtained in 4-a) or the second stream obtained in 4-b) to obtain a reaction mixture comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf). For the separation reference is made to separation step 4-b above.
At least part of the second stream obtained in step 4-b) can be recycled to step 4-a) or at least part of the second stream obtained in 4-d) can be recycled to either step 4-a) or 4-c) If a recycling is used, one can recycle directly at the inlet of the reactor or at an intermediate stage of the reactor, for example a separate dip pipe.
Step 4-f: Separating 2,3,3,3-Tetrafluoropropene (HFO-1234yf) from the First Stream Obtained in 4-b) or 4-d)
The first stream may be further separated into HCl and 2,3,3,3-tetrafluoropropene (HFO-1234yf). HFO-1234yf can be separated using appropriate known methods, preferably in a distillation step. For example a separation unit may be used, as a distillation column.
Step 4-g: Purifying at least one compound selected from 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf), HCl or HF, as obtained in one or more of steps 4-a), 4-b), 4-c), 4-d), 4-e), 4-f).
According to embodiments, the 2,3,3,3-tetrafluoropropene stream may undergo one or more further purification steps, i.e. the HFO-1234yf separated can subsequently be subjected to a purification stage in order to remove impurities and possibly traces of HF.
As used herein the term “light” organic impurities is meant to denote impurities with low boiling points, i.e. boiling points between the atmospheric boiling point of HCl and HFO-1234yf.
The process of production of 2,3,3,3-tetrafluoropropene (HFO-1234yf) may, according to embodiments, also comprise a step of separation of 2,3,3,3-tetrafluoropropene (HFO-1234yf) and HF from a reaction mixture or stream as produced according to one of steps 4 described above, comprising said two compounds, particularly the reaction mixture or stream obtained in step 4-b, and step 4-d described above, and for recovering the thus-separated HFO-1234yf and HF.
A respective method has been disclosed in applicant's own patent application WO 2013/007906, incorporated herein by reference.
In embodiments, therefore, the first stream of step 4-b) and/or 4-d) is cooled, optionally in the presence of an added amount of at least one compound (C1) chosen from chlorocarbons, hydrochlorocarbons, hydrochlorofluorocarbons, optionally fluorinated alcohols, optionally fluorinated ethers, ketones, esters, polyols and hydrofluorinated ethers in order to give an upper phase rich in HF and a lower organic phase rich in HFO-1234yf and optionally compound C1.
This stage of cooling in the presence of at least one compound C1 makes it possible to obtain an upper phase which is richer in HF with a very small amount of HFO-1234yf, which phase is capable of being used without any purification stage. The HF thus recovered can be directly recycled to a hydrofluorination reaction stage.
The lower organic phase comprises the compound C1, HFO-1234yf and possibly organic impurities. This organic phase can be subjected to a distillation stage in order to separate the compound C1 and HFO-1234yf. The compound C1 can be recycled to the cooling stage and/or to the reaction stage resulting in the formation of HFO-1234yf.
The HFO-1234yf/HF molar ratio in the composition to be separated is preferably between 0.5 and 2.5 and advantageously between 1.1 and 2.1.
The 2,3,3,3-tetrafluoropropene is preferably present in an azeotropic or quasi azeotropic amount with the HF in the composition to be separated.
In embodiments, the compound C1 to be added to the composition for the cooling stage is a hydrohalocarbon compound which preferably comprises three carbon atoms. Mention may in particular be made of pentachloropropanes, in particular 1,1,1,2,3-pentachloropropane (HCC-240db), 1,1,2,2,3-pentachloropropane (HCC-240aa) and 1,1,1,2,2-pentachloropropane (HCC-240ab); tetrachlorofluoropropanes, in particular 1,1,2,3-tetrachloro-1-fluoropropane (HCFC-241db); trichlorodifluoropropanes; dichlorotrifluoropropanes, in particular 1,2-dichloro-3,3,3-trifluoropropane (HCFC-243db); chlorotetrafluoropropanes, in particular 2-chloro-1,1,1,2-tetrafluoropropane (HCFC-244bb); tetrachloropropenes, in particular 1,1,2,3-tetrachloropropene (HCO-1230xa) and 1,1,1,2-tetrachloropropene (HCO-1230xf); and chlorotrifluoropropenes, in particular 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf).
Preferably, the compound to be added is the same as that which has reacted with HF to give the HFO-1234yf or the compound to be added is an intermediate in the hydrofluorination reaction resulting in the manufacture of the HFO-1234yf.
As the HFO-1234yf is prepared by the hydrofluorination reaction on HCC-240db, the compound C1 is preferably HCC-240db or HCFO-1233xf.
Likewise, when the HFO-1234yf is prepared by the hydrofluorination reaction on HCFO-1233xf, the compound C1 is preferably HCFO-1233xf
When the compound C1 is different from that which has reacted with HF to give HFO-1234yf, the preferred compound C1 is chosen from optionally fluorinated alcohols, optionally fluorinated ethers, ketones, esters, polyols and hydrofluorinated ethers.
Mention may in particular be made, as alcohol, of those having an alkyl group of 1 to 5 carbon atoms. The alcohol can also be fluorinated and the preferred fluorinated alcohol is chosen from the alkyl groups of 1 to 3 carbon atoms.
The ketones of formula RCOR′ with R and R′, which are identical or different, each representing an alkyl group of 1 to 5 carbon atoms may be suitable.
The esters of formula RCOOR′ with R and R′, which are identical or different, each representing an alkyl group of 1 to 5 carbon atoms may be suitable.
The ethers of formula ROR′ with R and R′, which are identical or different, each representing an alkyl group of 1 to 7 carbon atoms may be suitable.
The ethers may be partially or completely fluorinated. When the ethers are partially fluorinated, they are denoted by hydrofluorinated ether.
Preference is given, as hydrofluorinated ether, to that having a boiling point of between 0 and 250° C., advantageously between 20° C. and 200° C., and more advantageously between 20° C. and 150° C.
Mention may in particular be made of 2,2,2-trifluoroethyl difluoromethyl ether (HFE-245mf), 1,1,1,2,2-pentafluoroethyl methyl ether (HFE-245mc), 1,1,2,2-tetrafluoroethyl methyl ether (HFE-245pc), 1,1,2,3,3,3-hexafluoropropyl methyl ether (HFE-356mec) or 1,1,1,2,2,2-hexafluorodiethyl methyl ether (HFE-356mff).
The hydrofluorinated ether, such as heptafluoropropyl methyl ether (HFE-7000), nonafluorobutyl methyl ether/nonafluoroisobutyl methyl ether (HFE-7100), nonafluorobutyl ethyl ether (HFE-7200), decafluoro-3-methoxy-4-(trifluoromethyl)pentane (HFE-7300), 2-trifluoromethyl-3-ethoxydodecafluorohexane (HFE-7500) and the mixture of perfluoroisobutyl ethyl ether and perfluorobutyl ethyl ether (20-80% by weight) (HFE-8200), may be an advantage.
The polyols, such as ethylene glycols RO(CH2CH2O)nR′ with n between 1 and 3 and R and R′, which are identical or different, each representing a hydrogen atom or an alkyl group of 1 to 5 carbon atoms, may be suitable.
The amount of compound C1 to be added can represent from 5 to 95% by weight, with respect to the HFO-1234yf/HF mixture and preferably from 10 to 80% by weight, with respect to the HFO-1234yf/HF mixture.
The composition to be separated is preferably cooled to a temperature of between −20 and 40° C., and advantageously to a temperature of between −5 and 35° C. The cooling temperature depends both on the nature and on the amount of the compound C1 to be added. Thus, in the case of the addition a small amount of HCC-240db, the temperature of the cooling stage is preferably within the vicinity of 0° C., whereas it can reach ambient temperature (that is to say, 25° C.) in the presence of a greater amount of addition of the compound C1.
The pressure at which this cooling stage is carried out is between 0 and 40 bar, preferably between 0.3 and 25 bar and advantageously in the vicinity of the pressure of the reaction stage.
In addition to the HFO-1234yf and the HF, the composition to be separated can comprise organic impurities, such as HCFO-1233xf, and 1,1,1,2,2-pentafluoropropane (HFC-245cb).
These impurities are generally byproducts from the reaction stage.
Alternatively, the separation of HF and 2,3,3,3-tetrafluoropropene may be carried out by decantation. The decantation is carried out at a low temperature (below −5° C.) in order to get a better separation: less HF in the organic phase and less organic (especially HFO-1234yf) in the HF phase.
Furthermore, in order to remove residual amounts of HF which may be present in the 2,3,3,3-tetrafluoropropene stream after the separation step detailed above, the 2,3,3,3-tetrafluoropropene stream may be contacted with water and submitted to a drying step. The 2,2,2,3-tetrafluoropropene stream containing unrecoverable amounts of HF is treated with water in a first HF absorber to remove the bulk of the acid. The stream is then treated in a neutralization scrubber with circulating weak caustic solution (for example 20% NaOH or KOH) to remove further amounts of acid. The acid free stream is then optionally cooled to selectively condense water to reduce the amount of moisture followed by a drying step to remove traces of water. The drying step is carried out using solid products such as calcium sulfate, sodium sulfate, magnesium sulfate, calcium chloride, potassium carbonate, silica gel or molecular sieve (zeolite) such as siliporite.
In embodiments, 2,3,3,3-tetrafluoro-1-propene (HFO-1234yf) may be purified in a purification process in which 2,3,3,3-tetrafluoro-1-propene, comprising impurities based on halogen compounds, is fed to a distillation column to remove “light” organic impurities, i.e. organic impurities having boiling points between the atmospheric boiling point of HCl and HFO-1234yf. The “light” organic impurity may include an organic compound having a boiling point of from −84. to −35° C. The “light” organic impurity may be at least one selected from the group consisting of trifluoromethane (HFC-23), 1,1,1,2,2,2-hexafluoroethane (CFC-116), difluoromethane (HFC-32), 1,1,1,2,2-pentafluoroethane (HFC-125), 3,3,3-trifluoropropyne, 1,1,1-trifluoroethane (HFC-143a), 1-chloro-1,1,2,2,2-pentafluoroethane (CFC-115).
In embodiments of the present process, the 2,3,3,3-tetrafluoropropene stream is subjected to extractive distillation. A 2,3,3,3-tetrafluoropropene stream may be purified as disclosed in FR1563165, FR1563166, FR1563167, FR1563168, FR1563169 incorporated herein by reference.
