Abstract:
A process for the selective hydrogenation of the methyl acetylene and propadiene (MAPD) in a propylene rich stream is disclosed wherein the selective hydrogenation is carried out stepwise (a) first in a single pass fixed bed reactor and then (b) in a distillation column reactor containing a supported PdO hydrogenation catalyst which serves as a component of a distillation structure. Conversion and selectivity to propylene in improved over the use of the single pass fixed bed reactor alone.

Description:
BACKGROUND OF THE INVENTION 
     1. Field of the Invention 
     The present invention relates to removal of MAPD from hydrocarbon streams. More particularly the invention relates to a process wherein the MAPD is converted to valuable propylene by selective hydrogenation. More particularly the invention relates to a process where one of the hydrogenation reactors is a distillation column reactor wherein the propylene is concurrently separated from a stream containing unconverted MAPD. 
     2. Related Information 
     Methyl acetylene/propadiene (MAPD) is not a compound but covers the unstable compounds methyl acetylene and propadiene which may be depicted as follows:                           
     The MAPD compounds are highly reactive contaminants in a propylene stream. The most common method of removal is by selective hydrogenation which not only “removes” the contaminants but converts them to valuable product propylene. The propylene stream is then fractionated to remove a portion of the stream which contains unconverted MAPD. The stream containing the unconverted MAPD is called “Green Oil” and the fractionation tower used to separate the stream containing the unconverted MAPD from the propylene is called the “Green Oil Tower”. 
     Prior methods of the selective hydrogenation of MAPD in propylene streams have used both liquid and vapor phase reactors. Generally while conversion is good, in both cases the selectivity drops off rapidly with time. Thus, it would be desirable to improve selectivity in general and keep the selectivity over time. Both the vapor phase and liquid phase system utilize at least two reactors in parallel, with sometimes three, leaving at least one spare for regeneration. 
     The main drawback of this method is that the selectivity to propylene is not always as desired. The by-product of the hydrogenation, however, is propane which is not considered a contaminant when further processing of the propylene is carried out. However, it does reduce the amount of valuable propylene produced. 
     The term “reactive distillation” is used to describe the concurrent reaction and fractionation in a column. For the purposes of the present invention, the term “catalytic distillation” includes reactive distillation and any other process of concurrent reaction and fractional distillation in a column regardless of the designation applied thereto. 
     A preferred catalytic distillation is one in which a distillation structure also serves as the catalyst for the reaction. The use of a solid particulate catalyst as part of a distillation structure in a combination distillation column reactor for various reactions is described in U.S. Pat. Nos. (etherification) 4,232,177; (hydration) 4,982,022; (dissociation) 4,447,668; (aromatic alkylation) 5,019,669 and (hydrogenation) 5,877,363. Additionally U.S. Pat. Nos. 4,302,356; 4,443,559; 5,431,890 and 5,730,843 disclose catalyst structures which are useful as distillation structures. 
     Hydrogenation is the reaction of hydrogen with a carbon-carbon multiple bond to “saturate” the compound. This reaction has long been known and is usually done at super atmospheric pressures and moderate temperatures using a large excess of hydrogen over a metal catalyst. Among the metals known to catalyze the hydrogenation reaction are platinum, rhenium, cobalt, molybdenum, nickel, tungsten and palladium. Generally, commercial forms of catalyst use supported oxides of these metals. The oxide is reduced to the active form either prior to use with a reducing agent or during use by the hydrogen in the feed. These metals also catalyze other reactions, most notably dehydrogenation at elevated temperatures. Additionally they can promote the reaction of olefinic compounds with themselves or other olefins to produce dimers or oligomers as residence time is increased. 
