CA2019001A1 - Process for manufacturing syngas - Google Patents
Process for manufacturing syngasInfo
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- CA2019001A1 CA2019001A1 CA 2019001 CA2019001A CA2019001A1 CA 2019001 A1 CA2019001 A1 CA 2019001A1 CA 2019001 CA2019001 CA 2019001 CA 2019001 A CA2019001 A CA 2019001A CA 2019001 A1 CA2019001 A1 CA 2019001A1
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Abstract
Abstract An improved method for producing syngas wherein hot reactor effluent is briefly contacted with a C3 to C20 alkane quench material to remove part of the heat contained in the raw reactor effluent, and is then further quenched by thermal quenching means to achieve a temperature which discourages retrograde reactions.
Description
IMPROVED PROCESSES FOR MANUFACTURING SYNGAS
Field of the Invention This invention relates to improved processes for manufacturing syngas and higher hydrocarbons and oxygenates from methane and/or natural gas, and more particularly relates to a process wherein high quality heat wasted in prior art processes is captured, and additional hydrogen is provided for use in converting low value carbon oxides to more vaiuable products.
Backqround of the Invention Syngas is often manufactured as an intermediate in the conversion of feedstocks containing carbon and hydrogen to lS more desired hydrocarbons or oxygenates. This technique is particularly useful as a means for converting methane to higher hydrocarbons or oxygenates which are more readily transportable from the field and/or more desirable products such as alcohols.
A major source of methane is natural gas which typically contains about 40-95% methane depending on the particular source. Other constituents include about 10% ethane with the balance being made up of CO2 and smaller amounts of propane, butanes, pentanes, nitrogen, etc.
Primary sources for natural gas are reservoirs either alone or generally associated with hydrocarbon liquid reserves.
Most of the natural gas used for heating purposes comes from these sources. Quantities of natural gas are also known to be present in coal deposits and are by-products of crude oil refinery processes and bacterial decomposition of organic matter.
Prior to commercial use, natural gas must be processed to remove water vapor, condensible hydrocarbons and inert or poisonous constituents. Condensible hydrocarbons are generally removed by cooling natural gas to a low temperature, 27,418 and then washing it with a cold hydrocarbon liquid to absorb the condensible hydrocarbons. The condensible hydrocarbons are typically ethane and heavier hydrocarbor.s. This gas processing can occur at the wellhead or at a central processing station.
Processed natural gas typically comprises a major amount of methane, and minor amounts of ethane, propane, butanes, pentanes, carbon dioxide and nitrogen. Generally, processed natural gas comprises from about 70% to more than about 95% by volume of methane.
Most processed natural gas used commercially is distributed through extensive pipeline distribution networks.
As natural gas reserves in close proximity to gas usage declines, new sources that are more remote require transportation over further distances. The gas from many of these distant sources are not, however, amenable to transport by pipeline. For example, the gas from sources that are located in areas requiring economically unfeasible pipeline networks or in areas requiring transport across large bodies of water are not amenable to transport by pipeline. This problem has been addressed in several ways. One such solution has been to build a production facility at the site of the natural gas deposit to manufacture one specific product. Another approach has been to liquefy the natural gas and transport it in specially designed tanker ships. Natural gas can be reduced to 1/600th of the volume occupied in the gaseous state by cryogenic processing, and with proper procedures, safely stored or transported. These processes, which involve liquefying natural gas, transporting the liquified gas, and revaporizing it are complex, energy intensive and expensive.
Still another approach has been the conversion of natural gas to higher molecular weight hydrocarbons or oxygenates, preferably substantially liquid hydrocarbons or oxygenates, that can be easily handled and transported. The conversion of natural gas to higher order hydrocarbons or 27,418 -2~J~
oxygenates, especially ethane and ethylene, methanol or higher alcohols retains the material's versatility for uses as precursor materials in chemical processing. Known dehydrogenation and polymerization processes are available for the further conversion of ethane and ethylene to liquid hydrocarbons and of alcohols to numerous valuable products. In thase ways, easily transportable commodities may be derived from natural gas at the wellhead. A drawback in implementing such processes has been the lack of means for obtaining a sufficiently economical conversion rate of natural gas to higher molecular weight hydrocarbons.
To date, processes to convert natural gas to hydrocarbons or alcohols rely on the manufacture of a synthesis gas or syngas (carbon monoxide and hydrogen) with a subsequent conversion of syngas to the desired product. Examples of this category of processes included Shell Oil's Shell Middle Distillate (SMDS) process, the Gulf Badger Fischer-Tropsch process, the Mobil Oil MTG and MOGD processes, commercially available processes for the manufacture of methanol, and a variety of processes proposed for the manufacture of higher alcohols.
When methanol or other alcohols are the desired products or intermediates, or water the desired by-product in the production of aliphatic or mono-olefinic hydrocarbons or oxygenates from syngas, the raw syngas should theoretically have the following molar relationship:
(H2-C02)/(CO2-~CO) 2 This ratio will be referred to henceforth as the H2/COx ratio.
In practice for instance, a H2/COx ratio of 2-2.1 (mole/mole) is often used for the production of methanol. Since the H2/COx ratio in the described products is always less than or equal to the maximum possible from a methane feed (e.g. the ratio for methane and methanol is equal), water is the usual by-product of choice.
27,418 ~$~
While natural gas is a convenient source, syngas may also be derived from streams produced from the burning of coal and streams formed during numerous manufacturing or processing operations and in the above- or below-ground gasification or liquefication of coal, tar sands, oil shale and biomass, for example.
Several techniques are used to manufacture syngas, including partial oxidation, reforming and a process which can be viewed as a hybrid of the two, commonly referred to as autothermal reforming.
Partial oxidation relies on the partial combustion of hydrocarbons with an oxygen-containing gas to produce a high temperature mixture of carbon monoxide, hydrogen, water and carbon dioxide. When methane is used as the hydrocarbon feed, a hydrogen to carbon monoxide ratio of between 1.8 to 1.9 is the usual result, with the ratio being at least partially dependent on the composition of the feed stream. Carbon dioxide in the feed stream, for instance, will produce a somewhat lower ratio, and hydrogen or water in the feed stream will have a tendency to increase the ratio.
One way to achieve the desired H2/COX ratio is to treat the syngas with a shift catalyst to increase the hydrogen content at the expense of carbon monoxide according to the following formula:
H2O + CO = H2 +C2 Water is then rejected, thus resulting in a higher H2/COX ratio syngas. However, this technique has a detrimental side effect, since carbon dioxide concentrations well in excess of those needed for optimum H2/COX ratios are required to force the above equilibrium to the left. This requirement makes the approach relatively inefficient.
Classical reforming overcomes some of the H2/COX
ratio limitations encountered in partial oxidation, but presents other drawbacks. More of the available feed must be 27,418 h diverted to the reformer furnace to provide the heat to drive the idealized endothermic reforming reaction:
CH4 + H20 3H2 + CO
This feed derived fuel is generally burned with air, and the carbon it contains is lost as carbon oxides vented to the atmosphere, rather than to increase product yield. Thus, although thP H2/COX ratio in the raw syngas is greater than that necessary to achieve a stoichiometric optimum, that increase is offset by the decrease in yield attributable to vented combustion products.
In general, the raw syngas exits from the reformer at a lower temperature than is typical in either partial oxidation or autothermal reforming because of the metallurgical limitations of the tubes used in the reforming furnace.
Autothermal reforming is the means by which the industry through the use of current technology has attempted to balance the extremes in trade-offs between partial oxidation, with its low H2/COx ratio and reduced use of feed as fuel, and classical reforming, with its high H2/COX ratio and increased use of feed as fuel. As practiced, autothermal reforming consists of partially reforming a hydrocarbon such as methane, but not at the extremes practiced in classical reforming, and at lower fuel duty. The resultant effluent stream obtained through autothermal reforming is relatively high in hydrogen and also high in methane. This stream is subsequently mixed with oxygen or oxygen-containing gas and further partially combusted to reduce its methane and hydrogen content to produce a product stream having an improved H2/COX ratio. In addition, feed requirements as fuel are reduced.
Representative processes are disclosed in, for example, United States Patent Nos. 3,594,425; 3,631,111 and 4,390,314.
An important limitation and inefficiency of prior art processes, including partial oxidation, reforming and 27,418 autothermal reforming, is the method of heat removal from the processing units. The reactions are conducted under severe conditions, i.e., temperatures which may be in excess of 2000F. Following partial oxidation or other processes, the effluent stream must be rapidly cooled or quenched to a temperature at which the occurrence of retrograde reactions can be reduced. Typically the temperature of the reactor effluent must be lowered to 700C or even 600C to reduce the formation of carbonaceous deposits, the formation of undesirable aliphatic compounds from desirable olefins, etc. This temperature reduction generally is accomplished in from less than one second to a few seconds by quenching with water or other heat absorbing medium. This approach has major drawbacks which the present invention overcomes. Specifically, valuable high quality (high temperature) heat is wasted, and low quality (low temperature heat, which must then be disposed of at significant cost, is increased.
We have found that the thermal efficiency, and thus the economic attractiveness of syngas-based processes can be improved by the use of an alkane quench. The quenching of the high temperature partial oxidation reactor effluent, as well as the effluent from classical reforming and autothermal reforming processes, with a stream containing alkanes, is a means of efficiently raising the H2/CO~ ratio of the resultant stream without causing a detrimental rejection of carbon dioxide.