The 2,3,3,3-tetrafluoropropene stream from one or more impurities contained therein by providing a purification process comprising the steps of:
Said extractive agent may be a solvent selected from the group consisting of hydrocarbon, hydrohalocarbon, alcohol, ketone, amine, ester, ether, aldehyde, nitrile, carbonate, thioalkyl, amide and heterocycle. Said extractive agent may have a boiling point ranging from 10 to 150° C. The extractive agent may be selected with respect to one of said one or more impurities to be removed. Said extractive agent may have a separation factor S1,2 equal to or greater than 1,1, said separation factor being determined by the formula S1,2=(γ1,S*P1)/(γ2,S*P2) wherein
Said extractive agent may also have a absorption capacity C2,S greater than or equal to 0.20, said absorption capacity being determined by the formula C2,S=1/(γ2,S) wherein γ2,S represent the activity coefficient of said one of said one or more impurities to be removed in the extractive agent at infinite dilution; advantageously the absorption capacity may be greater than or equal to 0.40, preferably greater than or equal to 0.60, more preferably greater than or equal to 0.80, in particular greater than or equal to 1.0. Said one or more impurities may be, for example, 3,3,3-trifluoropropene (HFO-1243zf), trans-1,3,3,3-tetrafluoro-1-propene (HFO-1234ze-E), chloromethane (HCC-40), 1,1-difluoroethane (HFC-152a), chloropentafluoroethane (CFC-115), 1,1,1,2-tetrafluoroethane (HFC-134a) or trans-1,2,3,3,3-pentafluoropropene (HFO-1225ye-E). For example, when the impurity to be removed is 3,3,3-trifluoropropene (HFO-1243zf), the extractive agent may be selected from 2-methoxy-1-propene, 1,2-epoxypropane, ethoxy-ethene, dimethoxymethane, methylacetate, isobutanal, isopropylformate, ethylacetate, butanone, n-propylformate, 1,2-dimethoxyethane, isopropylacetate, 2-methylbutanal, ethyl propionate, 1,2-dimethoxypropane, dioxane, 3-pentanone, 2-pentanone, trimethoxymethane, 1,3-dioxane, 3,3-dimethyl-2-butanone, 4-methyl-2-pentanone, diethylcarbonate, n-butylacetate, 2-hexanone, 5-hexen-2-one, 1-ethoxy-2-propanol, hexanal, 2-(dimethylamino)-ethanol, 2-methylpyrazine, 1-methylpiperazine, valeronitrile, 4-methyl-2-hexanone, 1-methoxy-2-acetoxypropane, 2,6-dimethylmorpholine, methylhexanoate, 1-propoxy-2-propanol; advantageously from ethoxy-ethene, dimethoxymethane, methylacetate, isobutanal, isopropylformate, ethylacetate, butanone, 1,2-dimethoxyethane, isopropylacetate, dioxane, 3-pentanone, 2-pentanone, trimethoxymethane, 1,3-dioxane, 3,3-dimethyl-2-butanone, 4-methyl-2-pentanone, diethylcarbonate, n-butylacetate, 1-ethoxy-2-propanol, hexanal; preferably from dimethoxymethane, butanone, isopropylacetate, dioxane, trimethoxymethane, 1,3-dioxane, n-butylacetate, 1-ethoxy-2-propanol, hexanal; in particular from dimethoxymethane, isopropylacetate, dioxane, trimethoxymethane, 1,3-dioxane, n-butylacetate, 1-ethoxy-2-propanol, hexanal.
For example, when the impurity to be removed is trans-1,3,3,3-tetrafluoro-1-propene (HFO-1234ze-E), the extractive agent may be selected from ethylamine, isopropylamine, diethylether, ethoxy-ethene, dimethoxymethane, n-propylamine, methyl-t-butylether, diethylamine, propanone, methylacetate, isobutanal, tetrahydrofurane, isopropylformate, diisopropylether, 2-ethoxy-2-methyl-propane, ethylacetate, butanone, diethoxymethane, isopropylacetate, 3-pentylamine, 2-methoxyethanamine, tert-butylacetate, dioxane, 3-pentanone, 1,1-diethoxyethane, 2-pentanone, trimethoxymethane, n-pentylamine, 1,3-dioxane, 3,3-dimethyl-2-butanone, sec-butylacetate, 4-methyl-2-pentanone, 1,2-diaminoethane, 1-methoxy2-propanol, diethylcarbonate, n-butylacetate, 1-ethoxy-2-propanol, hexanal; advantageously from ethylamine, isopropylamine, diethylether, dimethoxymethane, n-propylamine, diethylamine, diisopropylether, 2-ethoxy-2-methyl-propane, butanone, diethoxymethane, isopropylacetate, 3-pentylamine, 2-methoxyethanamine, tert-butylacetate, dioxane, trimethoxymethane, n-pentylamine, 1,3-dioxane, sec-butylacetate, 1,2-diaminoethane, 1-methoxy-2-propanol, n-butylacetate, 1-ethoxy-2-propanol, hexanal; preferably from éthylamine, isopropylamine, diethylether, dimethoxymethane, n-propylamine, diethylamine, diisopropylether, 2-ethoxy-2-methyl-propane, diethoxymethane, isopropylacetate, 3-pentylamine, 2-methoxyethanamine, tert-butylacetate, dioxane, trimethoxymethane, n-pentylamine, 1,3-dioxane, sec-butylacetate, 1,2-diaminoethane, 1-methoxyl-propanol, n-butylacetate, 1-ethoxy-2-propanol, hexanal.
For example, when the impurity to be removed is chloromethane (HCC-40), the extractive agent may be selected from methylformate, 2-methoxy-1-propene, ethoxy-ethene, propanone, methylacetate, isobutanal, isopropylformate, ethylacetate, butanone, n-propylformate, 1,2-dimethoxyethane, isopropylacetate, 1-methoxy-2-propanamine, 2-methoxyethanamine, 2-methylbutanal, tert-butylacetate, ethylpropionate, dioxane, 3-pentanone, 2-pentanone, 2-methoxy-1propanamine, trimethoxymethane, 1,3-dioxane, 3,3-dimethyl-2-butanone, 2-ethoxyethanamine, sec-butylacetate, n-methyl-1,2-ethanediamine, 4-methyl-2-pentanone, 1,2-diaminoethane, butyronitrile, 1-methoxyl-propanol, 1,2-propanediamine, 2,6-dimethyl-5-heptenal, 1-(dimethylamino)-2-propanol, diethylcarbonate, n-butylacetate, 2-hexanone, n-ethylethylenediamine, 5-hexen-2-one, 2-methyl pyridine, 2-methoxyl-propanol, 1-ethoxy-2-propanol, hexanal, 2-(dimethylamino)-ethanol, 2-methylpyrazine, 2-ethoxy-1-propanol, 1,3-propanediamine, valeronitrile, 2,6-dimethylpyridine, 4-methyl-2-hexanone, 1-methoxy-2-acetoxypropane, 4-methyl pyridine, 2,6-dimethylmorpholine, methylhexanoate, 2-propoxyethanol, 1-propoxy-2-propanol; advantageously from methylformate, ethoxy-ethene, propanone, methylacetate, isobutanal, isopropylformate, ethylacetate, butanone, isopropylacetate, 2-methoxyethanamine, tert-butylacetate, dioxane, 3-pentanone, 2-pentanone, 1,3-dioxane, 3,3-dimethyl-2-butanone, sec-butylacetate, 4-methyl-2-pentanone, 1,2-diaminoethane, 1-methoxyl-propanol, 1,2-propanediamine, diethylcarbonate, n-butylacetate, 2-methoxyl-propanol, 1-ethoxy-2-propanol, hexanal; preferably from methylformate, propanone, butanone, isopropylacetate, 2-methoxyethanamine, tert-butylacetate, dioxane, 1,3-dioxane, sec-butylacetate, 1,2-diaminoethane, 1-methoxy2-propanol, 1,2-propanediamine, n-butylacetate, 2-methoxyl-propanol, 1-ethoxy-2-propanol, hexanal; in particular from methylformate, isopropylacetate, 2-methoxyethanamine, tert-butylacetate, dioxane, 1,3-dioxane, sec-butylacetate, 1,2-diaminoethane, 1-methoxy2-propanol, 1,2-propanediamine, n-butylacetate, 2-methoxyl-propanol, 1-ethoxy-2-propanol, hexanal.
Alternatively, 2,3,3,3-tetrafluoropropene may be separated from chloromethane with an extractive agent selected from pentane, hexane, heptane, octane, nonane, decane, undecane, dodecane, dichloromethane, trichloromethane, perchloromethane, 1,2-dichloropropane, perchloroethylene, methanol, ethanol, propanol, butanol, pentanol, 1,3-dioxolane, tetrahydrofuran, acetonitrile, acetone, methyl ethyl ketone, diethyl ketone, methyl isobutyl ketone, dimethyl carbonate, dimethylformamide, dimethylacetamide, N-methylpyrrolidone, N-formylmorpholine, γ-butyrolactone, and dimethylsulfoxide.
For example, 2,3,3,3-tetrafluoropropene may also be separated from 1,1,1,2-tetrafluoroethane by using a extractive solvent selected from the group consisting of an alcohol, a ketone, an ester, an amide, a hydrofluoroether having 2 to 4 carbon atoms, a sulfoxide, a nitrile, and dichloropropane, preferably the extractive solvent may be selected from methanol, ethanol, propanol, butanol, acetone, methyl ethyl ketone, diethyl ketone, γ-butyrolactone, dimethylformamide, dimethylacetamide, N-methylpyrrolidone, 1,1,2,2-tetrafluoroethyl-2, 2, 2-trifluoroethyl ether, dimethyl sulfoxide and acetonitrile.
In embodiments, 2,3,3,3-tetrafluoro-1-propene (HFO-1234yf) may be purified in a purification process in which 2,3,3,3-tetrafluoro-1-propene, comprising impurities based on halogen compounds, is brought into contact with an adsorbent, preferably molecular sieves and advantageously molecular sieves having a pore opening with an average diameter between 5 and 11 Å, preferably between 5 and 9 Å.
It has been found that the impurities based on halogenated compounds present in an HFO-1234yf can be removed (partially or totally) by placing a crude HFO-1234yf in contact with an adsorbent.
Molecular sieves, also known as synthetic zeolites, are chemical compounds widely used in the industry as adsorbents, especially for drying gases or liquids. They are metal aluminosilicates that have a three-dimensional crystal structure formed from an assembly of tetrahedra. These tetrahedra are formed by four oxygen atoms that occupy the apices, and which surround either a silicon atom or an aluminum atom placed at the center. These structures generally contain cations to make the system electrically neutral, such as those derived from sodium, potassium or calcium.