     Selective hydrogenation of hydrocarbon compounds has been known for quite some time. Peterson, et al in “The Selective Hydrogenation of Pyrolysis Gasoline” presented to the Petroleum Division of the American Chemical Society in September of 1962, discusses the selective hydrogenation of C 4  and higher diolefins. Boitiaux, et al in “Newest Hydrogenation Catalyst”,  Hydrocarbon Processing , March 1985, presents a general, non enabling overview of various uses of hydrogenation catalysts, including selective hydrogenation of a propylene rich stream and other cuts. Conventional liquid phase hydrogenations as presently practiced required high hydrogen partial pressures, usually in excess of 200 psi and more frequently in a range of up to 400 psi or more. In a liquid phase hydrogenation the hydrogen partial pressure is essentially the system pressure. 
     UK Patent Specification 835,689 discloses a high pressure, concurrent trickle bed hydrogenation of C 2  and C 3  fractions to remove acetylenes. The selective hydrogenation of MAPD in propylene streams utilizing catalytic distillation alone is disclosed in International Application WO 95/15934. 
     It is an advantage of the present process that the propadiene and methyl acetylene contained within the hydrocarbon stream contacted with the catalyst are selectively converted to propylene with very little if any formation of oligomers or little if any saturation of the mono-olefins contained in the feed. 
     SUMMARY OF THE INVENTION 
     The present invention comprises the selective hydrogenation of methyl acetylene and propadiene (MAPD) contained within a propylene rich stream to purify the stream and obtain greater amounts of the propylene. In a class of preferred embodiments the propylene rich stream is fed along with hydrogen first to a standard single pass fixed bed reactor and the effluent then fed to a catalytic distillation column reactor and contacted with hydrogen in a reaction zone containing a hydrogenation catalyst, such as supported palladium oxide catalyst, preferably in the form of a catalytic distillation structure. The hydrogenation catalyst in the single pass fixed bed reactor may the same or different from that in the catalytic distillation column reactor. Hydrogen is provided as necessary to support the reaction and, it is believed, to reduce the oxide and maintain it in the hydride state. The distillation column reactor is operated at a pressure such that the reaction mixture is boiling in the bed of catalyst. If desired, a bottoms stream containing any higher boiling material (the Green Oil) may be withdrawn to effectuate a complete separation. 
     The single pass fixed bed reactor(s) may be any of those known in the art, as may the hydrogenation catalyst. 
     The hydrogen rate must be adjusted such that it is sufficient to support the hydrogenation reaction and replace hydrogen lost from the catalyst but kept below that required for hydrogenation of propylene and, in the case of the catalytic distillation reactor, to prevent flooding of the column which is understood to be the “effectuating amount of hydrogen ” as that term is used herein. Generally the mole ratio of hydrogen to methyl acetylene and propadiene in the feed to the fixed bed of the present invention will be about 1.05 to 2.5 preferably 1.4 to 2.0. 
     In some embodiments the process may be described as comprising the steps of: 
     (a) feeding (1) a first stream comprising propylene, methyl acetylene and propadiene and (2) a second stream containing hydrogen to a single pass fixed bed reactor wherein a portion of the methyl acetylene and propadiene react with the hydrogen to produce propylene; 
     (b) feeding the effluent from step (a) to a distillation column reactor into a feed zone; 
     (c) concurrently in said distillation column reactor 
     (i) contacting unreacted methyl acetylene and propadiene with hydrogen in a distillation reaction zone with a hydrogenation catalyst capable of acting as a distillation structure thereby reacting a further portion of said methyl acetylene and propadiene with said hydrogen to form additional propylene, and 
     (ii) separating the propylene contained by fractional distillation and 
     (e) withdrawing the separated propylene along with any propane and lighter compounds, including any unreacted hydrogen, from said distillation column reactor as overheads. Optionally the process may include withdrawing any C 4  or higher boiling compounds from said distillation column reactor as bottoms. There is no significant loss of propylene from the hydrogenation. 
    
    
     BRIEF DESCRIPTION OF THE DRAWING 
     FIG. 1 is flow diagram in schematic form of one embodiment of the invention. 
     FIG. 2 is flow diagram in schematic form of a second embodiment of the invention. 
     FIG. 3 is flow diagram in schematic form of a third embodiment of the invention. 
    