Thermal or physical quenching has heretofore been employed in prior art cracking processes involving gross volumes of heavy hydrocarbon feeds where heat is fed into the system to promote feedstock cracking, after which the reaction is quenched to prevent retrograde reactions. Generally speaking, prior art processes that employ hydrocarbons as a heat sink are doing so as a mechanism to reduce the process stream volume and the size of heat recovery equipment downstream which is required if water vaporization is used to 27,418 absorb heat and lower stream temperatures. Such prior art techniques, however, including equipment are those described in United States Patent Nos. 4,520,224, 4,288,408, 4,556,749 and 4,384,160, all rely on specific heat capacity, phase change, heat of vaporization, and the like to dilute the concentration of heat thus lowering stream temperature. These prior art processes are applicable to subsequent unit operations, i.e., those performed downstream of the process of the present invention. The prior art processes produce lower quality heat downstream resulting in the requirement of more heat rejection equipment in the form of cooling towers, vaporizers, heat exchangers and the like. Further, the prior art cracking processes which employ some form of thermal or physical quenching, including those that employ a hydrocarbon quench medium such as recycled heavy oil, produce methane in a form only useful as process fuel, the very material the present invention is designed to convert to more transportable, higher hydrocarbons.
United States Patent No. 3,022,148 discloses the use of an organic quench in the non-catalytic partial oxidation of hydrocarbon oil at a temperature of approximately 2500F, specifically, a hot oil stream wherein the oil has an average boiling range of from 500-750F. The requirement of the use of a high boiling point oil stream as a quench medium is not only energy intensive and therefore very costly, but impractical from the standpoint of converting methane to higher hydrocarbons or oxygenates at remote sites where the heavy, high boiling points oils would have to be transported in for processing.
The use of an alkane quench in syngas based processes, on the other hand, has a number of advantages. The product yield, thermal efficiency, and hence economic attractiveness is improved by the alkane quench techniques of 27,418 the present invention. Further, the quench material is likely to be readily available for processing at the site.
Thus, the present invention overcomes drawbacks of the prior art and improves the economic feasibility of both catalytic and non-catalytic partial oxidation, classical reforming and autothermal reforming processes.
Summary of the Invention The improved processes of this invention for producing syngas comprises contacting a hydrocarbon feedstock with an oxygen-containing material in a reaction zone of a processor to produce an effluent stream comprising reaction products consisting essentially of COx H2, and optionally c2+
hydrocarbons or oxygenates; and removing heat from the effluent stream by briefly contacting hot reactor effluent with a C3-C20 alkane quench material having a temperature of up to lQ00C to produce a chemically quenched stream wherein part of the heat contained in the raw reactor effluent has been removed through endothermic cracking reactions. The temperature of the chemically quenched effluent stream is further reduced by thermal means, using, for example, water, recirculating refractory oils, heat exchangers, or other heat absorbing means to achieve a temperature which further discourages or retards retrograde reactions. Thereafter, water, C4+ hydrocarbons and like byproducts are removed, and the effluent stream is processed in a synthesis reactor and product recovery system via methods well known in the art whereby oxygenates and hydrocarbons are formed, separated and stored. The process may be carried out in a partial oxidation, classical reforming or autothermal reforming reactor.
In another embodiment, streams comprising methane, light hydrocarbons, carbon oxides, and hydrogen are converted to C2+ hydrocarbons and oxygenates via pyrolysis at 1500 D F
employing an alkane quench to lower the temperature of the 27,418 reactor effluent stream as described above, and thereafter processing said stream by methods well known in the art.
More particularly, in one embodiment, the process of the present invention provides a method for converting a methane rich feedstock to syngas and olefins comprising the steps of: reacting the feedstock with an oxygen-containing material in a reaction zone of an partial oxidation reactor to produce an effluent stream comprising reaction products consisting essentially of COx, H2, and C2+ hydrocarbons; and intxoducing an alkane quench material having a temperature of up to 1000C to produce a chemically quenched effluent stream having a reduced temperature. Thereafter, the process of this invention may comprise the additional steps of further reducing the temperature of the chemically quenched stream by thermal means to retard retrograde reactions; and separating and removing C2+ hydrocarbon, water, hydrogen and carbon oxide products from the product stream. If desired, at least a portion of unreacted methane, unreacted alkane quench material, alkanes formed during chemical quenching, and mixtures thereof contained in the product stream may be recycled to the reaction zone. In addition, at least a portion of the C2+ hydrocarbons may be reacted to form higher homologs, oxygenates and alcohols.
C2+ hydrocarbons may be separated and removed from unreacted methane either from the effluent stream following the chemical quench or from the product stream following the thermal quench.
The hydrocarbon quench of this invention provides a number of major benefits. First, it absorbs some of the heat otherwise removed by wasteful techniques through an extremely fast endothermic reaction. Second, at least a portion of the hydrocarbon quench is converted to reactive intermediate or valuable end products. A further advantage is the production of additional hydrogen that can be used to improve the value of 27,418 end products or increase yields by converting carbon oxides to valuable products or intermediates.
It is presently preferred to employ the process of this invention in the vicinity of a gas extraction plant.
Ethane, propane, butanes, other natural gas liquids, condensates, etc., which are presently several of the preferred quench media for use in the gasifier, can be isolated at low cost in many gas plants. Furthermore, in remote locations where methane conversion technologies are likely to be implemented, these components usually have a low value, and are expensive to transport to markets.
The present invention is also applicable to producing streams comprising syngas and olefins from syngas containing streams derived from other sources, such as from coal deposits and formation during numerous mining operations and in the above- or below-ground gasification or liquefaction of coal, tar sands, oil shale and biomass, for example.
Brief Description of the Drawina FIG. 1 shows a process flow scheme for a representative partial oxidation process leading to the production of C2+ products in accordance with this invention.
FIG. 2 shows a process flow scheme for a representative classical reforming process in accordance with this invention.
FIG. 3 shows a process flow scheme for a representative autothermal reforming process in accordance with this invention.
Detailed Description of the Invention Referring to FIG. 1, there is depicted one preferred embodiment in which partial oxidation and quench operations are accomplished in different zones of a reaction vessel. Other configurations for contacting hot reaction gases with hydrocarbon quench materials, such as separate units, may be 27,418 employed, alternative hydrocarbon feeds may be used, and alternative syngas processing techniques may be employed to result in alternative products.
Methane, illustrative of a feedstock such as natural gas, is introduced into the system by line 12 where it is mixed with oxygen, air, or other oxygen-source gas supplied through line ll. The resulting mixture is fed into a first zone 14 of partial oxidation reactor 13 where it reacts to form syngas.
The hot gaseous feedstream passes on to hydrocarbon quench zone where it is mixed and reduced in temperature by the introduction of alkane quench material, after which the reactor stream progresses to final quench zone 16 at which point- a diluent such as water or oil is introduced to further cool the stream and prevent retrograde reactions.
Depending on the level of temperature reduction and downstream processing desired, several alternative flow configurations may be used, two of which are for the manufacture of representatiYe products such as Cl-C4 alcohols and are described below. If the effluent is cooled below the condensation point of water, the reactor effluent passes through line 17 into feed gas adjust zone 18 where some C2+
hydrocarbon product and water, etc. ara withdrawn through line 19 and C02 byproduct are withdrawn and recycled through line 20. Thy stream is then treated in synthesis reactor and product recovery system 30. Mixed alcohols (a mixture of Cl-C4 alcohols) are fed into mixed alcohol stabilization and storage zone 31 Yia line 32. The retaining product is fed into methanol reactor and product recovery system 34 via line 35.
Water is recovered from reactor 30 through line 36. Methane, SO hydrogen, and inert are separated and removed from reactor 34 via line 37a to prevent exces ive buildup of inerts. Carbon dioxide recovered from synthesis reactor 34 may be recycled via line 37 to the syngas generator. Methanol is fed to methanol 27,418 ~J~ 9~
product stabilization and storage unit 38 via line 39 and may also be recycled to synthesis reactor 30 via line 35a.
If desired, the process flow scheme depicted in FIG.
1 may be changed to allow for the production of intermediates, such as olefins (e.g., ethylene, propylene, etc.), oxygenates such as alcohols (e.g., methanol, ethanol, etc.) or other hydrocarbons the incorporation of carbonylation, hydroformylation, carbon oxide hydrogenation and other processes known to those skilled in the art to produce derivative products.
As shown in FIG. 2, in one embodiment of the present invention as applied to classical reforming of natural gas, natural gas and steam are fed into team reformer 40 via line 41. The effluent stream exitinq reformer 40 via line 42 is contacted with alkane quench material in quench zone 43 and subsequently fed via line 44 into gay cooler 45 where it is further cooled. The further cooled effluent stream exits gas cooler 45 via line 46 and is fed into C02 removal zone. C02 is removed via line 48 And the effluent stream is fed into compressors 50 via line 49 where it it converted to a high pre sure tree co~pri~ing synthesis gas and olefins.