The molecular sieves that are suitable for use are preferably those of the type A and of the type X and advantageously those of the type X.
In the case of molecular sieves “of the type A”, the tetrahedra are assembled such that they compose a truncated octahedron. These octahedra are themselves arranged in a simple cubic crystal structure, forming a network whose cavities have an approximate diameter of 11.5. These cavities are accessible via apertures, or pores, that can be partially blocked with cations. When these cations are derived from sodium, these cavities have an aperture diameter of 4.1 angstrom, which then gives a “4 A” molecular sieve. The crystal structure of such a sieve may be represented by the following chemical formula:
Na12[(AlO2)12(SiO2)12].XH2O in which x, which represents the number of water molecules belonging to the structure (water of crystallization), may be up to 27, which represents 28.5% by weight of the anhydrous zeolite.
After removing the water of crystallization by heating at a temperature from about 500 to 700° C., the cavities of these substances are available for the selective adsorption of various gases or liquids. Thus, the pores of the various types of zeolite allow the passage and adsorption in the corresponding cavities only of molecules whose effective diameter is less than or equal to the effective pore diameter. In the case of drying gases or liquids, it is thus water molecules that are retained by selective adsorption in the cavities mentioned previously, the substance to be dried being itself not or only sparingly adsorbed.
The size of the apertures (or pores) may, moreover, be modified according to the different types of molecular sieve. Thus, by exchanging a large proportion of the sodium ions of a 4 A molecular sieve with potassium ions, the 3 A molecular sieve is obtained, the pores of which have a diameter of about 3 angstrom. The 5 A molecular sieve is prepared by replacing the sodium ions with calcium ions, the effective pore diameter then being about 5 angstrom.
The elementary cell of zeolite X is a tetrahedron whose apices are occupied by polyhedra of the same type as those present in zeolite A, each being connected to four other polyhedra by virtue of an octahedral substructure, formed by a double ring containing eight oxygen atoms. The center of each edge is always occupied by an oxygen atom, whereas the silicon and aluminum atoms occupy the various apices of the polyhedra. The empirical formula is of the structure Na88Al88Si104O384.220H2O.
This process is suitable for the purification of a crude HFO-1234yf with a purity of at least 85% by weight, preferably greater than 90% by weight and advantageously greater than 95% by weight.
The crude HFO-1234yf subjected to the purification step may originate directly from the effluent obtained from the manufacturing step, after optional separation such as decantation or distillation.
The impurities based on saturated halogenated compounds present in the HFO-1234yf are especially HFC-245eb (CF3-CHF—CH2F), HFC-245cb (CF3-CF2-CH3), HFC-236ea (CF3-CHF—CHF2), 1,1,1,2-tetrafluoro-3-chloropropane and tetrafluoropropane. The impurities based on unsaturated halogenated compounds are especially fluoropropenes, such as 1,1,1,2,3-pentafluoropropene, 1,1,1,3,3-pentafluoropropene and 1,1,1-trifluoropropene.
The placing in contact with the adsorbent to purify the crude HFO-1234yf may be performed in the gaseous phase or in the liquid phase at a temperature of between −20° C. and +80° C.; and preferably between +10° C. and +40° C., and at a pressure from 100 to 2200 kPa, preferably at atmospheric pressure.
For the gaseous-phase treatment, a flow rate corresponding to a throughput of between 10 and 40 g/h of crude HFO-1234yf may be used for an amount of adsorbent of between 10 and 50 g.
When step 4g) relates in particular to the purification of 1,1,1,2,2-pentafluoropropane (HFC-245cb), the purification may be an extractive distillation as described above. The purification of 1,1,1,2,2-pentafluoropropane may be carried out by extractive distillation as disclosed in the patent application FR1563163 incorporated herein by reference.
This step may be carried out to separate 1,1,1,2,2-pentafluoropropane (HFC-245cb) from any of the following impurities 1,1-difluoroethane (HFC-152a), 1,1,1,2-tetrafluoroethane (HFC-134a), trans-1,3,3,3-tetrafluoro-1-propene (HFO-1234ze-E), cis-1,3,3,3-tetrafluoro-1-propene (HFO-1234ze-Z), trans-1,2,3,3,3-pentafluoropropene (HFO-1225ye-E), cis-1,2,3,3,3-pentafluoropropene (HFO-1225ye-Z), or 3,3,3-trifluoropropene (HFO-1243zf), if any of these compounds are present in a 1,1,1,2,2-pentafluoropropane stream to be purified.
The purification may be carried out by:
In particular, the 1,1,1,2,2-pentafluoropropane stream may be separated from trans-1,3,3,3-tetrafluoropropene with an extractive agent selected from ethylamine, isopropylamine, n-propylamine, diethylamine, propanone, tetrahydrofurane, ethylacetate, butanone, 3-pentylamine, 2-methoxyethanamine, dioxane, 3-pentanone, 2-pentanone, n-pentylamine, 1,3-dioxane, 1,2-diaminoethane, 1,2-propanediamine, 2-methoxyethanol, n-butylacetate, 1-ethoxy-2-propanol.
When step 4g) relates in particular to the purification of 2-chloro-3,3,3-trifluoropropene, the purification may be an extractive distillation as described above. The purification of 2-chloro-3,3,3-trifluoropropene may be carried out by extractive distillation as disclosed in the patent application FR1563164 incorporated herein by reference. Preferably, this step may be carried out to separate 2-chloro-3,3,3-trifluoropropene from any of the following impurities E-1-chloro-3,3,3-trifluoro-1-propene (HCFO-1233zdE), 1,1,1,3,3-pentafluoropropane (HFC-245fa) or 1,1,1,3,3,3-hexafluoropropane (HFC-236fa), if any of these compounds are present in a 1,1,1,2,2-pentafluoropropane stream to be purified.
For example, a 2-chloro-3,3,3-trifluoropropene stream may be purified from 1,1,1,3,3-pentafluoropropane (HFC-245fa) by using an extractive agent selected from ethanedial, propanone, methylacetate, methylglyoxal, ethylacetate, butanone, propionitrile, dioxane, trimethoxymethane, 1,3-dioxane, 1,3,5-trioxane, 1,2-diaminoethane, 1-methoxyl-propanol, diethylcarbonate, 2-methoxyl-propanol, 1-methoxy-2-acetoxypropane, dimethylformamide, 3-methoxy-1-butanol, diacetone alcohol, methylacetoacetate, n,n-dimethylpropanamide, dimethylmalonate, diethylsulfoxide, 2-(2-methoxyethoxy)ethanol, trimethylphosphate, diethylmalonate; preferably an extractive agent selected from propanone, methylacetate, ethylacetate, butanone, dioxane, trimethoxymethane, 1,3-dioxane, 1,3,5-trioxane, 1,2-diaminoethane and 1-methoxy-2-propanol.
For example, a 2-chloro-3,3,3-trifluoropropene stream may be purified from E-1-chloro-3,3,3-trifluoro-1-propene by using an extractive agent selected from isopropylmethylamine, methyl-t-butylether, diethylamine, propanone, methylacetate, 2-butanamine, n-methylpropylamine, tetrahydrofurane, 1-butylamine, ethylacetate, butanone, n-propylformate, dimethoxypropane, diisopropylamine, 1,2-dimethoxyethane, 3-methyl-2-butanamine, diethoxymethane, isopropylacetate, 3-pentylamine, n-methylbutylamine, 1-methoxy-2-propanamine, 2-methoxyethanamine, tert-butylacetate, ethylpropionate, 1,2-dimethoxypropane, dioxane, 3-pentanone, 1,1-diethoxyethane, 2-pentanone, 2-methoxy-1propanamine, trimethoxymethane, n-pentylamine, 3,3-dimethyl-2-butanone, 1,3-dioxane, piperidine, 2-ethoxyethanamine, sec-butylacetate, n-methyl-1,2-ethanediamine, 2,2-diethoxypropane, 1,2-diaminoethane, 1-methoxy2-propanol, 1,2-propanediamine, 2,6-dimethyl-5-heptenal, 1-(dimethylamino)-2-propanol, 3-methyl-3-pentanol, 2-ethylbutylamine, diethylcarbonate, n-butylacetate, 2-hexanone, n-ethylethylenediamine, 2-methoxyl-propanol, 1-ethoxy-2-propanol, 4-methyl-2-hexanamine, hexylamine, methoxycyclohexane, 2-(dimethylamino)-ethanol, cyclohexylamine, n-ethyl-2-dimethylaminoethylamine, ethoxyethanol, 2-ethoxy-1-propanol, 1-methylpiperazine, 1,3-propanediamine, 2-heptanamine, n,n-diethylethylenediamine, 4-methyl-2-hexanone, 1,1,1-triethoxyethane, 1-methoxy-2-acetoxypropane, 4-methylpyridine, n,n′-diethyl-1,2-ethanediamine, 2,6-dimethylmorpholine, methylhexanoate, 2-propoxyethanol, 1-propoxy-2-propanol, 2-heptanone, dimethylformamide, 2-isopropoxyethanol, 2-methylpiperazine, cyclohexanone, 1-heptanamine, 2-ethoxyethanolacetate, 1,4-butanediamine, 2,4-dimethyl pyridine, 2-methoxy-3-methyl pyrazine, 4-methoxy-4-methyl-pentan-2-one, 3-ethoxy-1-propanol, 3-methoxy-1-butanol, diglyme, 2-(diethylamino)-ethanol, 2,2-diethoxyethanamine, 2-methoxy-n-(2-methoxyethyl)ethanamine, 2-(ethylamino)ethanol, 3-octanone, diacetone alcohol, diethylaminopropylamine, 2-ethylhexylamine, 1-butoxy-2-propanol, 2-butoxyethanol, 2-octanone, methylheptanoate, triethylenediamine, n,n-dimethylpropanamide, 2-propanol-1-methoxy-propanoate, 1,5-pentanediamine, cycloheptanone, 3,4-dimethylpyridine, 1-octanamine, benzylmethylamine, 1,1,3,3-tetramethoxypropane, dihexylphthalate, diethylpropanolamine, 2-butoxyethanolacetate, diethylsulfoxide, 2-(2-methoxyethoxy)ethanol, 4-methylbenzenemethanamine, diethyleneglycolmonoethylether, 2-propylcyclohexanone, trimethylphosphate, 2-methyl-2,4-pentanediol, methyl benzoate, diethylmalonate, 2-methoxypyrimidine; preferably an extractive agent selected from diethylamine, propanone, methylacetate, tetrahydrofurane, ethylacetate, butanone, diethoxymethane, isopropylacetate, tert-butylacetate, dioxane, 3-pentanone, 1,1-diethoxyethane, 2-pentanone, n-pentylamine, 1,3-dioxane, sec-butylacetate, 1,2-diaminoethane, 1-methoxyl-propanol, n-butylacetate, 1-ethoxy-2-propanol.