    
     DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT 
     The feed for the instant process may generally be the bottoms from a deethanizer of an ethylene plant. However, any streams which contain propylene contaminated with MAPD would be a candidate for the process. In a particular embodiment of the invention conventional single pass fixed bed reactors, preferably vapor phase, are combined with a distillation column reactor to enhance conversion and selectivity in place of the Green Oil Tower. The advantages of using a distillation column reactor in place of the Green Oil Tower are: 
     1. Propylene production from the MAPD converters can be increased by a factor of 2.5 thereby increasing unit productivity. The catalytic distillation hydrogenation uses hydrogen more efficiently, thereby providing an operating savings. The attributes of a distillation type reaction, e.g. limited product contact with catalyst and the washing effect of internal reflux, result in less build up of coke on the catalyst and longer reaction times between regenerations for further operations savings. 
     2. Regeneration frequency is reduced by eliminating the operation of the second bed. 
     3. The conversion of the Green Oil Tower to catalytic distillation hydrogenation is simple in scope and can be accomplished during a normal plant turnaround. The impact on plant operations is minimal with the potential of simplifying the operation of the unit by eliminating a reactor. 
     4. Spare reactor vessels can be used for other services. 
     Catalyst suitable for the present process include 0.05-5 wt % PdO on extruded alumina, such as, 0.3 wt % PdO on ⅛″ Al 2 O 3  (alumina) extrudates, hydrogenation catalyst, supplied by United Catalysts Inc. designated as G68F. Typical physical and chemical properties of the catalyst as provided by the manufacturer are as follows: 
     
       
         
               
               
               
             
           
               
                   
                 TABLE I 
               
               
                   
                   
               
             
             
               
                   
                 Designation 
                 G68F 
               
               
                   
                 Form 
                 spheres 
               
               
                   
                 Nominal size 
                 3 × 6 Mesh 
               
               
                   
                 Pd. wt % 
                 0.3 
               
               
                   
                 Support 
                 High purity alumina 
               
               
                   
                   
               
             
          
         
       