PIG. 3 depicts one embodiment ox the present invention a applied to autothermal reforming. Natural gas and steam are fod into feed preheater via line 61. The preheated stream exits preheater 60 via line 62 where it is mixed with oxygen and fed into an autothermal reformer reactor 63. The effluent stream exiting reactor 63 via line 64 is cooled by contact with an alkane quench materlal in quench zone 65. The cooled stream is passed through gay cooler 67 via line 66.
Steam and water are removed via lines 68 and 69, respectively, and C02 ~ub3equently removed in zone 70. The stream then is fed into compressor zone 71 via line 72 and a high pressure stream comprising synthesis gas and olefin~ is recovered via line 73.
27,418 2 3 :~
As shown in FIGS. 2 and 3, the steam generated by the processes is employed to generate electricity.
The hydrocarbon quench employed in the practice of this invention may be a C2-C20, and preferably C3-C20 straight or branched chain alkane, or mixtures thereof.
The term "C2-C20 alkane", as used herein, refers to a straight or branched-chain alkane, i.e., ethane, propane, n-butane, 2-methylpropane, n-pentane, 2-methylbutane, 2,2-dimethylpropane, n-hexane, 2-methylpentane, 2,3-dimethylbutane, n-heptane, 2,3-dimethylpentane, n-octane, 3~ethyl-2,3-dimethylhexane, n-dodecane, n-octadecane, n-eicosane, and the like. Suitable quench materials also may include naphthas.
The term "naphthas", as used herein, refers to hydrocarbon product mixtures that distill at a boiling range below 180-190C.
It is preferred to employ paraffinic hydrocarbons as the quench material, more particularly light alkanes, i.e. C3-C5 alkanes. However, the presence of moderate amounts of aromatics, while not desirable, would not generally be detrimental to the process of this invention. Rather, aromatics would serve as diluents. However, the presence of naphthenic components are not desirable since they have a lesser tendency to produce desirable olefins. The term "naphthene~ or ~naphthenic components" refers to multiringed compounds containing both aromatic and saturated rings, also referred to as cycloalkylaryl compounds.
It is especially preferred to employ light hydrocarbon, preterably C3-C5 alkanes, in the practice of this inVQntiOn. However, as a practical matter, the actual selection of the hydrocarbon quench material will depend upon what is available in the field. Varying the selection of the alkane will vary the product distribution.
The use of an alkane quench material to cool the effluent stream causes endoth~rmic dehydrogenation and 27,41~
2~ Jo decomposition of said alkane quench material and a substantial increase in C2+ hydrocarbon products.
Generally speaking, the alkane quench comprises from 1 to 2S weight percent of the process stream just prior to the alkane quench addition, preferably from 5 to 20 weight percent and most preferably from 10-20 weight percent. The exact amount of quench material used will depend on the volume and composition of the quench material available, the conditions at which the partial oxidation, reforming or autothermal reforming segment is run, and the desired downstream composition.
The alkane quench material may be employed at ambient temperature, but is preferably heated to temperatures of up to 10007C, preferably, above 300C. In cases where there is a cryogenic unit at the site, temperatures as low as -lS0C may be employed. However, for best results, it is preferred to introduce the alkane quench material into the quench reactor at elevated temperatures as discussed above, most preferably, if feasible, at the temperature of the effluent stream. The optimum preheat temperature will be governed by the composition of the quench, and process conditions upstream and downstream of the quench operation.
The upper limitation on the temperature of the quench material is the thermal stability of the material. This will, in part, depend upon whether or not the material is diluted and the diluent. For example, if a quench material is diluted with carbon dioxide, steam or other diluents, it may be heated to a higher temperature without degrading.
The process of this invention may be run under a wide range of pressures, i.e. atmospheric to 1500 psi. It is preferable to operate at pressures of at least 20 psi, more preferably, from about S0 psi to 1000 psi.
The temperature at which the partial oxidation, reforming, autothermal reforming or pyrolysis reaction is run will depend upon the type of reactor, whether a catalyst is 27,418 2 so employed and if so, the particular catalyst employed.
Generally speaking, in the case of a partial oxidation operation, the temperature should be maintained at from about 1000-1450~C, depending upon the particular catalyst and equipment employed. In the case of pyrolysis, the temperature should be maintained from about 1300 to 2000C. For classical reforming, the temperature should be maintained at from about 800 to 950C and for autothermal reforming, from about 800 to 1250C.
It is critical that the temperature of the effluent stream leaving the reactor be maintained at a minimum of 600 D C
and preferably not drop below 700C for best results.
The use of an alkane chemical quench, immediately followed by a thermal quench results in extremely short effective contact times, on the order of from a few milliseconds to less than 2 seconds, since heat transfer is not limited either on heating or cooling by the resistance of tube walls as is the case in conventional ethylene production units.
This residence time is critical. It is critical that the alkane quench contact time be as brief as possible, preferably under 2 seconds, i.e. from 0.02 to 2 seconds. It is believed that all reaction occurs in under 1.5 seconds, with most of the reaction occurring in milliseconds, i.e. 0.02-0.5 second, and it is preferred that residence time not exceed 1.5 seconds.
The residence time of the alkane quench material within the reactor will depend upon the configuration of the reactor and other conditions which are well known and within the skill of the art.
The following examples illustrate the effects of various parameters on the practice of the invention through computer simulation of a process carried out in a reactor of 5"
I.D. (0.127 m) with a length of 200 m, with appropriate scale-up to the respective amounts of feed indicated in each example, to more accurately model the commercial application of the 27,418 2~
invention. It is to be understood, however, that the benefits and effects of the practice of the invention are largely independent of the configuration of the reactor used. Thus, the applicability and utility of the invention and the principles enunciated herein are not to be construed as limited to such a reactor configuration. A flow rate of 2000 kg/hr at 1300C was assumed for the effluent of the partial oxidation reactor. The reactor is operating at 55 atm. except in Example 6 where the pressure is 30 atm. At the inlet of the quench reactor, this stream is instantaneously and perfectly mixed with a quench stream of known composition, temperature and flow rate. The model simulates the changes in composition and temperature from thi point onwards. The quench reactor is assumed to operate at a specified constant pressure. No pressure drop is taken into account. All s-mulations were carried out assuming an adiabatic operation. An overall material balance model was then constructed with product recovery, purge and recycle streams to simulate a commercial facility.
Table 1 sets forth the mole % and weight % of effluents A and B used for purposes of Examples l Effluent A Effluent B
mole % wt % mole % wt %
H2 53.19 8.20 34.28 3.76 25 CH4 0.86 1.05 0.05 0.04 C0 29.07 62.24 39.37 60.08 C2 2.65 8.91 7.29 17.47 H2O 14.23 19.60 19.01 18.65 Table 2 sets forth the linear relationship between the hydrocarbon quench and the water quench needed to reach a target temperature after mixing. The amounts of water and hydrocarbon quench are expressed as a weight ratio relative to the amount of partial oxidation reactor effluent.
Table 2 Effluent A Effluent B
27,418 TarqetHC quench a b a b 700Cpropane 0.814-1.819 0.602 -1.819 800Cpropane 0.571-1.611 0.421 -1.611 900Cpropane 0.392-1.455 0.289 -1.455 900Cethane 0.384-1.424 0.283 -1.424 (weight ratio hydrocarbon quench)= a~b * (weight ratio water quench) Example 1 Following partial oxidation of 2000 kg/hr of effluent A feed at a temperature of 1300C and 55 atm., the reactor effluent is quenched with 785 kg/hr of propane at 25C, a mixing temperature of 9o0C, and a residence time of 1.331 sec. The final quench effluent temperature is 710C. The reactor inlet composition after mixing and the quench effluent composition after residence times of 0.065, 0.197, 0.995 and 1.331 (final) seconds are set forth below in Table 3.
Table 3 Component Wt % After Mixinq Wt % After Ouench/Sec 0.065 0.197 0.996 1.331 H2 5.896.388 6.433 6.478 6.483 CO 44.7044.697 44.697 44.697 44.697 25 CO2 6.406.399 6.399 6.399 6.399 H2O 14.0814.075 14.075 14.075 14.075 CH4 0.753.397 3.881 4.583 4.703 C2H4 4.158 5.090 5.752 5.934 C2H6 --0.512 0.805 1.027 1.045 30 C3H6 11.133 12.359 13.648 13.793 C3H828.19 8.961 6.319 3.032 2.548 n-C4H10 -- 0.114 0.158 0.230 0.243 iso-C4H10 0.066 0.072 0.079 0.080 27,418 Example 2 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 250 kg/hr of 100% propane at a temperature of 25C and 368 kg/hr of water at 25C. The final quench temperature is 815C. The reactor inlet composition after mixing and the quench effluent composition after residence times of 0.060, 0.182, 0.991 and 1.215 second, respectively, are set forth below in Table 4.
Table 4 Component Wt % After Mixina Wt % After Ouench/Sec 0.060 0.182 0.911 1.215 H2 6.266.480 6.487 6.488 6.488 JO 47.5547.548 47.548 47.548 47.548 15 CO2 6.816.807 6.807 6.807 6.807 H2O 29.0329.030 29.030 29.030 29.030 CH4 0.80 2.144 2.325 2.433 2.441 C2H4 2.577 2.989 3.305 3.333 C2H6 -- 0.139 0.120 0.067 0.057 20 C3H6 4.331 4.357 4.234 4.216 C3H8 9 55 0.882 0.268 0.017 0.008 n-C4H10 -- 0.035 0.040 0.042 0.042 iso-C4HlO 0.027 0.029 0.030 0.030 Example 3 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 1141 kg/hr of 100%
propane at a temperature of 25C. The final quench temperature of 642C is achieved after 1.375 sec. The reactor inlet composition after mixing the quench effluent compositions at 0.067, 0.202, 1.208 and 1.375 seconds, respectively, are set forth below in Table 5.