In the following specific processes for the fluorination of the 1,1,1,2,3-pentachloropropane (HCC-240db) feedstock (step 4) already known in the state of the art, which may be used at least partly as fluorination step 4 of the present process are summarized by way of example below. Patent application WO 2013/088195 of the present applicant, which is incorporated herein by reference. describes a catalytic fluorination of a 1,1,1,2,3-pentachloropropane (HCC-240db) feedstock with HF in the presence of catalyst preferably in a vapor phase to produce a reaction mixture comprising HCl, HF, 1,1,1,2,2-pentafluoropropane (HFC-245cb), 2-chloro-3,3,3-trifluoropropene (HCFO-1233xf) and 2,3,3,3-tetrafluoropropene (HFO-1234yf).
The first reaction step can be performed in a single reactor. The effluent stream exiting the reactor may optionally comprise additional components such as 1,1,1,2,2-pentafluoropropane (HFC-245cb) and unreacted HF.
The product stream of the first step (a) is then sent to a separation step (b), preferably distillation, to give a first stream comprising HCl and HFO-1234yf and a second stream comprising HF, 2-chloro-3,3,3-trifluoropropene and optionally 1,1,1,2,2-pentafluoropropane. The second stream is then fed in a second reactor optionally with fresh HF in conditions sufficient to give a product stream comprising HFO-1234yf, HFC-245cb, together with unreacted HCFO-1233xf and HF. This product stream is sent directly to step (a).
Step 4-a) of the process involves contacting fresh HCC-240db as obtained in step 3 of the process of the invention and/or the reaction products from step 4-c) with HF in the reaction zone in the presence of a catalyst, preferably in the gas phase, under conditions sufficient to give fluorination products comprising mainly HCFO-1233xf and HFO-1234yf.
Typically, the step 4-a) is carried out with a molar ratio HF:Organics, preferably the Organics are the starting materials, more preferably HCC-240db, in particular HCC-240db as obtained in step 3 of the present process, from 4:1 to 100:1, preferably 5:1 to 50:1. Typically, the process is carried out at a pressure from 0.1 to 50 bar absolute, preferably 0.3 to 15 bar absolute. Typically, the process is carried out at a temperature of from 100 to 500° C., preferably from 200 to 450° C. Contact times (catalyst volume divided by the total flow rate of reactants and co-feeds, adjusted to the operating pressure and temperature) are typically from 1 to 50 sec, preferably from 2 to 40 sec.
An oxygen co-feed may be used to extend the catalyst lifetime, typically the molar ratio of oxygen/Organics is from 0.005 to 2, preferably 0.01 to 1.5. The oxygen can be introduced as an oxygen-containing gas such as air, pure oxygen, or an oxygen/nitrogen mixture. A chlorine co-feed may also be used instead of the oxygen co-feed (with the same operating conditions). Chlorine can be introduced as a chlorine-containing gas such as pure chlorine, or a chlorine/nitrogen mixture.
The catalyst is for example a catalyst based on a metal including a transition metal oxide or a derivative or halide or oxyhalide such a metal as described above.
The product stream of step 4-a) comprising HCl, 2-chloro-3,3,3-trifluoropropene, 2,3,3,3-tetrafluoropropene, unreacted HF, and optionally 1,1,1,2,2-pentafluoropropane enters a separation unit, for example a distillation column, to give a first stream comprising HCl and 2,3,3,3-tetrafluoropropene and a second stream comprising HF, 2-chloro-3,3,3-trifluoropropene and optionally 1,1,1,2,2-pentafluoropropane. Step 4-b) can be performed at a temperature preferably from −90 to 150° C. and more preferably from −85 to 100° C., and at a pressure preferably from 0.1 to 50 bar abs and more preferably from 0.3 to 5 bar abs.
The first stream leaves the reaction system and may enter an acid production unit to produce a stream comprising hydrochloric acid and a stream comprising HFO-1234yf.
HFO-1234yf and intermediate products are readily recovered by any means known in the art, such as by scrubbing, washing, extraction, decantation and preferably distillation. Any stream can also be further purified by distillation techniques.
Step 4-c) is a fluorination reaction, preferably gas phase, of the second stream of step 4-b) with HF in the presence of a catalyst and it comprises mainly fluorination of 2-chloro-3,3,3-trifluoro-1-propene obtained in step 4-a) in 2,3,3,3-tetrafluoro-1-propene, the final product.
Step 4-c) can be carried out in a single or multiple gas-phase reactor. This step of the process, as well as the entire process, is preferably run continuously. This step involves mainly contacting HCFO-1233xf with HF in the reaction zone in a gas phase, under conditions sufficient to convert the HCFO-1233xf to fluorination products comprising HFO-1234yf and HFC-245cb. Such conditions are given below. In addition to the fluorinated products, unreacted HCFO-1233xf, unreacted HF and other co-produced underfluorinated intermediates which may be present in minor amounts are sent directly to step 4-a).
Typically, this step is carried out with a molar ratio HF:Organics from 4:1 to 100:1, more preferably 5:1 to 50:1. Typically, this step is carried out at a pressure from 0.1 to 50 bars, preferably 0.3 to 15 bars absolute. Typically, this step is carried out at a temperature of from 100 to 500° C., preferably from 200 to 450° C.
Contact times (catalyst volume divided by the total flow rate of reactants and co-feeds, adjusted to the operating pressure and temperature) are typically from 1 to 100 sec, preferably from 5 to 50 sec.
An oxygen co-feed may be used to extend the catalyst lifetime, typically the molar ratio of oxygen/Organics is from 0.005 to 2, preferably 0.01 to 1.5. The oxygen can be introduced as an oxygen-containing gas such as air, pure oxygen, or an oxygen/nitrogen mixture. A chlorine co-feed may also be used in lieu of the oxygen co-feed (with the same operating conditions).
Chlorine can be introduced as a chlorine-containing gas such as pure chlorine, or a chlorine/nitrogen mixture.
The catalyst described above can be used in this step. It can be similar to the one used in step 4-a) or different.
Reaction steps 4-a) and 4-c) are implemented in a dedicated reactor for reactions involving halogens. Such reactors are known to those skilled in the art, and can include linings as mentioned above. The reactor may also include means of heat exchange, if necessary.
Besides advantages described above, the reaction step 4-c) which is a critical step can be performed in the absence of the huge amount of HCl generated in the first step and also in some embodiment such as when the reactor of step 4-c) is placed above that of the reactor of step (a), loading and unloading of the catalyst is easier. Moreover, since unreacted HCFO-1233xf coming from step 4-c) also reacts in step 4-a), the yield of HFO-1234yf based on pentachloropropane is higher.
The present invention can be practiced in a compact plant since only one separation cycle is needed and is also low energy consuming.
Patent application WO 2012/052797 of the present applicant, which is incorporated herein by reference describes a process for preparing 2,3,3,3-tetrafluoropropene, comprising the following steps of catalytic reaction of 1,1,1,2,3-pentachloropropane with HF into product 2-chloro-3,3,3-trifluoropropene; and catalytic reaction of the thus-obtained 2-chloro-3,3,3-trifluoropropene into 2,3,3,3-tetrafluoropropene. The second stage of this method of preparation of the HFO-1234yf is a fluorination reaction of the 2-chloro-3,3,3-trifluoro-1-propene (HCFO-1233xf) obtained at the previous stage in 2,3,3,3-tetrafluoro-1-propene, the final product. The two stages can be implemented continuously or in a discontinuous way, with intermediate storage of the HCFO-1233xf. This second stage can comprise direct fluorination in the presence of HF, on a catalyst, in gas phase. This gas phase reaction is carried out in the presence of a fluorination catalyst. The reaction may be carried out in a single gas-phase reactor. The temperatures, pressures and contact times are easily determined by the skilled worker. Typical conditions are given above.
The catalyst can be present in any suitable form, such as fixed or fluidized bed, preferably in a fixed bed. The direction of flow may be downward or upward.
This step can be carried out in an apparatus as depicted in
The pressure need not be the same in steps 4-a) and 4-c). Preferably the pressure in the second reactor is lower than the pressure in the first one so to gain in terms of pumps in the unit. The other process conditions need not be the same in steps 4-a) and 4-b) either.
Liquid-Phase Reaction of HCC-240db into HCFO-1233xf
HCC-240db can be fluorinated in liquid phase into HCFO-1233xf, and that process conditions can be selected so as to achieve the reaction with a substantial selectivity into the desired product.
In a preferred embodiment, the liquid phase process is carried out in an organic phase. Using an organic phase rather than an HF phase favors the reaction into HCFO-1233xf. The prior art reported above disclose reaction mixtures comprising a substantial part of HF, hence an acidic phase. In an acidic phase, only saturated products are produced. The applicant has found that, there exist conditions that allow fluorination into HCFO-1233xf. Notably, when the reaction is carried out in an organic phase (comprised of the HCC-240db starting material and/or solvent), then HCFO-1233xf can be formed. When HF is added to an initial medium, it will not remain in the medium since it reacts and the amount of HF (or concentration) will be very low, compared to the other products.
The term “organic phase” can thus be defined as referring to a reaction phase comprising the catalyst and the starting material and possibly a solvent if used, but substantially free of HF. Especially the process carried out in an “organic phase” refers to the process in which the initial load does not comprise any HF, in contrast with the prior art.
Because of particular operating conditions, gaseous HCFO-1233xf can be removed from the reactor under gaseous phase, keeping polymerization reactions at a low level.
The liquid phase fluorination of HCC-240db into HCFO-1233xf is carried out in the presence of a catalyst.
The reaction can be implemented in a liquid solvent medium, the reaction zone being either loaded at the beginning with a starting amount of organic (the starting material) and/or the necessary quantity of solvent, or fed continuously with this quantity of solvent (possibly preliminary mixed with the raw material). When carried out with solvent, it is preferred that the solvent be loaded at the beginning; injections with a view of adjusting the quantity of solvent may however be carried out if necessary.
The reaction conditions (notably pressure) are such that the reactants are liquid. According to an embodiment the reactants are liquid while the reaction product is gaseous. The fact that the reaction products are gaseous allows their recovery in a gaseous phase at the exit of the reaction zone. The intermediate product, especially the HCFC-242 compound (trichlorodifluoropropane), is preferably liquid under the reaction conditions, even though it can be stripped away in the gaseous flow.