     
     This catalyst may be used in either the single pass reactors or the distillation column reactor (Green Oil Tower). However to be used in the distillation column reactor it must be in a form to also serve as a distillation structure which may, for some distillation columns, be a simple loading of the catalyst. There are several methods and structures available for better performance which are variously described in U.S. Pat. Nos. 4,215,011; 4,439,350; 4,443,559; 5,057,468; 5,189,001; 5,262,012; 5,266,546; 5,348,710; 5,431,890; and 5,730,843 all of which are hereby incorporated by reference. One of the preferred catalyst structures being that described in U.S. Pat. No. 5,730,843. 
     The structure of U.S. Pat. No. 5,730,843 comprises a rigid frame made of two substantially vertical duplicate grids, spaced apart and held rigid by a plurality of substantially horizontal rigid members and a plurality of substantially horizontal wire mesh tubes mounted to the grids to form a plurality of fluid pathways among the tubes. For use as a catalytic distillation structure, which serves as both the distillation structure and the catalyst, at least a portion of the wire mesh tubes contain a particulate catalytic material. The catalyst within the tubes provides a reaction zone where catalytic reactions may occur and the wire mesh provides mass transfer surfaces to effect a fractional distillation. The spacing elements provide for a variation of the catalyst density and loading and structural integrity. 
     One typical embodiment of this invention is shown in FIG.  1 . It includes a depropanizer  10  to which a C 3 + stream containing propane, propylene, MAPD, higher boiling olefins and higher boiling paraffins are fed via flow line  101 . The bottoms from the depropanizer contain C 4  and higher boiling materials which are taken via flow line  102  for further processing. The overheads from the depropanizer contain the propylene, MAPD, and propane which are taken via flow line  103  and are condensed in condenser  24  and collected in receiver  22  and fed to a via flow line  106  to first vapor phase single pass reactor  20  containing a fixed bed of hydrogenation catalyst. Non condensibles are removed via line  109 . Typical operating conditions include 215-315 psig and 100-250° F. A portion of the condensed overheads is returned to the depropanizer  10  as reflux via flow line  110 . Hydrogen is added to the reactor via flow line  104 . In the reactor  20  a portion of the MAPD is converted to propylene and propane. The first reactor may act as a guard bed to remove catalyst poisons such as arsine, mercury or methanol. The effluent from the first reactor  20  is taken via flow line  105  to second single pass reactor  30  containing a fixed bed of the same or similar hydrogenation catalyst as in reactor  20  wherein additional MAPD is converted to propylene and propane. The effluent from the second reactor  30  is fed via flow line  106  to distillation column reactor  40 . Hydrogen as needed is added via flow line  107 . This embodiment may also be configured to use the two primary reactors  20  and  30  alternatively by switching the overheads from depropanizer  10  between the reactors by selection of through valve  106   a  and the corresponding adjustment of the flow line  105  through valve  105   a . In this configuration a single reactor serves as the guard bed for the catalytic distillation column, while the catalyst in the other reactor is regenerated or replaced. 
     In a variant of this embodiment the system can be operated and low pressures, i.e., 90-120 psig. In this case the feed to reactors  20  and  30  may be through line  109 . The operating temperature is between 100-250° F. Conversely, in yet another variant the system can be operated at higher pressures in an all liquid mode by pumping the material in the feed line  106  to about 400 psig and heating prior to entering reactors  20  and  30 . The temperatures are in the same range. 
     Distillation column reactor  40  is seen to contain a bed  41  of the same or similar catalyst as in the two reactors  20  and  30  but in a form so as to act as both catalyst and distillation structure. A stripping section  42  containing standard distillation apparatus such as bubble cap trays, sieve trays or packing is provided below the bed  41  to assure that all of the C 3 &#39;s are removed in the overheads. A rectification section  44  also containing standard distillation apparatus such as bubble cap trays, sieve trays or packing is provided above the bed  41  to assure complete separation. The Green Oil containing unconverted MAPD is removed in the bottoms via flow line  108  and returned to the depropanizer where any C 3  material is taken as overheads and recycled to the reactors. 
     The overheads from the distillation column reactor  40  are taken via flow line  111  and passed through partial condenser  50  where the C 3 &#39;s are condensed and collected in receiver/separator  60 . Uncondensed material, included unreacted hydrogen are taken via flow line  112  with the hydrogen being recycled if desired. The C 3  liquid is taken via flow line  113  with a portion being returned to the distillation column reactor  40  as reflux via flow line  114 . 
     Using the processing sequence of this invention the overall selectivity to propylene at start of run is 75%. At end of run the selectivity is about 50%. This is compared to a start of run of about 50% and end of run of 0% for conventional vapor phase converters. 
     In a another embodiment of this invention as shown in FIG. 2 the reactors  20  and  30  are deleted. The C 3  feed is directly to the distillation column reactor via flow line  106 . The depropanizer and distillation column are otherwise identical to FIG.  1 . 
     In another embodiment (shown in FIG. 3) a C 3  splitter  70  is positioned after the distillation column reactor  40 . The overheads  111  go directly to splitter  70  via flow line  113  and reflux to the distillation column reactor  40  comes from the splitter  70  via flow line  114 . Otherwise the operation of the depropanizer  10  and distillation column  40  is the same as in FIG.  2 . The overheads from the splitter in flow line  121  contain the lighter material. Any condensible material is condensed in condenser  80  and collected in receiver/separator  90  for reflux to the splitter via flow line  124 . Non condensibles are vented flow line  122 . Propylene is taken from the column below the overheads via flow  123  and propane is removed as bottoms via flow line  128 . 
     The above embodiments are offered as typical but not limiting illustrations of the flexibility offered by the process and scheme of this invention. Additional variations, configurations and conditions based on the invention would be readily evident to those skilled in the art.