27,418 Table 5 Component Wt % AftPr M _inq Wt % After Ouench/Sec 0.067 0.2Q2 1.028 1.375 H2 5.225.599 5.673 5.769 5.784 CO 39.6339.631 39.631 39.631 39.631 C2 5.675.673 5.673 5.673 5.673 H2O 12.4812.480 12.480 12.480 12.480 CH4 0.672.133 2.512 3.094 3.201 10 C2H4 2.253 2.716 3.336 3.440 C2H6 --0.355 0.566 0.981 1.069 C3H6 8.342 10.191 12.767 13.209 C3H8 36.3323.476 20.485 16.175 15.414 n-c4Hlo --0.022 0.032 0.047 0.050 15 iso-c4Hlo 0.035 0.041 0.Q48 0.049 Example 4 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 250 kg/hr of 100% propane at a temperature of 25C and 553 kg/hr of water at 25C.
Temperatures of approximately 730C are achieved after a residence time of 1.25 sec. The reactor inlet composition after mixing and the quench effluent composition at 0.062, 0.185, 0.939 and 1.253 seconds are set forth below in Table 6.
Table 6 .Com~onent Wt % After Mixinq Wt % After Ouench~Sec 0.062 0.186 0.939 1.253 H2 5.855.993 6.024 6.050 6.052 30 CO 44.4144.410 44.410 44.410 44.410 C2 6.366.357 6.357 6.357 6.357 H2O 33.7133.714 33.714 33.714 33.714 CH4 0.75 1.384 1.630 1.989 2.046 C2H4 - 1.072 1.464 2.064 2.168 35 C2H6 0.081 0.125 0.159 0.156 C3H6 3.041 3.740 4.319 4.360 C3H8 8.92 3.923 2.495 0.873 0.668 n-C4H10 -- 0.013 0.025 0.045 0.048 iso-C4H10 0.012 0.016 0.020 0.021 Example 5 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 250 kg/hr of 100% ethane at a temperature of 25C and 364 kg/hr of water at 25C. The 27,418 reactor inlet composition after mixing and the quench effluent composition after residence times of 0.060, 0.181, 0.917 and 1.225 (final), respectively, are set forth below in Table 7.
Table 7 5 Component Wt % After Mixinq Wt % After Quench/Sec 0.060 0.181 0.gl7 1.225 H2 6.276.546 6.616 6.706 6.720 CO 47.6247.621 47.621 47.621 47.621 C2 6.8260817 ~.817 6.817 6.817 10 H2O 28.9228.921 28.921 28.321 28.921 CH4 0.80 0.914 0.963 1.052 1.070 C2H4 3.790 4.760 6.014 6.211 C2H6 9.56 5.194 4.019 2.431 2.171 C3H6 -- 0.002 0.003 0.005 0.005 n-C4H10 -- 0.190 0.276 0.430 0.462 i s o- C4 H 10 Example 6 20 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 250 kg/hr of 100% propane at a temperature of 25C and 553 kg/hr of water at 25C and a pressure of 30 arm. Temperatures of approximately 730C are achieved after a residence time of 0.68 sec. The reactor inlet composition after mixing and the quench effluent composition at residence times of 0.033, 0.101, 0.510 and 0.681 seconds (final), respectively, are set forth below in Table 8.
Table 8 Component Wt % After Mixinq Wt % After Ouench/Sec 30 0.033 0.101 0.510 0.681 H2 5.855.968 6.007 6.045 6.049 CO 44.4144.410 44.410 44.410 44.410 C2 6.366.357 6.357 6.357 6.357 H2O 33.7133.714 33.714 33.714 33.714 35 CH4 0.75 1.223 1.461 1.832 1.896 C2H4 -- 0.841 1.247 1.881 1.995 C2H6 -- 0.043 0.073 0.111 0.114 C3H6 2.454 3.287 4.104 4.188 C3H8 8.92 4.977 3.422 1.505 1.233 40 n-C4H10 0.005 0.011 0.024 0.027 isO-c4Hlo 0.008 0.012 0.016 0.017 27,418 Example 7 Following the process of Example 1, 2000 kg/hr of effluent B at 1300C is quenched with 578 kg/hr of 100% propane at a temperature of 25C and 0 kg/hr of water. The reactor inlet composition after mixing and the quench effluent composition after residence times of 0.091, 0.368, 1.393 and 1.860 sec., respectively, are set forth below in Table 9. The quench effluent temperature after each time interval is 747, 726, 711 and 708C, respectively.
Table 9 Component Wt % After Mixing Wt % After ~uench/Sec 0.091 0.368 1.393 1.860 H2 2.923.327 3.367 3.391 3.394 Co 46.6146.610 46.610 46.610 46.~10 15 CO2 13.5513.5S3 13.553 13.553 13.553 H20 14.4714.469 14.469 14.469 14.469 CH4 0.032.263 2.753 3.205 3.298 C2H4 3.472 4.116 4.764 4.913 C2H6 --0.543 0.731 0.839 0.842 20 C3H6 9.208 10.329 11.005 11.089 C3H8 22.40 6.398 3.862 1.902 1.558 n-C4H10 -- 0.103 0.150 0.199 0.209 isO-c4Hlo 0.054 0.060 0.064 0.065 Example 8 Following the process of Example 1, 2000 kg/hr of effluent B at 1300C is quenched with 250 kg/hr of 100% propane at a temperature of 25C and 225 kg/hr of water at 25C. The reactor inlet composition after mixing and the final quench effluent composition for residence times of 0.129, 0.345, 1.297 and 1.729 seconds, respectively, are set forth below in Table 10. Temperatures at each residence time are 799, 794, 792 and 791C, respectively.
27,418 Table 10 Component Wt % After Mixinq Wt % After Ouench~Sec _ 129 0.345 1.297 1.729 H2 3.043.266 3.270 3.271 3.271 5 CO 48.5548.549 48.549 48.549 48.549 C2 14.1214.117 14.117 14.117 14.117 H2O 24.1824.162 24.162 24.162 24.162 CH4 0.03 1.535 1.686 1.780 1.790 C2H4 2.735 3.066 3.321 3.362 10 C2H6 0.150 0.121 0.069 0.059 C3H6 ~~ 4.698 4.702 4.608 4.589 C3H8 10.10 0.700 0.230 0.025 0.014 n-C4H10 0.055 0.061 0.063 0.063 isO-c4Hlo ~~ 0.032 0.034 0.035 0.035 Example 9 Following the process of Example 1, 2000 kg/hr of effluent B at 1300C is quenched with 842 kg/hr of 100% propane at a temperature of 25C and 0 kg/hr of water. The reactor inlet composition after mixing and the quench effluent composition for residence times of 0.141, 0.331, 1.438 and 1.922 seconds, respectively, are set forth below in Table 11.
Temperatures at each of the above residence times are 685, 667, 641 and 636~C, respectively.
table 11 Component Wt % After Mixina Wt % After Quench/Sec 0.141 0.331 1.438 1.922 H2 2.652.997 3.041 3.108 3.120 CO 42.2842.280 42.280 42.280 42.280 30 C2 12.29~2.294 12.294 12.294 12.294 H2O 13.1213.125 13.125 13.125 13.125 CH4 0.031.436 1.681 2.120 2.209 C2H4 2.101 2.381 2.827 2.911 C2H6 --0.406 0.561 0.896 0.972 35 C3H6 ~~7.856 8.999 10.874 11.231 C3H8 29.6317.450 15.573 12.394 1~.774 n-C4H10 --0.023 0.030 0.041 0.044 iso-C4H10 0.032 0.035 0.040 0.041 27,418 Example 10 Following the process of Example 1, 2000 kg/hr of effluent B at 1300~C is quenched with 250 kg~hr of 100% propane at a temperature of 25C and 368 kg/hr of water at 25C. The reactor inlet composition after mixing and the final quench effluent composition for residence times of 0.131, 0.362, 1.328 and 1.774 seconds, respectively, are set forth below in Table 12. Temperatures at each residence time are 734, 722, 710 and 708C, respectively.
Table 12 Component Wt After Mixing Wt % After Ouench/Sec 0.131 0.362 1.328 1.774 H2 2.873.036 3.059 3.079 3.082 CO 45.9045.898 4~.898 45.898 45.898 15 CO2 13.3513.346 13.346 13.346 13~346 H2O 28.3028.304 28.304 28.304 28.304 C~4 0.030.800 1.006 1.288 1.346 C2H4 ~~1.242 1.547 1.980 2.078 C2H6 -- 0O135 0.189 0.240 0.242 20 C3H6 3.575 4.132 4.629 4.692 C3~8 9 553.629 2.467 1.162 0.993 n-C4H10 -- 0.021 0.034 0.054 0.058 iso-C4H10 0.015 0.017 0.020 0.021 Example 11 Following the process of Example 1, 2000 kg/hr of effluent B at 1300C is quenched with 250 kg/hr of 100% ethane at a temperature of 25C and 222 kg/hr of water at 25DC. The reactor inlet composition after mixing and the quench effluent composition for residence times of 0.127, 0.341, 1.295 and 1.732 seconds, respectively, are set forth in Table 13.