This stage is in particular implemented under a pressure higher than 2 bar. Advantageously, the pressure lies between 4 and 50 bar, in particular between 5 and 25 bar.
For example, the reaction may be implemented at a temperature ranging between 30° C. and 200° C., preferably between 40° C. and 170° C., advantageously between 50° C. and 150° C.
The molar ratio HF:starting compound lies generally between 0.5:1 and 50:1, preferably between 3:1 and 20:1. Values of about 5:1 can be used with advantage. The amount of HF added will correspond to the stoichiometry of the reaction (here 3), to which one will add the amount of HF that is present in the exiting streams (HF and organics) which are usually azeotropic mixtures.
The other reaction conditions, notably flow rates, can be determined by the skilled person according to common general knowledge, depending on the temperature, pressure, catalyst, reactant ratios, and the like. One shall take care that further fluorination reactions should be avoided so that HCFO-1233xf is the main product obtained (apart intermediate products).
The solvent, if used, is an inert organic solvent under the reaction conditions. Such a solvent will be generally saturated, advantageously in C2 to C6, in order to avoid the reactions of addition. Such solvents can for example be those mentioned in patent application FR2733227. Such solvents have a boiling point (measured at atmospheric pressure), for example higher than 40° C., advantageously higher than 50° C., in particular higher than 60° C. Higher reaction temperatures will imply higher pressures, so that the boiling point of the solvent under the conditions of reaction is higher than the temperature of implementation of the reaction.
One can in particular mention as a solvent the saturated compounds of ethane, propane or butane, substituted by at least two atoms of halogen, chosen among chlorine and fluorine, or a mixture thereof. As an example one can mention 1,2-dichloroethane, 1,2,3-trichloropropane, 1-chloro-1-fluoroethane, 1,1-difluoroethane, 1,1-dichloroethane and 1,3-dichloro-1-fluorobutane, tetrachlorofluoropropane isomers, trichlorodifluoropropane isomers and dichlorotrifluoropropane isomers, 1,1,1,3,3-pentafluorobutane and 1,1,2-trichloro-2,2-difluoroethane, or a mixture thereof. Nitrated solvents like nitromethane or nitrobenzene and sulfones like tetramethylene sulfone (also known as sulfolane) or dimethyl sulfone may also be used. A preferred solvent is the 1,1,2-trichloro-2,2-difluoroethane (HCFC-122). One can also use possibly reactive solvents, in so far as the product of their reaction is a nonreactive solvent. For example, one can also use the precursor of HCFC-122, namely HCFC-121 (1,1,2-trichloro-2-fluoroethane) or perchloroethylene. The solvent can be present in a quantity for a dilution ratio from at least 20%, preferably between 20% and 80%, advantageously between 40% and 60%.
The reaction is catalyzed. The catalysts may be catalysts known by the person skilled in the art of fluorinations in liquid phase.
One can use an acid of Lewis, a catalyst containing a metal halide, in particular containing halide of antimony, tin, tantalum, titanium, metals of transition such as molybdenum, niobium, iron halides, cesium, oxides of metals of transition, halides of metals of the IVb group, halides of metals of the Vb group, a fluorinated chromium halide, a fluorinated chromium oxide or a mixture of both. One can advantageously use metal chlorides and fluorides. Examples of such catalysts include: SbCl5, SbCl3, TiCl4, SnCl4, TaCl5, NbCl5, TiCl4, FeCl3, MoCl6, CsCl, and their corresponding fluorinated derivatives. Pentavalent metal halides are suitable. Advantageously one will use a catalyst containing an ionic liquid. These ionic liquids are particularly interesting for fluorination by HF in liquid phase. One will be able to mention the ionic liquids described in patent applications WO2008/149011 (in particular from page 4, line 1 to page 6 line 15, included by reference) and WO01/81353 in the name of the applicant, as well as the reference “liquid-phase HF Fluorination”, Multiphase Homogeneous Catalysis, Ed. Wiley-VCH, (2002), 535.
One can operate with variable ratios catalyst/organic (including solvent if used), but in general one will prefer that this molar ratio lies between 2 mol % and 90 mol %, preferably between 4 mol % and 80 mol % and more preferably between 6 mol % and 75 mol %.
The starting material can be substantially pure HCC-240db.
A chlorine stream may be used to increase the lifetime of the catalyst, typically in a quantity from 0.05 to 20 mole o, preferably 0.5 to 15 mole % of chlorine per mole of starting compound HCC-240db. Chlorine may be introduced pure or mixed with an inert gas such as nitrogen. The use of an ionic catalyst allows using small quantities of chlorine.
It is also possible that the product of the reaction be stripped using a light gas allowing its drive by mechanical entrainment. Removing gaseous HCFO-1233xf from the liquid phase reactor keep polymerization reactions at a low level (since polymerizable material is in a low amount in the medium) as well as side-reactions (such as addition onto the double bond of the HCFO-1233xf). The addition of a gaseous compound can be advantageous for the reaction, which can be favored for example by the improvement of agitation (bubbling).
This gas can be inert as the nitrogen or helium or the gas can be preferably HCl. When HCl is used, the reaction performs despite the addition into the medium of HCl, which is a reaction product. Advantageously, this added gas is anhydrous hydrochloric acid. The flow of the stripping gas is determined according to the operating conditions. For example, the flow of HCl, compared to the flow of starting product is such that the molar ratio HCl:starting product lies between 0.5:1 and 5:1, advantageously, between 1:1 and 3:1.
The fluorination process in liquid phase according to the invention can be implemented continuously or semi-continuously. According to the preferred embodiment, the process is continuous.
The reactants (starting product and HF) and other compounds used in the reaction (chlorine, anhydrous HCl) can be fed in the reactor at the same place or at different places of the reactor. A preferred embodiment is when the gaseous compounds are injected in the bottom of the reactor, in particular in order to enhance the mechanical stripping and the mixing.
If a recycling is used, one can recycle directly at the inlet of the reactor or on a separate dip pipe.
Gas-Phase Reaction of HCC-240db into HCFO-1233xf.
HCC-240db can be catalytically fluorinated in gas phase into HCFO-1233xf as described above, and HCFO-1233xf may be produced along with HFO-1234yf and optionally HFC-245cb, and separation and recycling of HCFO-1233xf and optionally HFC-245cb into the gas phase reactor is one embodiment.
Hence, also a process is provided comprising the steps of: (i) contacting 1,1,1,2,3-pentachloropropane (HCC-240db) with hydrogen fluoride HF in gas phase in the presence of a fluorination catalyst under conditions sufficient to produce a reaction mixture comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf); (ii) separating the reaction mixture into a first stream comprising 2,3,3,3-tetrafluoropropene (HFO-1234yf) and a second stream comprising 2-chloro-3,3,3-trifluoro-1-propene (HCFO-1233xf); (iii) recycling at least a part of the second stream at least in part back to step (i).
This recycling can take various forms, as is depicted in the following figures.
The reactants can be fed to the reactor at the same location, at different locations, or using staged feeding at staged locations along the reactor. A preferred feeding system is to blow the gaseous reactants at the bottom of the reactor. Recycling can be done at the entry of the reactor or at an intermediate stage of the reactor; preferably at the entry of the reactor. It is also possible to recycle part of the stream exiting the reactor.
Reactions are implemented in a dedicated reactor for reactions involving halogens. Such reactors are known to those skilled in the art. The reactor may also include means of heat exchange, if necessary.
The final product is readily recovered by any means known in the art, such as by scrubbing, washing, extraction, decantation and preferably distillation. It can also be further purified by distillation techniques.
In the following examples 1 to 19 the process for the production of high purity 1,1,1,2,3-pentachloropropane is illustrated.
For the avoidance of doubt, where reference is made to units of pressure (kPa) herein it is the absolute value which is identified. Where values are presented as percentages herein, they are percentages by weight unless otherwise stated. Where the purity of a composition or material is presented by percentage or ppm herein, unless otherwise stated, this is a percentage/ppm by weight.
For clarity, Examples 1 to 7 exemplify or relate to the telomerisation reaction (and subsequent treatment steps) of step 1) of the HCC-240db production process, which is an optional step. Examples 8 to 12 exemplify or relate to the dehydrochlorination reaction (and subsequent treatment steps) of step 2) of the of the HCC-240db production process, which is an optional step. Examples 13 to 19 exemplify or relate to the chlorination reaction (and subsequent treatment steps) of step 3) of the HCC-240db production process.
Abbreviations used:
TeCPa=1,1,1,3-tetrachloropropane
TCPe=1,1,3-trichloropropene
PCPa=1,1,1,2,3-pentachloropropane
TeCM: tetrachloromethane
TeCPna: tetrachloropentane
HCE=hexachloroethane
DCPC=dichloropropanoylchloride
Bu3PO4: Tributylphosphate
Ethylene and carbon tetrachloride were reacted to produce 1,1,1,3-tetrachloropropane in the presence of catalyst which was either i) recovered from a reaction mixture using conventional distillation techniques, or ii) recovered from a reaction mixture using the inventive aqueous treatment step for catalyst described herein. The reaction mixture additionally comprised 1,1,1,3-tetrachloropropane (present in the recycle stream) and tetrachloropentane (a chlorinated alkane impurity commonly formed as a byproduct in the presence of telomerisation reactions between carbon tetrachloride and ethylene).
These test examples show that using the aqueous treatment step to recover catalyst, the performance of the catalyst is significantly higher as compared to catalyst recovered using conventional distillation techniques.
Gas chromatography was used to monitor the progress of the reaction.
A stainless steel autoclave with a volume of 405 ml, equipped with a stirrer, a thermowell for temperature measurement and a sampling tube (with valve) was filled with the reaction mixture described below and closed. Heating was provided by means of an oil bath placed on a magnetic (heating) stirrer. Ethylene was fed by a copper capillary tube from 10 l cylinder placed on weighing scale. The gaseous atmosphere in the autoclave was replaced by ethylene flushing. After pressurizing with ethylene to 5 bar, the autoclave was heated up to 105° C., then the ethylene supply to the autoclave was opened. Ethylene supply was controlled manually for a first ten minutes (to maintain the reaction temperature to 112° C.), and later was maintained at a constant pressure of 9 bar. The reaction was allowed to react defined time period. Than the reactor was cooled and reaction mixture was withdrawn after opening of depressurised reactor.
In the first example, the distillation residue was directly used as a recycled catalyst (Comparative Example 1-1). In the second example, the distillation residue was extracted with 5% hydrochloric acid and a filtered organic fraction was used as a catalyst (Example 1-2).
90.1 g of a distillation residue comprising 63.7% TeCPa, 22.8% TeCPna and 7.49% Bu3PO4 was mixed with 400 g of TeCM. The mixture was then introduced into the autoclave where 5.0 g of iron was added. After flushing with ethylene, the mixture was heated in the autoclave up to 110° C. At this temperature and at a pressure of 9 bar of ethylene, the reaction mixture was allowed to react for 4.5 hours. The first sample was taken after 3 hours. The concentration of residual TeCM at the end of the experiment was 19.7% (33.0% after 3 hours).
90.1 g of a distillation residue comprising 63.7% TeCPa, 22.8% TeCPna and 7.49% Bu3PO4 was extracted with 370 g of 5% HCl. A bottom organic layer was filtered and mixed with 400 g TeCM. The mixture was then introduced into the autoclave where 5.0 g of iron was added. After flushing with ethylene, the mixture was heated in the autoclave up to 110° C. At this temperature and at a pressure of 9 bar of ethylene, the reaction mixture was allowed to react for 4.5 hours. The first sample was taken after 3 hours. The concentration of residual TeCM at the end of the experiment was 5.5% (24.6% after 3 hours).
Comparative Example 1-3 was carried out using identical conditions as those employed in Comparative Example 1-1, except that differing concentrations of tetrachloromethane and tributylphosphate were used.
Examples 1-4 and 1-5 were carried out using identical conditions as those employed in Example 1-2, except that differing concentrations of tetrachloromethane and tributylphosphate were used.
The results of Comparative Example 1-1 and Example 1-2, and Comparative Example 1-3 and Examples 1-4 and 1-5 are shown in the following table. As can be seen, the percentage of tetrachloromethane which was converted to 1,1,1,3-tetrachloropropane is significantly higher in Examples 1-2, 1-4 and 1-5 than in Comparative Examples 1-1 and 1-3 demonstrating that the performance of an aqueous treatment step when recovering the catalyst has a profound positive effect on the system. This is due to the high viability of the catalyst recovered from the distillate residue and also potentially due to the removal of impurities (e.g. oxygenated impurities) from the reaction mixture which otherwise may retard the reaction.
The same stainless steel autoclave as described above for the batch experiments was used as a stirred flow continuous reactor. The reactor was initially filled with approximately 455 g of reaction mixture. After pressurizing with ethene to 5 bar, the autoclave was heated up to 105° C., then the ethylene supply to the autoclave was opened, with continuous feed of the raw material and continuous withdrawal of the reaction mixture started.
Feedstock solution with dissolved catalyst was fed into the autoclave from a stainless steel tank. The tank was placed above the reactor, and thus, a pump was not used for feeding the reactor. Reactor and tank were under an atmosphere of ethene distributed by copper capillaries from the cylinder, with conditions in the cylinder selected to prevent commencement of the reaction. Sampling of the reaction mixture was carried out by sampling tube every five minutes. To monitor the course of the reaction, the container with the feedstock and dissolved catalyst, cylinder of ethene and the withdrawn reaction mixture were weighed. The reaction mixture was always collected for an hour and after that, the collecting vessel is replaced.
Continuous experiments comparing the activity of recycled catalyst (i.e. a distillation residue were conducted with and without performance of an aqueous treatment step. In the first case, the distillation residue was directly used as a recycled catalyst (Comparative Example 1-6). In the latter cases, the reaction mixture, after aqueous treatment of the distillation residue with 5% HCl, was used as a raw material containing recycled catalyst (Examples 1-4 and 1-5).
587.5 g of the distillation residue comprising 63.7% TeCPa, 22.8% TeCPna and 7.49% Bu3PO4 was mixed with 2200 g of TeCM. This mixture comprised 78.7% TeCM, 11.8% TeCPa, 5.8% TeCPna and was used as a feed stream for the continuous arrangement. The reaction vessel constituted an autoclave was filled with reaction mixture and 8 g of fresh iron. The reaction was carried out at 110° C. with a pressure of ethylene of 9 bar. The residence time was 2.7 hours. During the reaction, the amount of reacted TeCM ranged between 75-76%.
587.5 g of the distillation residue comprising 63.7% TeCPa, 22.8% TeCPna 7.49% Bu3PO4 was added dropwise over 1.5 hour into 1001.5 g of boiling 5% HCl. This mixture was then stripped. From the overhead product, an organic phase was collected and an aqueous phase was returned as a reflux. Distillation was terminated after an hour when all of the distillation residue was added. The residue, after stripping, was diluted with 200 g of TeCM and then separated in a separatory funnel. A bottom organic phase was filtered and together with distilled residue was mixed with 2000 g of TeCM. This mixture comprised 81.2% TeCM, 10.8% TeCPa and 5.3% TeCPna. It was used as a feed stream for the continuous arrangement of the experiment. The reaction vessel (autoclave) was filled with the older reaction mixture and 8 g of fresh iron. The reaction was carried out at 110° C. and a pressure of ethene of 9 bar. Residence time was 2.7 hours/flow rate. During the time of the reaction the amount of reacted TeCM ranged between 83-85%.
Comparative Example 1-8 was carried out using identical conditions as those employed in Comparative Example 1-6, except that differing concentrations of tetrachloromethane and tributylphosphate were used.
Example 1-9 was carried out using identical conditions as those employed in Example 1-7, except that differing concentrations of tetrachloromethane and tributylphosphate were used.
High purity 1,1,1,3-tetrachloropropane may be obtained according to step 1) of the process involving an alkylation step (
In the alkylation step shown in
Particulate iron is intermittently fed into the continuously stirred tank reactor 3 using a controlled feed. The ongoing addition of particulate iron is employed because, as the alkylation reaction proceeds, particulate iron dissolves into the reaction mixture. It has been found that optimal results are obtained by maintaining the presence of particulate iron in the reaction mixture, in this example with the addition of 1 to 2% by weight of the reaction mixture in the primary alkylation zone. This results in the reaction mixture extracted from the primary alkylation zone having a dissolved iron content of 0.2 to 0.3% by weight of the reaction mixture. Carbon tetrachloride is fed into the continuously stirred tank reactor 3 via line 12 in liquid form. In the illustrated embodiment, the carbon tetrachloride stream is used to trap gaseous ethene extracted from the reaction mixture. However, the use of carbon tetrachloride in this way is not essential; a feed of fresh carbon tetrachloride as the sole or main source of carbon tetrachloride could be fed into the reactor 3.
Tributyl phosphate/ferric chloride catalyst is also fed into the continuously stirred tank reactor 3 via line 12. The tributyl phosphate present in that stream is partly obtained from the aqueous treatment process illustrated in
In the illustrated embodiment, a single primary alkylation zone is employed, located in the continuously stirred tank reactor 3. Of course, if required, a plurality of primary alkylation zones could be employed, for example in one or more continuously stirred tank reactors, that could be operated in parallel and/or in series.
The primary alkylation zone is operated under superatmospheric pressure (5 to 8 bar gauge) and elevated temperature (105° C. to 110° C.) and for a residence time of 100-120 minutes. These conditions are selected to cause the carbon tetrachloride and ethene to form 1,1,1,3-Tetrachloropropane in an alkylation reaction. However, it has been found that the total conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane is undesirable as this also results in the formation of unwanted impurities. Thus the level of conversion of the carbon tetrachloride to the chlorinated C3-6 alkane of interest is controlled and, in this embodiment, is not permitted to proceed beyond 95% Control of the progress of the alkylation reaction is achieved partly through use of reaction conditions which do not favour the total conversion of carbon tetrachloride to 1,1,1,3-tetrachloropropane, through control of the residence time of the reaction mixture in the continuously stirred tank reactor.
Reaction mixture comprising i) unreacted carbon tetrachloride and ethene, ii) 1,1,1,3-Tetrachloropropane (the chlorinated C3-6 alkane of interest in this embodiment) and iii) tributyl phosphate/iron chloride catalyst is extracted from the primary alkylation zone (the continuously stirred tank reactor 3) and fed into a plug/flow reactor 4 (in which the principal alkylation zone is located).
The reaction mixture is extracted such that particulate iron catalyst present in the primary alkylation zone 3 is not extracted and thus the reaction mixture is substantially free of particulate material. Further, in the illustrated embodiment, no additional catalyst is added into the plug/flow reactor 4, although the plug/flow reactor 4 may provide with a catalyst bed. Additionally, no further ethene is added into the plug/flow reactor 4.
In the illustrated embodiment, the operating pressure in the principal alkylation zone 4 is the same as that in the primary alkylation zone 3. The residence time of the reaction mixture is about 30 minutes, which in the illustrated embodiment was sufficient to result in substantially all of the ethene present being used up in the reaction. Of course, it will be understood that for different reactor types and operating conditions, different resident times may be optimal.
When the determined optimal residence time of the reaction mixture in the principal alkylation zone has been reached, reaction mixture is extracted therefrom via line 5, while being maintained at elevated pressure and temperature, and fed into flash evaporation vessel 6. In this vessel, the extracted reaction mixture is subjected to depressurisation, to atmospheric pressure. This pressure drop causes evaporation of the ethene present in the reaction mixture. The 1,1,1,3-tetrachloropropane-rich mixture, now with substantially no ethene present, is extracted from the flash vessel via line 7 and subjected to the distillation step shown in
The evaporated ethene is extracted from the flash vessel 6 via line 8 and fed through a condenser 9. The ethene is then fed via line 10 into absorption column 11 where it is contacted with a stream of carbon tetrachloride and tributyl phosphate/iron chloride catalyst, recovered from the reaction mixture in the aqueous treatment step shown in
The flow of cooled carbon tetrachloride/catalyst through the absorption column 11 has the effect of trapping the ethene therein. The obtained liquid flow of carbon tetrachloride/catalyst/ethene is then fed back into the continuously stirred tank reactor 3.
As is apparent from
Turning now to
The light ends and tetrachloroethene streams 110.1, 110.3 may be used in the production of carbon tetrachloride, advantageously minimising the production of waste products. This can be achieved through use of a high temperature chlorinolysis process.
The carbon tetrachloride stream 110.2 is recycled back into the continuously stirred tank reactor shown with reference numeral 3 in
A 1,1,1,3-tetrachloropropane-rich mixture which also comprises catalyst is extracted as a residue from boiler 102 via line 103 and is subjected to the catalyst recovery step shown in
In that step, the 1,1,1,3-tetrachloropropane-rich mixture is fed into a batch distillation boiler 204 via line 202, along with a weak (1-5%) hydrochloric acid solution via line 201.