Temperatures at each residence time are 808, 790, 769, and 765DC, respectively.
27,418 Table 13 COmDOnent Wt After Mixina Wt % After Ouench/Sec 0.127 0.341 1.295 1.732 H2 3 . 043 . 326 3. 381 3 . 448 3 . 461 5 CO 48.6148.608 48.608 48.608 48.608 C2 14.1314.134 14.134 14 . 134 14.134 H2O 24.0724.070 24.070 24.070 24.070 CH4 0.03 0.170 0.214 0.284 0.301 C2H4 3.959 4.712 5.646 5.835 10 C2H6 10.11 5.487 4.687 3.362 3.113 C3H6 - 0.002 0.003 0.004 0.004 C3H8 0.003 0.003 0.002 0.002 n-C4H10 -- 0.239 0.318 0.442 0.472 15 iso-C4H10 -_ 0.000 0.000 0.000 0.000 In a representative partial oxidation process which does not employ an alkane quench, natural gas at an initial pressure of 300 psig is compressed, preheated and injected into a gasifier, operated at a pressure of about 800 psi, with steam, where it is partially oxidized at 1300-1400 C by combustion with 99.5~ oxygen from an air separation plant. The hot gases enter the water heat recovery section where they are cooled by heat exchange to produce 1500 psi saturated steam, then further cooled in a water scrubbing section. The gases then flow to an amine scrubbing unit for removal of CO2. The C2 content of the untreated gas is about 3% and it may not be necessary to remove in some applications, such as methanol synthesis.
27,418 Typical design and performance parameters for the partial oxidation of natural gas are as follows:
Oxygen/natural gas, ~wt/wt) 1.29 Steam/natural gas, (wt/wt) 0.20 Exit gas pressure, psi 800 Exit gas temperature, F 2500 Dry gas produced, scf/lb natural gas 62.0 CO + H2 produced, scf/lb natural gas 59.4 CO + H2 produced, mol/mol oxygen 3.89 CO + H2 produced, mol/mol steam 14.1 The feed preheat system includes a zinc oxide sulfur guard bed to assure that any catalysts in downstream synthesis gas processing steps are protected from sulfur contaminants. A
portion of the 1500 psi saturated steam is used to preheat the oxygen feed and to inject Examples 12-15 Use of Alkane Ouench with Partial Oxidation of 1762 Pounds of Methane Following the general procedure described above, the partial oxidation of methane with no alkane quench (Example 12); 785 kg propane quench at 25C per 2000 kg raw syngas (Example 13); 1141 kg propane quench at 25C per 2000 kg raw syngas (Example 14); and 250 kg propane + 268 kg water quench at 25C per 2000 kg raw syngas (example 15) produced the results set forth in Table 14.
Table 14 Time Temp Temp (H2-C2L _2H4 Eauiv- Lb. C3H8 Onch. Lb (Sec) mix C ex C H2~ (CO+C021 Svnaas Alkane Net Gross Net Examule 12 0.065 1300 1300 1.8262 1.3900 1255 ---- 1255 Base No Qnch Example 13 35 0.065 900 752 1.9810 1.7321 1361 939 2300 1731 1180 0.197 900 734 1.~950 1.7449 1371 1054 2425 1731 1342 0.996 900 713 2.009 1.7577 1378 1191 2569 1731 1545 1.331 900 710 2.0105 1.7591 1378 1211 2589 1731 1575 27,418 Example 14 0.067 800 698 1.9583 1.7113 1346 734 2080 2515 890 0.202 800 676 l.g842 1.7351 1363 8~4 2257 2515 1097 1.028 800 647 2.0178 1.7658 13~2 1115 2497 2515 1395 1.375 800 642 2.0230 1.7706 1385 1153 2538 2515 1448 Example 15 0.060 9Q0 824 1.8891 1.6479 ~2~8399 1697 551 500 10 0.182 900 818 1.8911 1.6497 1299424 1723 551 536 0.911 900 816 1.8914 1.6500 1300435 1735 551 550 1.125 900 815 1.8914 ~.6500 1300436 1736 551 551 In a representative reforming process not employing 15 an alkane quench, natural gas feed and steam are heated and passed through catalyst-filled tubes in the reformer furnace. Natural gas is burned in the furnace to supply the endothermic heat of reaction and the sensible heat to the hot product gases. A heat recovery system which 20 recovers heat from the furnace flue gases and the reformed gases to preheat the feed and generate high pressure superheated steam is employed. A zinc oxide guard bed is included in the feed preheat system to ensure that the reformer catalyst is protected from sulfur contamination.
The gases are further cooled to condense the unreacted steam, processed through an amine unit to remove CO2, and compressed to 750 psig for delivery.
The high pressure, superheated steam is expanded through a turbine to supply power for the gas compressors 30 and intermediate pressure steam for the reformer furnace.
Low pressure steam is generated by heat recovery from product gas and flue gas to provide reboil steam for the C2 removal unit.
The synthesis gas has a H2/CO ratio of 5.4 and 3S contains 3.37% of methane and 0.1% of nitrogen.
27,41~
2 g Typical design and performance parameters are as follows:
Steam/carbon, moles/atom3.30 Steam/natural gas feed, wt/wt 3.68 Reformer exit pressure, psia 275.00 Reformer exit temperature, F 1560.00 Dry gas produced, scf/lb feed 91.9 CO + H2 produced, scf/lb feed 80.9 CO + H2 produced, mol/mol steam 1.04 Examples 16-20 Use of Alkane Ouench with Reforming 936 Pounds Methane Employing a natural gas which is low in CO2 and nitrogen content, and following the general procedure described above, reforming of 936 pounds of methane with no alkane quench (Example 16); 200 kg propane quench at 25C per 2000 kg raw syngas (Example 17); 500 kg propane quench at 25C per 2000 kg raw syngas (Example 18); 500 kg propane quench at 600C per 2000 kg raw syngas (Example 19); and 200 kg propane quench at 600C per 2000 kg raw syngas (Example 20) produced the results set forth in Table 15.
Table 15 Time Temp Temp (H2-C2L C2H4 Equiv- Lb- C3Hg Onch. Lb (Sec) mix,C ex,~C H2/CO ~CO+CO21 Synqas Alkane Net Gross Net Example 16 -- 857 857 5.3751 2.6931 918 ---- 918 Base No Qch Example 17 0.654 779 720 5.5232 3.0004 918 284 1202 441 348 Example 18 0.682 690 741 5.5340 3.0070 918 270 1188 1102 317 27,418 Example 19 0.633 793 694 5.6920 3.1039 982 594 1576 1102 730 5 Example 20 0.628 827 759 5.5377 3.0092 960 329 1289 441 409 In a typical autothermal reforming process, using the typical design and performance parameters for the autothermal reforming of natural gas set forth in Table 16 below, the natural gas feed and steam, preheated to about 1000F, and 99.5% oxygen preheated to about 500F, are fed into catalyst filled reactors where reforming is carried out. The feed preheat system includes a zinc oxide guard bed to insure that downstream catalysts are protected from sulfur contamination.
The exit gases at about 1750F enter the heat recovery section where they are first cooled by generating high pressure unsaturated steam. The gases are further cooled to condense unreacted steam, fed to an amine unit to remove CO2,. and compressed to a pressure of 750 psig. The saturated steam, superheated in a natural gas fired unit, is expanded through turbines to provide power before being injected in the gasifier or used for stripping steam in the CO2 removal unit. The amount of steam generated is in balance with the steam demands, although some electric power is purchased.
Table 16 Desiqn and Performance Parameters Steam/carbon, mols/atom 2.10 Steam/natural gas, wt/wt 2.34 Oxygen/natural gas, wt/wt 1.10 Reformer exit pressure, psia 290 Reformer exit temperature, F 1750 Dry gas production, scf/lb natural gas 73.9 CO + H2 production, scf/lb natural gas 65.1 27,418 3 !~
Co + H2 production, mol/mol oxygen 4.98 CO+ H2 production, mol/mol steam 1.32 Examples 21-22 Use of Alkane Ouench in Autothermal Reforminq of 984 Pounds of Methane Following the general procedure described above, autothermal reforming of 984 pounds of methane with no alkane quench (Example 22); and as modified by the process of the 10 present invention using 500 kg of propane quench at 600C per 2000 kg raw syngas to first cool the exit gases (Example 23) produced the results set forth in Table 17.
Table 17 Time Temp Temp (H2-C2L C2~4 Equiv- Lb. C3H~ Onch. Lb (Sec) mix C ex.C H2/CQ rCO+CO21 Syngas Alkane Net Gross Net Example 22 --- 940 940 3.3781 1.7988 793 Base 793 Base No qch Example 23 0.069 848 746 3.6138 1.9496 836 539 1375 1102 657 25 0.208 848 729 3.6476 1.9712 842 634 1476 1102 782 0.348 848 720 3.6610 1.9800 844 674 1518 1102 836 0.686 848 712 3.6751 2.0000 8~7 722 1569 1102 902 It will be understood by those skilled in the art that the optimum amount and type of quench for the particular situation, within the parameters of this invention, will be determined by local feedstock availability, markets, end products desired, and the like.
The above discussion is intended to be only illustrative of the invention. The full spirit and scope of the invention should be determined by reference to the following claims.
27,418
Field of the Invention This invention relates to improved processes for manufacturing syngas and higher hydrocarbons and oxygenates from methane and/or natural gas, and more particularly relates to a process wherein high quality heat wasted in prior art processes is captured, and additional hydrogen is provided for use in converting low value carbon oxides to more vaiuable products.
Backqround of the Invention Syngas is often manufactured as an intermediate in the conversion of feedstocks containing carbon and hydrogen to lS more desired hydrocarbons or oxygenates. This technique is particularly useful as a means for converting methane to higher hydrocarbons or oxygenates which are more readily transportable from the field and/or more desirable products such as alcohols.
A major source of methane is natural gas which typically contains about 40-95% methane depending on the particular source. Other constituents include about 10% ethane with the balance being made up of CO2 and smaller amounts of propane, butanes, pentanes, nitrogen, etc.
Primary sources for natural gas are reservoirs either alone or generally associated with hydrocarbon liquid reserves.
Most of the natural gas used for heating purposes comes from these sources. Quantities of natural gas are also known to be present in coal deposits and are by-products of crude oil refinery processes and bacterial decomposition of organic matter.
Prior to commercial use, natural gas must be processed to remove water vapor, condensible hydrocarbons and inert or poisonous constituents. Condensible hydrocarbons are generally removed by cooling natural gas to a low temperature, 27,418 and then washing it with a cold hydrocarbon liquid to absorb the condensible hydrocarbons. The condensible hydrocarbons are typically ethane and heavier hydrocarbor.s. This gas processing can occur at the wellhead or at a central processing station.
Processed natural gas typically comprises a major amount of methane, and minor amounts of ethane, propane, butanes, pentanes, carbon dioxide and nitrogen. Generally, processed natural gas comprises from about 70% to more than about 95% by volume of methane.
Most processed natural gas used commercially is distributed through extensive pipeline distribution networks.
As natural gas reserves in close proximity to gas usage declines, new sources that are more remote require transportation over further distances. The gas from many of these distant sources are not, however, amenable to transport by pipeline. For example, the gas from sources that are located in areas requiring economically unfeasible pipeline networks or in areas requiring transport across large bodies of water are not amenable to transport by pipeline. This problem has been addressed in several ways. One such solution has been to build a production facility at the site of the natural gas deposit to manufacture one specific product. Another approach has been to liquefy the natural gas and transport it in specially designed tanker ships. Natural gas can be reduced to 1/600th of the volume occupied in the gaseous state by cryogenic processing, and with proper procedures, safely stored or transported. These processes, which involve liquefying natural gas, transporting the liquified gas, and revaporizing it are complex, energy intensive and expensive.
Still another approach has been the conversion of natural gas to higher molecular weight hydrocarbons or oxygenates, preferably substantially liquid hydrocarbons or oxygenates, that can be easily handled and transported. The conversion of natural gas to higher order hydrocarbons or 27,418 -2~J~
oxygenates, especially ethane and ethylene, methanol or higher alcohols retains the material's versatility for uses as precursor materials in chemical processing. Known dehydrogenation and polymerization processes are available for the further conversion of ethane and ethylene to liquid hydrocarbons and of alcohols to numerous valuable products. In thase ways, easily transportable commodities may be derived from natural gas at the wellhead. A drawback in implementing such processes has been the lack of means for obtaining a sufficiently economical conversion rate of natural gas to higher molecular weight hydrocarbons.
To date, processes to convert natural gas to hydrocarbons or alcohols rely on the manufacture of a synthesis gas or syngas (carbon monoxide and hydrogen) with a subsequent conversion of syngas to the desired product. Examples of this category of processes included Shell Oil's Shell Middle Distillate (SMDS) process, the Gulf Badger Fischer-Tropsch process, the Mobil Oil MTG and MOGD processes, commercially available processes for the manufacture of methanol, and a variety of processes proposed for the manufacture of higher alcohols.
When methanol or other alcohols are the desired products or intermediates, or water the desired by-product in the production of aliphatic or mono-olefinic hydrocarbons or oxygenates from syngas, the raw syngas should theoretically have the following molar relationship:
(H2-C02)/(CO2-~CO) 2 This ratio will be referred to henceforth as the H2/COx ratio.
In practice for instance, a H2/COx ratio of 2-2.1 (mole/mole) is often used for the production of methanol. Since the H2/COx ratio in the described products is always less than or equal to the maximum possible from a methane feed (e.g. the ratio for methane and methanol is equal), water is the usual by-product of choice.
27,418 ~$~
While natural gas is a convenient source, syngas may also be derived from streams produced from the burning of coal and streams formed during numerous manufacturing or processing operations and in the above- or below-ground gasification or liquefication of coal, tar sands, oil shale and biomass, for example.
Several techniques are used to manufacture syngas, including partial oxidation, reforming and a process which can be viewed as a hybrid of the two, commonly referred to as autothermal reforming.
Partial oxidation relies on the partial combustion of hydrocarbons with an oxygen-containing gas to produce a high temperature mixture of carbon monoxide, hydrogen, water and carbon dioxide. When methane is used as the hydrocarbon feed, a hydrogen to carbon monoxide ratio of between 1.8 to 1.9 is the usual result, with the ratio being at least partially dependent on the composition of the feed stream. Carbon dioxide in the feed stream, for instance, will produce a somewhat lower ratio, and hydrogen or water in the feed stream will have a tendency to increase the ratio.
One way to achieve the desired H2/COX ratio is to treat the syngas with a shift catalyst to increase the hydrogen content at the expense of carbon monoxide according to the following formula:
H2O + CO = H2 +C2 Water is then rejected, thus resulting in a higher H2/COX ratio syngas. However, this technique has a detrimental side effect, since carbon dioxide concentrations well in excess of those needed for optimum H2/COX ratios are required to force the above equilibrium to the left. This requirement makes the approach relatively inefficient.
Classical reforming overcomes some of the H2/COX
ratio limitations encountered in partial oxidation, but presents other drawbacks. More of the available feed must be 27,418 h diverted to the reformer furnace to provide the heat to drive the idealized endothermic reforming reaction:
CH4 + H20 3H2 + CO
This feed derived fuel is generally burned with air, and the carbon it contains is lost as carbon oxides vented to the atmosphere, rather than to increase product yield. Thus, although thP H2/COX ratio in the raw syngas is greater than that necessary to achieve a stoichiometric optimum, that increase is offset by the decrease in yield attributable to vented combustion products.
In general, the raw syngas exits from the reformer at a lower temperature than is typical in either partial oxidation or autothermal reforming because of the metallurgical limitations of the tubes used in the reforming furnace.
Autothermal reforming is the means by which the industry through the use of current technology has attempted to balance the extremes in trade-offs between partial oxidation, with its low H2/COx ratio and reduced use of feed as fuel, and classical reforming, with its high H2/COX ratio and increased use of feed as fuel. As practiced, autothermal reforming consists of partially reforming a hydrocarbon such as methane, but not at the extremes practiced in classical reforming, and at lower fuel duty. The resultant effluent stream obtained through autothermal reforming is relatively high in hydrogen and also high in methane. This stream is subsequently mixed with oxygen or oxygen-containing gas and further partially combusted to reduce its methane and hydrogen content to produce a product stream having an improved H2/COX ratio. In addition, feed requirements as fuel are reduced.
Representative processes are disclosed in, for example, United States Patent Nos. 3,594,425; 3,631,111 and 4,390,314.
An important limitation and inefficiency of prior art processes, including partial oxidation, reforming and 27,418 autothermal reforming, is the method of heat removal from the processing units. The reactions are conducted under severe conditions, i.e., temperatures which may be in excess of 2000F. Following partial oxidation or other processes, the effluent stream must be rapidly cooled or quenched to a temperature at which the occurrence of retrograde reactions can be reduced. Typically the temperature of the reactor effluent must be lowered to 700C or even 600C to reduce the formation of carbonaceous deposits, the formation of undesirable aliphatic compounds from desirable olefins, etc. This temperature reduction generally is accomplished in from less than one second to a few seconds by quenching with water or other heat absorbing medium. This approach has major drawbacks which the present invention overcomes. Specifically, valuable high quality (high temperature) heat is wasted, and low quality (low temperature heat, which must then be disposed of at significant cost, is increased.
We have found that the thermal efficiency, and thus the economic attractiveness of syngas-based processes can be improved by the use of an alkane quench. The quenching of the high temperature partial oxidation reactor effluent, as well as the effluent from classical reforming and autothermal reforming processes, with a stream containing alkanes, is a means of efficiently raising the H2/CO~ ratio of the resultant stream without causing a detrimental rejection of carbon dioxide.