Advantageously, the water present in the acid solution deactivates the catalyst system and protects it from thermal damage. This enables the catalyst system, to be recovered from the 1,1,1,3-tetrachloropropane-rich mixture, and it can be easily reactivated, post-recovery, and reused in the alkylation process without any significant loss in catalytic activity.
The batch distillation boiler is operated at a temperature of about 100° C., to create a gaseous mixture comprising 1,1,1,3-tetrachloropropane and water vapour.
The gaseous mixture produced in the boiler 204, is then subjected to steam distillation (or steam stripping) of crude 1,1,1,3-tetrachloropropane in column 210, which is coupled to the boiler 204. The crude 1,1,1,3-tetrachloropropane is extracted from the distillation column 210 via line 211, condensed with a condenser 212, fed via line 213 to a reflux liquid-liquid separator 214. Water from the gaseous mixture is fed back to the distillation column 210 via line 215, and a portion is taken off via line 216 for a further distillation step, shown in more detail in
The operating temperature of the boiler 204 is then reduced to stop steam stripping, resulting in the condensation of the water vapour present therein. This results in the formation of a biphasic mixture containing an aqueous phase and an organic phase containing the catalyst system, which has not be subjected to steam stripping. To facilitate extraction of the organic phase, a haloalkane extraction agent (in this case, 1,1,1,3-tetrachloropropane) is added to the boiler 204 via line 203 to increase the volume of that phase.
Extraction of the organic phase from the biphasic mixture is achieved by the sequential extraction of the phases from the boiler 204 via line 205. The organic phase is extracted from the boiler 204 via line 205 and is filtered 206. A filter cake is removed from the filter 206 via line 207. The organic phase is extracted via line 208 and, in this embodiment, fed back to the primary alkylation zone, as shown in
The stripped crude 1,1,1,3-tetrachloropropane product is subjected to a further distillation step shown in
The chlorinated pentane/pentene stream 310.2 may be used in the production of carbon tetrachloride, advantageously minimising the production of waste products. This can be achieved through use of a high temperature chlorinolysis process.
The purified 1,1,1,3-tetrachloropropane product stream 310.1 is extracted from the system and may be combined with the major product stream (identified with reference numeral 110.4 in
The heavy ends residue extracted from the boiler 302 via line 303 is either disposed of or further processed.
Using the apparatus and process conditions outlined above, 2635 kg of carbon tetrachloride (CTC, 99.97% purity) was continuously processed with an average hourly loading 78.2 kg/h to produce 1,1,1,3-tetrachloropropene (1113TeCPa). Basic parameters of disclosed process carried out according to Example 2 are as following.
The full impurity profile of the purified product of the above-described embodiment is presented in the following table. Please note that the figures are given as a weighted average of the profiles for the product obtained in line 110.4 in
These examples were carried out using the equipment and techniques outlined above in the ‘Continuous Arrangement’ in Example 1, except where otherwise stated. The molar ratio of the chlorinated C3-6 alkane product (in this case, 1,1,1,3-tetrachloropropane):carbon tetrachloride in the reaction mixture was controlled to differing levels, principally by the residence time of reaction mixture in the alkylation zone. Temperature was maintained at 110° C. and pressure was maintained at 9 Bar. The selectivities towards product of interest are reported in the following table:
As can be seen from this example, when the molar ratio of product:starting material exceeds 95:5 when the process is operated on a continuous basis, there is a notable reduction in selectivity towards the product of interest.
These examples were carried out using the equipment and techniques as illustrated in
As can be seen from this example, when the molar ratio of product:starting material exceeds 95:5 when the process is operated on a continuous basis, there is a notable reduction in selectivity towards the product of interest.
These examples were carried out using the equipment and techniques as illustrated in
These examples were carried out using the equipment and techniques as illustrated in
As can be seen from this example, there is a conversion of ethylene between 75-93% in the plug-flow reactor. Thus if plug-flow reactor is employed there is more efficient ethylene utilization in the reaction section. The serial plug-flow reactor allows the CSTR reactor to be operated at an optimal pressure, without needing complex and uneconomical ethylene recovery processes. The serial plug reactor therefore controls the ethylene use in an efficient closed loop.
Fractional distillation equipment consisting of a 2-litre glass distillation four-neck flask equipped with condenser, thermometer, heating bath and vacuum pump system was set up. The distillation flask was initially filled with 1976 grams of reaction mixture obtained using the apparatus and techniques illustrated in
During distillation, pressure was gradually reduced from an initial pressure of 100 mmHg to a final pressure of 6 mmHg. During this period, 1792 grams of distillate (in different fractions) were extracted. During distillation, there was visible HCl gas formation and furthermore chlorobutane (the breakdown product from tributylphosphate ligand) was also formed in significant amounts namely between 1 to 19% for the distillate fractions. Upon these observations being made, the distillation was interrupted, distillation residue was weighed and analyzed and was found to have a Tetrachloropropane content of 16%. It was no longer possible to continue distillation without significant degradation of the Tributylphosphate ligand.
The 1,1,1,3-tetrachloropropane is converted to 1,1,3-trichloropropene in the continuously stirred tank reactor 403, which fulfils the role of the dehydrochlorination zone. The residence time of the reaction mixture in the reactor 403 is limited to prevent the excessive conversion of 1,1,1,3-tetrachloropropane to 1,1,3-trichloropropene and thus, the molar ratio of 1,1,3-trichloropropene:1,1,1,3-tetrachloropropane does not exceed 50:50.
A proportion of 1,1,3-trichloropropene is extracted from the reaction mixture through the use of distillation column 408. Reaction mixture is fed into the bottom of the distillation column 408 and a 1,1,3-trichloropropene rich stream is withdrawn as overhead vapours via line 409. A partial condenser 410 functions to extract gaseous hydrogen chloride from the 1,1,3-trichloropropene rich stream via line 411. The 1,1,3-trichloropropene rich stream is then fed via line 412 to a reflux divider 413, and a stream of purified 1,1,3-trichloropropene is taken off via line 415. A proportion of the 1,1,3-trichloropropene rich stream is fed back as a reflux to distillation column 408 via line 414.
A mixture comprising catalyst, unreacted 1,1,1,3-tetrachloropropane and a limited amount of 1,1,3-trichloropropene is extracted via line 404 from the reactor 403 to a filter 405. The obtained filter cake is extracted via line 406 and the filtrate is passed via line 407 for aqueous treatment, as shown in
In
A biphasic mixture is formed in the tank 505 and the organic phase is extracted from the tank 505 via line 506, filtered 507 and taken via line 509 for further treatment, as shown in
1,1,1,3-tetrachloropropane and 1,1,3-trichloropropene dissolved in the aqueous layer present in the washing tank 505 are extracted therefrom by means of a steam distillation column 511. Stripped chlorinated alkanes are passed via line 512 from the distillation column 511 to a condenser 513 and then via line 514 to a reflux liquid-liquid separator 515 where two layers are formed. The stripped 1,1,1,3-tetrachloropropane is then taken off as an organic phase via line 517 and an aqueous phase is refluxed back to the distillation column via line 516.
Turning to
A heavy ends residue is extracted from boiler 602 via line 603 and filtered 604. The obtained filter cake and liquid residue are extracted via lines 605 and 606 respectively and recycled or treated.
Using the apparatus and process conditions outlined above, 3563 kg of 1,1,1,3-Tetrachloropropane (1113TeCPa, 99.925% purity) was continuously processed with an average hourly loading 63.1 kg/h to produce 1,1,3-trichloropropene (113TCPe). Basic parameters of disclosed process carried out according to Example 8 are as following.
The full impurity profile of the purified product of the above-described embodiment is presented in the following table. The figures are given as a weighted average of the profiles for the product obtained in line 415 in
As can be seen, step 2) of the process can be operated to produce highly pure chlorinated alkene material.
This example was conducted using the apparatus and techniques employed in Example 8 above, except where otherwise stated. The continuously stirred tank reactor was operated at a temperature of 149° C. and at atmospheric pressure. The molar ratio of 1,1,3-trichloropropene:1,1,1,3-tetrachloropropane in the reactor was controlled such that it did not exceed 30:70. Using the apparatus and process conditions outlined in Example 8 above, 1543.8 kg of 1,1,1,3-Tetrachloropropane (1113TeCPa, 99.901% purity) was continuously processed with an average hourly loading 47.5 kg/h to produce 1,1,3-trichloropropene (113TCPe). Catalyst was added in the form of FeCl3 aqueous solution to provide a catalyst content of 66 ppm, based on feedstock 1113TeCPa. Basic parameters of disclosed process carried out according to Example 8 are as following.
The full impurity profile of the product of the above-described embodiment is presented in the following table. The figures are given as a weighted average of the profiles for the product obtained in line 415 in
As can be seen, when the dehydrochlorination reaction is controlled such that the molar ratio of 1,1,3-trichloropropene:1,1,1,3-tetrachloropropane does not exceed 30:70, the process of step 2) can be operated to produce highly pure chlorinated alkene material with the very high selectivity and in high yield. Of note is that 3,3,3-tetrachloropropene is only formed in trace amounts. This is particularly advantageous as 3,3,3-tetrachloropropene is a very reactive olefin contaminant with a free induced (activated) double bond and can be a precursor of highly problematic oxygenated impurities.
These examples were conducted using the apparatus and techniques employed in Example 8 above, except where otherwise stated. In each of these trials, the reaction progress was controlled such that there was a different ratio between 1,1,3-trichloropropene:1,1,1,3-Tetrachloropropane in the reaction mixture present in the reactor (403,
As can be seen from this example, when the specific apparatus and techniques employed, an increase in the molar ratio of the product to starting material (increased amount of the product in the reaction mixture) in step 2), this corresponds to an increase in the formation of heavy oligomers. Further, if the 1,1,3-trichloropropene concentration is high, catalyst deactivation was also observed.
An Erlenmeyer glass flask was filled with pure distilled 1,1,3-trichloropropene with purity of >99%. The test construction material sample was immersed in the liquid and the system was closed with a plastic plug.
Samples of the Trichloropropene were regularly taken from the flask. The construction material samples were weighed before and after trail. The temperature of the liquid was ambient laboratory conditions, around 25° C.