Thermal or physical quenching has heretofore been employed in prior art cracking processes involving gross volumes of heavy hydrocarbon feeds where heat is fed into the system to promote feedstock cracking, after which the reaction is quenched to prevent retrograde reactions. Generally speaking, prior art processes that employ hydrocarbons as a heat sink are doing so as a mechanism to reduce the process stream volume and the size of heat recovery equipment downstream which is required if water vaporization is used to 27,418 absorb heat and lower stream temperatures. Such prior art techniques, however, including equipment are those described in United States Patent Nos. 4,520,224, 4,288,408, 4,556,749 and 4,384,160, all rely on specific heat capacity, phase change, heat of vaporization, and the like to dilute the concentration of heat thus lowering stream temperature. These prior art processes are applicable to subsequent unit operations, i.e., those performed downstream of the process of the present invention. The prior art processes produce lower quality heat downstream resulting in the requirement of more heat rejection equipment in the form of cooling towers, vaporizers, heat exchangers and the like. Further, the prior art cracking processes which employ some form of thermal or physical quenching, including those that employ a hydrocarbon quench medium such as recycled heavy oil, produce methane in a form only useful as process fuel, the very material the present invention is designed to convert to more transportable, higher hydrocarbons.
United States Patent No. 3,022,148 discloses the use of an organic quench in the non-catalytic partial oxidation of hydrocarbon oil at a temperature of approximately 2500F, specifically, a hot oil stream wherein the oil has an average boiling range of from 500-750F. The requirement of the use of a high boiling point oil stream as a quench medium is not only energy intensive and therefore very costly, but impractical from the standpoint of converting methane to higher hydrocarbons or oxygenates at remote sites where the heavy, high boiling points oils would have to be transported in for processing.
The use of an alkane quench in syngas based processes, on the other hand, has a number of advantages. The product yield, thermal efficiency, and hence economic attractiveness is improved by the alkane quench techniques of 27,418 the present invention. Further, the quench material is likely to be readily available for processing at the site.
Thus, the present invention overcomes drawbacks of the prior art and improves the economic feasibility of both catalytic and non-catalytic partial oxidation, classical reforming and autothermal reforming processes.
Summary of the Invention The improved processes of this invention for producing syngas comprises contacting a hydrocarbon feedstock with an oxygen-containing material in a reaction zone of a processor to produce an effluent stream comprising reaction products consisting essentially of COx H2, and optionally c2+
hydrocarbons or oxygenates; and removing heat from the effluent stream by briefly contacting hot reactor effluent with a C3-C20 alkane quench material having a temperature of up to lQ00C to produce a chemically quenched stream wherein part of the heat contained in the raw reactor effluent has been removed through endothermic cracking reactions. The temperature of the chemically quenched effluent stream is further reduced by thermal means, using, for example, water, recirculating refractory oils, heat exchangers, or other heat absorbing means to achieve a temperature which further discourages or retards retrograde reactions. Thereafter, water, C4+ hydrocarbons and like byproducts are removed, and the effluent stream is processed in a synthesis reactor and product recovery system via methods well known in the art whereby oxygenates and hydrocarbons are formed, separated and stored. The process may be carried out in a partial oxidation, classical reforming or autothermal reforming reactor.
In another embodiment, streams comprising methane, light hydrocarbons, carbon oxides, and hydrogen are converted to C2+ hydrocarbons and oxygenates via pyrolysis at 1500 D F
employing an alkane quench to lower the temperature of the 27,418 reactor effluent stream as described above, and thereafter processing said stream by methods well known in the art.
More particularly, in one embodiment, the process of the present invention provides a method for converting a methane rich feedstock to syngas and olefins comprising the steps of: reacting the feedstock with an oxygen-containing material in a reaction zone of an partial oxidation reactor to produce an effluent stream comprising reaction products consisting essentially of COx, H2, and C2+ hydrocarbons; and intxoducing an alkane quench material having a temperature of up to 1000C to produce a chemically quenched effluent stream having a reduced temperature. Thereafter, the process of this invention may comprise the additional steps of further reducing the temperature of the chemically quenched stream by thermal means to retard retrograde reactions; and separating and removing C2+ hydrocarbon, water, hydrogen and carbon oxide products from the product stream. If desired, at least a portion of unreacted methane, unreacted alkane quench material, alkanes formed during chemical quenching, and mixtures thereof contained in the product stream may be recycled to the reaction zone. In addition, at least a portion of the C2+ hydrocarbons may be reacted to form higher homologs, oxygenates and alcohols.
C2+ hydrocarbons may be separated and removed from unreacted methane either from the effluent stream following the chemical quench or from the product stream following the thermal quench.
The hydrocarbon quench of this invention provides a number of major benefits. First, it absorbs some of the heat otherwise removed by wasteful techniques through an extremely fast endothermic reaction. Second, at least a portion of the hydrocarbon quench is converted to reactive intermediate or valuable end products. A further advantage is the production of additional hydrogen that can be used to improve the value of 27,418 end products or increase yields by converting carbon oxides to valuable products or intermediates.
It is presently preferred to employ the process of this invention in the vicinity of a gas extraction plant.
Ethane, propane, butanes, other natural gas liquids, condensates, etc., which are presently several of the preferred quench media for use in the gasifier, can be isolated at low cost in many gas plants. Furthermore, in remote locations where methane conversion technologies are likely to be implemented, these components usually have a low value, and are expensive to transport to markets.
The present invention is also applicable to producing streams comprising syngas and olefins from syngas containing streams derived from other sources, such as from coal deposits and formation during numerous mining operations and in the above- or below-ground gasification or liquefaction of coal, tar sands, oil shale and biomass, for example.
Brief Description of the Drawina FIG. 1 shows a process flow scheme for a representative partial oxidation process leading to the production of C2+ products in accordance with this invention.
FIG. 2 shows a process flow scheme for a representative classical reforming process in accordance with this invention.
FIG. 3 shows a process flow scheme for a representative autothermal reforming process in accordance with this invention.
Detailed Description of the Invention Referring to FIG. 1, there is depicted one preferred embodiment in which partial oxidation and quench operations are accomplished in different zones of a reaction vessel. Other configurations for contacting hot reaction gases with hydrocarbon quench materials, such as separate units, may be 27,418 employed, alternative hydrocarbon feeds may be used, and alternative syngas processing techniques may be employed to result in alternative products.
Methane, illustrative of a feedstock such as natural gas, is introduced into the system by line 12 where it is mixed with oxygen, air, or other oxygen-source gas supplied through line ll. The resulting mixture is fed into a first zone 14 of partial oxidation reactor 13 where it reacts to form syngas.
The hot gaseous feedstream passes on to hydrocarbon quench zone where it is mixed and reduced in temperature by the introduction of alkane quench material, after which the reactor stream progresses to final quench zone 16 at which point- a diluent such as water or oil is introduced to further cool the stream and prevent retrograde reactions.
Depending on the level of temperature reduction and downstream processing desired, several alternative flow configurations may be used, two of which are for the manufacture of representatiYe products such as Cl-C4 alcohols and are described below. If the effluent is cooled below the condensation point of water, the reactor effluent passes through line 17 into feed gas adjust zone 18 where some C2+
hydrocarbon product and water, etc. ara withdrawn through line 19 and C02 byproduct are withdrawn and recycled through line 20. Thy stream is then treated in synthesis reactor and product recovery system 30. Mixed alcohols (a mixture of Cl-C4 alcohols) are fed into mixed alcohol stabilization and storage zone 31 Yia line 32. The retaining product is fed into methanol reactor and product recovery system 34 via line 35.
Water is recovered from reactor 30 through line 36. Methane, SO hydrogen, and inert are separated and removed from reactor 34 via line 37a to prevent exces ive buildup of inerts. Carbon dioxide recovered from synthesis reactor 34 may be recycled via line 37 to the syngas generator. Methanol is fed to methanol 27,418 ~J~ 9~
product stabilization and storage unit 38 via line 39 and may also be recycled to synthesis reactor 30 via line 35a.
If desired, the process flow scheme depicted in FIG.
1 may be changed to allow for the production of intermediates, such as olefins (e.g., ethylene, propylene, etc.), oxygenates such as alcohols (e.g., methanol, ethanol, etc.) or other hydrocarbons the incorporation of carbonylation, hydroformylation, carbon oxide hydrogenation and other processes known to those skilled in the art to produce derivative products.
As shown in FIG. 2, in one embodiment of the present invention as applied to classical reforming of natural gas, natural gas and steam are fed into team reformer 40 via line 41. The effluent stream exitinq reformer 40 via line 42 is contacted with alkane quench material in quench zone 43 and subsequently fed via line 44 into gay cooler 45 where it is further cooled. The further cooled effluent stream exits gas cooler 45 via line 46 and is fed into C02 removal zone. C02 is removed via line 48 And the effluent stream is fed into compressors 50 via line 49 where it it converted to a high pre sure tree co~pri~ing synthesis gas and olefins.
PIG. 3 depicts one embodiment ox the present invention a applied to autothermal reforming. Natural gas and steam are fod into feed preheater via line 61. The preheated stream exits preheater 60 via line 62 where it is mixed with oxygen and fed into an autothermal reformer reactor 63. The effluent stream exiting reactor 63 via line 64 is cooled by contact with an alkane quench materlal in quench zone 65. The cooled stream is passed through gay cooler 67 via line 66.
Steam and water are removed via lines 68 and 69, respectively, and C02 ~ub3equently removed in zone 70. The stream then is fed into compressor zone 71 via line 72 and a high pressure stream comprising synthesis gas and olefin~ is recovered via line 73.