The major changes in the quality of the trichloropropene are shown in the following table, as a % change in purity:
In a second set of trials, an Erlenmeyer glass flask equipped with a back cooler and oil heating bath with controlled temperature was filled with pure distilled 1,1,3-Trichlorpropene with a purity of >99%. The test material sample was immersed in the liquid and the system was partially closed using a plastic plug. Samples of Trichloropropene were regularly taken from the flask. The material samples were weighed before and after trail. The temperature of the liquid was controlled at 100° C. The major changes in the liquid Trichloropropene are shown in the following table:
As can be seen from this example, the use of carbon steel appears to be challenging as it is not compatible with the process fluid consisting of 1,1,3-trichloropropene. Stainless steel and titanium have also poor performance, resulting in the formation of significant amounts of oligomers are formed. From the tested metal materials, the Ni-alloy Hastelloy C-276 has excellent results. It can be seen also that glass (or enamel) and other non-metallic material, such as phenolic resin impregnated graphite, are also more suitable.
In many downstream reactions in which chlorinated alkenes are used as starting materials, the presence of oxygenated organic impurities is problematic. This example demonstrates that certain impurities have a surprising propensity to form such compounds.
A four neck glass flask equipped with a stirrer, thermometer, back cooler, feed and discharge neck and cooling bath was filled with water and chlorine gas was bubbled into the water to produce a weak solution of hypochlorous acid. When an appropriate amount of chlorine had been introduced into the water, a feedstock consisting obtained from the process of Example 8 comprising 1,1,3-trichloropropene with a purity of 98.9% was slowly dropped into the prepared solution of hypochlorous acid for a period of 90 min and cooled. The pressure was atmospheric and the operating temperature was close to 20° C. The same procedure was repeated with 3,3,3-Trichloropropene having a purity of 68.1%. After reaction completion the systems formed biphasic mixtures. The organic phase (product) was extracted and then analyzed by gas chromatography. The results are shown in the following table:
As can be seen from this example, 1,1,3-Trichlorpropene reacts with chlorine in water to produce 1,1,1,2,3-pentachloropropane, while 3,3,3-trichloropropene reacts significantly to produce corresponding tetrachlorohydrines, especially 1,1,1,3-tetrachloropropan-2-ol.
In other words, 1,1,3-Trichlorpropene reacts to produce a product of commercial interest, while 3,3,3-trichloropropene reacts to the produce an oxygenated impurity which cannot be easily removed from the 1,1,1,2,3-pentachloropropane. As is apparent from Examples 8 and 9 above, the processes of step 2) can be advantageously employed to produce 1,1,3-trichloropropene resulting in the formation of only trace amounts of 3,3,3-trichloropropene.
A schematic diagram of the equipment used to perform the primary conversion step and principal conversion step in step 3-a) is provided as
The operating temperature within the primary reaction zone is 0° C. to 20° C. Operating the reactor within this range was found to minimise the formation of pentachloropropane isomers, which are difficult to separate from the target product, 1,1,1,2,3-pentachloropropane. Thorough mixing of the reaction mixture and mild temperatures, but also controlling the proportion of 1,1,1,2,3-pentachloropropane present in the reaction mixture, was found to minimise serial reactions of 1,1,3-trichloropropene and the formation of 1,1,1,3,3-pentachloropropane (which is difficult to separate from 1,1,1,2,3-pentachloropropane). To increase the rate of reaction at the low temperatures, the reaction mixture is exposed to visible light.
The reaction mixture is then passed up through the reactor 702 for the principal conversion step, which is performed as a reduced temperature conversion step. Cooling of the reaction mixture is achieved using cooling tubes, and the reaction mixture is passed through a series of upstream and downstream principal reaction zones (not shown), resulting in zonal chlorination of 1,1,3-trichloropropene. To drive the reaction towards completion, the reaction mixture in the downstream principal reaction zone is exposed to ultraviolet light. Advantageously, this fully utilizes the chlorine starting material such that the obtained reaction mixture which is extracted from the downstream-most principal reaction zone has very low levels of dissolved chlorine. Operating the principal reaction zones at such temperatures has been found to minimise the serial reactions of 1,1,3-trichloropropene, which result in the formation of unwanted and problematic impurities, such as hexachloropropane.
A 1,1,1,2,3-pentachloropropane rich stream is extracted from reactor 702 via line 708. Off-gas is extracted from the reactor 702 via line 711. The 1,1,1,2,3-pentachloropropane rich stream is subjected to cooling using a product cooler 709 and passed via line 710 for a hydrolysis step. A schematic diagram illustrating the equipment used to conduct this step is presented as
In that equipment, the 1,1,1,2,3-pentachloropropane rich stream is fed into washing tank 803 via line 802. Water is fed into the washing tank via line 801 to form a biphasic mixture. The organic phase (containing the 1,1,1,2,3-pentachloropropane rich product) can easily be separated from the aqueous phase by the sequential removal of those phases via line 804. The extracted phases are filtered 805 with the filter cake being removed 806. The 1,1,1,2,3-pentachloropropane rich product is then fed via line 807 for further processing while wastewater is removed via line 808.
The hydrolysis step is especially effective at removing oxygenated organic compounds, such as chlorinated propionyl chloride and their corresponding acids and alcohols, which may be formed during upstream steps in the process. While the formation of such compounds can be avoided by excluding the presence of oxygen from the upstream stages of the synthesis, doing so increases the cost of production. Thus, the hydrolysis step assists with the economic and straightforward removal of such otherwise problematic (owing to the difficulty of removing them, e.g. by distillation) impurities.
To maximise the purity of the obtained 1,1,1,2,3-pentachloropropane, a vacuum distillation step was performed, using the apparatus shown in
The vacuum distillation column 907 is provided with a liquid side stream withdrawal which can be used to prevent contamination of the product stream with light molecular weight compounds which may be formed in the boiler.
The 1,1,1,2,3-pentachloropropane rich product from the apparatus shown in
Distillate is taken from the distillation column 907 via line 908, fed via condenser 909, intermediate line 910 and liquid divider 911 to yield a streams of i) 1,1,3-trichloropropene via line 913.1 which is recycled to the primary reaction zone, ii) 1,1,1,3-tetrachloropropane via line 913.2 and purified 1,1,1,2,3-pentachloropropane via line 913.3. A reflux stream 912 from divider 911 is fed back into the vacuum distillation column 907.
Using the apparatus and process conditions outlined above, 3062 kg of 1,1,3-trichloropropene (113TCPe, purity 97.577%) was continuously processed with an average hourly loading 44.9 kg/h to produce 1,1,1,2,3-pentachloropropane (11123PCPa). Basic parameters of the process are as follows:
The purified product obtained in line 913.3 in
The process of Example 13 was repeated four times and samples of 1,1,1,2,3-pentachloropropane were obtained following distillation using the apparatus illustrated in
Crude 1,1,1,2,3-pentachloropropane compositions were obtained using the apparatus depicted in
As is apparent, the washing step can be successfully employed to minimise the content of oxygenated organic impurities in compositions rich in chlorinated alkanes of interest.
A batch operated reactor consisting of a four neck glass flask equipped with a stirrer, thermometer, back cooler, feed and discharge neck and cooling bath was set up. The feedstock consisted of 1,1,3-trichloropropene comprising perchlorethylene and oxygenated impurities in amounts observed in commercially sourced supplies.
Minor amounts of HCl gas were formed and these together with traces of chlorine were cooled down by means of a back cooler/condenser and then absorbed in a caustic soda scrubber. Chlorine was introduced into the liquid reaction mixture via dip pipe in various amounts for a period of 90 minutes. The temperature of reaction was maintained at 26 to 31° C. Pressure was atmospheric. The chlorine was totally consumed during the reaction. The reaction mixture was sampled and analyzed by gas chromatography and the results of this analysis are shown in the following table:
As can be seen, increasing the conversion of the chlorinated alkene starting material to the chlorinated alkane product of interest results in an increase in the formation of impurities in the reaction mixture. These disadvantageous results arise as conversion of the starting material to product approaches total conversion.
This example was carried out in as described in Example 16 above. 1,1,3-trichloropropene (purity 94.6% containing 5% of 1,1,1,3-tetrachloropropane as an impurity) was used as the feedstock.
4 trials at different reaction temperature were conducted. The samples of reaction mixture were taken at 80%, 90%, 95% and 100% of stoichiometric quantity of chlorine dosed (based on 113TCPe in the feedstock) and then analyzed by gas chromatography. The results of this analysis are shown in the following table:
These results demonstrate that increasing the conversion of the chlorinated alkene starting material to the chlorinated alkane product of interest results in a decrease in the selectivity of the reaction towards the chlorinated alkane isomer of interest. These disadvantageous results arise as conversion of the starting material to product approaches total conversion.
This chlorination step was carried out as described in Example 16 above. 1,1,3-trichloropropene (purity 99,4%) was used as a feedstock.
Chlorine was introduced into the liquid reaction mixture at 120% of the stoichiometric quantity towards feedstock 1,1,3-trichloropropene for a period of 90 minutes and was totally consumed during the reaction. The reaction temperature was 80° C. and reactor pressure was atmospheric. The samples of reaction mixture were taken by 80%, 95%, 110% and 120% of stoichiometric quantity of the chlorine dosed was analyzed by gas chromatography. Reaction selectivity is expressed in the table below as a ratio between sum of major impurities (1,1,3,3-Tetrachloropropene, 1,1,1,2,3,3-Hexachloropropane, 1,1,1,2.2.3-Hexachloropropane) to the product 1,1,1,2,3-pentachloropropane:
These results demonstrate that increasing the conversion of the chlorinated alkene starting material to the chlorinated alkane product of interest results in an increase in the formation of unwanted impurities. These disadvantageous results arise as conversion of the starting material to product approaches total conversion. As can be seen, the degree of conversion (and thus the formation of impurities) can advantageously and conveniently be achieved by controlling the amount of chlorine into the reaction zone, such that there is no molar excess of chlorine:chlorinated alkene starting material.
To demonstrate the effectiveness of the hydrolysis step of step 3-b) at removing oxygenated compounds from the chlorinated alkane product of interest, samples of crude reaction mixture reaction mixture were obtained using the apparatus depicted in
As can be seen from this example there is about 97.5% efficiency in the removal of this specific oxygenated impurity.
The equipment used is similar to the one described in
The fluorination of pure HCC-240db, obtained according to the present process, is performed with 9.2 liter of a commercial bulk Cr catalyst. Prior to use, the catalyst is activated.
No air is added in order to maintain the catalyst activity.
The system was continuously fed with anhydrous HF and HCC-240db (0.36 kg/hr) at P=5 bar. The contact time was 19 s and the reaction temperature was 350° C. The contact time is defined as the ratio of the volume of catalyst bed on the total volume flow rate in the experimental conditions of temperature and pressure.
The conversion of HCC-240db was 100%. The crude 1234yf depicted in
| Filing Document | Filing Date | Country | Kind |
|---|---|---|---|
| PCT/IB2016/000576 | 4/13/2016 | WO | 00 |