27,418 2 3 :~
As shown in FIGS. 2 and 3, the steam generated by the processes is employed to generate electricity.
The hydrocarbon quench employed in the practice of this invention may be a C2-C20, and preferably C3-C20 straight or branched chain alkane, or mixtures thereof.
The term "C2-C20 alkane", as used herein, refers to a straight or branched-chain alkane, i.e., ethane, propane, n-butane, 2-methylpropane, n-pentane, 2-methylbutane, 2,2-dimethylpropane, n-hexane, 2-methylpentane, 2,3-dimethylbutane, n-heptane, 2,3-dimethylpentane, n-octane, 3~ethyl-2,3-dimethylhexane, n-dodecane, n-octadecane, n-eicosane, and the like. Suitable quench materials also may include naphthas.
The term "naphthas", as used herein, refers to hydrocarbon product mixtures that distill at a boiling range below 180-190C.
It is preferred to employ paraffinic hydrocarbons as the quench material, more particularly light alkanes, i.e. C3-C5 alkanes. However, the presence of moderate amounts of aromatics, while not desirable, would not generally be detrimental to the process of this invention. Rather, aromatics would serve as diluents. However, the presence of naphthenic components are not desirable since they have a lesser tendency to produce desirable olefins. The term "naphthene~ or ~naphthenic components" refers to multiringed compounds containing both aromatic and saturated rings, also referred to as cycloalkylaryl compounds.
It is especially preferred to employ light hydrocarbon, preterably C3-C5 alkanes, in the practice of this inVQntiOn. However, as a practical matter, the actual selection of the hydrocarbon quench material will depend upon what is available in the field. Varying the selection of the alkane will vary the product distribution.
The use of an alkane quench material to cool the effluent stream causes endoth~rmic dehydrogenation and 27,41~
2~ Jo decomposition of said alkane quench material and a substantial increase in C2+ hydrocarbon products.
Generally speaking, the alkane quench comprises from 1 to 2S weight percent of the process stream just prior to the alkane quench addition, preferably from 5 to 20 weight percent and most preferably from 10-20 weight percent. The exact amount of quench material used will depend on the volume and composition of the quench material available, the conditions at which the partial oxidation, reforming or autothermal reforming segment is run, and the desired downstream composition.
The alkane quench material may be employed at ambient temperature, but is preferably heated to temperatures of up to 10007C, preferably, above 300C. In cases where there is a cryogenic unit at the site, temperatures as low as -lS0C may be employed. However, for best results, it is preferred to introduce the alkane quench material into the quench reactor at elevated temperatures as discussed above, most preferably, if feasible, at the temperature of the effluent stream. The optimum preheat temperature will be governed by the composition of the quench, and process conditions upstream and downstream of the quench operation.
The upper limitation on the temperature of the quench material is the thermal stability of the material. This will, in part, depend upon whether or not the material is diluted and the diluent. For example, if a quench material is diluted with carbon dioxide, steam or other diluents, it may be heated to a higher temperature without degrading.
The process of this invention may be run under a wide range of pressures, i.e. atmospheric to 1500 psi. It is preferable to operate at pressures of at least 20 psi, more preferably, from about S0 psi to 1000 psi.
The temperature at which the partial oxidation, reforming, autothermal reforming or pyrolysis reaction is run will depend upon the type of reactor, whether a catalyst is 27,418 2 so employed and if so, the particular catalyst employed.
Generally speaking, in the case of a partial oxidation operation, the temperature should be maintained at from about 1000-1450~C, depending upon the particular catalyst and equipment employed. In the case of pyrolysis, the temperature should be maintained from about 1300 to 2000C. For classical reforming, the temperature should be maintained at from about 800 to 950C and for autothermal reforming, from about 800 to 1250C.
It is critical that the temperature of the effluent stream leaving the reactor be maintained at a minimum of 600 D C
and preferably not drop below 700C for best results.
The use of an alkane chemical quench, immediately followed by a thermal quench results in extremely short effective contact times, on the order of from a few milliseconds to less than 2 seconds, since heat transfer is not limited either on heating or cooling by the resistance of tube walls as is the case in conventional ethylene production units.
This residence time is critical. It is critical that the alkane quench contact time be as brief as possible, preferably under 2 seconds, i.e. from 0.02 to 2 seconds. It is believed that all reaction occurs in under 1.5 seconds, with most of the reaction occurring in milliseconds, i.e. 0.02-0.5 second, and it is preferred that residence time not exceed 1.5 seconds.
The residence time of the alkane quench material within the reactor will depend upon the configuration of the reactor and other conditions which are well known and within the skill of the art.
The following examples illustrate the effects of various parameters on the practice of the invention through computer simulation of a process carried out in a reactor of 5"
I.D. (0.127 m) with a length of 200 m, with appropriate scale-up to the respective amounts of feed indicated in each example, to more accurately model the commercial application of the 27,418 2~
invention. It is to be understood, however, that the benefits and effects of the practice of the invention are largely independent of the configuration of the reactor used. Thus, the applicability and utility of the invention and the principles enunciated herein are not to be construed as limited to such a reactor configuration. A flow rate of 2000 kg/hr at 1300C was assumed for the effluent of the partial oxidation reactor. The reactor is operating at 55 atm. except in Example 6 where the pressure is 30 atm. At the inlet of the quench reactor, this stream is instantaneously and perfectly mixed with a quench stream of known composition, temperature and flow rate. The model simulates the changes in composition and temperature from thi point onwards. The quench reactor is assumed to operate at a specified constant pressure. No pressure drop is taken into account. All s-mulations were carried out assuming an adiabatic operation. An overall material balance model was then constructed with product recovery, purge and recycle streams to simulate a commercial facility.
Table 1 sets forth the mole % and weight % of effluents A and B used for purposes of Examples l Effluent A Effluent B
mole % wt % mole % wt %
H2 53.19 8.20 34.28 3.76 25 CH4 0.86 1.05 0.05 0.04 C0 29.07 62.24 39.37 60.08 C2 2.65 8.91 7.29 17.47 H2O 14.23 19.60 19.01 18.65 Table 2 sets forth the linear relationship between the hydrocarbon quench and the water quench needed to reach a target temperature after mixing. The amounts of water and hydrocarbon quench are expressed as a weight ratio relative to the amount of partial oxidation reactor effluent.
Table 2 Effluent A Effluent B
27,418 TarqetHC quench a b a b 700Cpropane 0.814-1.819 0.602 -1.819 800Cpropane 0.571-1.611 0.421 -1.611 900Cpropane 0.392-1.455 0.289 -1.455 900Cethane 0.384-1.424 0.283 -1.424 (weight ratio hydrocarbon quench)= a~b * (weight ratio water quench) Example 1 Following partial oxidation of 2000 kg/hr of effluent A feed at a temperature of 1300C and 55 atm., the reactor effluent is quenched with 785 kg/hr of propane at 25C, a mixing temperature of 9o0C, and a residence time of 1.331 sec. The final quench effluent temperature is 710C. The reactor inlet composition after mixing and the quench effluent composition after residence times of 0.065, 0.197, 0.995 and 1.331 (final) seconds are set forth below in Table 3.
Table 3 Component Wt % After Mixinq Wt % After Ouench/Sec 0.065 0.197 0.996 1.331 H2 5.896.388 6.433 6.478 6.483 CO 44.7044.697 44.697 44.697 44.697 25 CO2 6.406.399 6.399 6.399 6.399 H2O 14.0814.075 14.075 14.075 14.075 CH4 0.753.397 3.881 4.583 4.703 C2H4 4.158 5.090 5.752 5.934 C2H6 --0.512 0.805 1.027 1.045 30 C3H6 11.133 12.359 13.648 13.793 C3H828.19 8.961 6.319 3.032 2.548 n-C4H10 -- 0.114 0.158 0.230 0.243 iso-C4H10 0.066 0.072 0.079 0.080 27,418 Example 2 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 250 kg/hr of 100% propane at a temperature of 25C and 368 kg/hr of water at 25C. The final quench temperature is 815C. The reactor inlet composition after mixing and the quench effluent composition after residence times of 0.060, 0.182, 0.991 and 1.215 second, respectively, are set forth below in Table 4.
Table 4 Component Wt % After Mixina Wt % After Ouench/Sec 0.060 0.182 0.911 1.215 H2 6.266.480 6.487 6.488 6.488 JO 47.5547.548 47.548 47.548 47.548 15 CO2 6.816.807 6.807 6.807 6.807 H2O 29.0329.030 29.030 29.030 29.030 CH4 0.80 2.144 2.325 2.433 2.441 C2H4 2.577 2.989 3.305 3.333 C2H6 -- 0.139 0.120 0.067 0.057 20 C3H6 4.331 4.357 4.234 4.216 C3H8 9 55 0.882 0.268 0.017 0.008 n-C4H10 -- 0.035 0.040 0.042 0.042 iso-C4HlO 0.027 0.029 0.030 0.030 Example 3 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 1141 kg/hr of 100%
propane at a temperature of 25C. The final quench temperature of 642C is achieved after 1.375 sec. The reactor inlet composition after mixing the quench effluent compositions at 0.067, 0.202, 1.208 and 1.375 seconds, respectively, are set forth below in Table 5.
27,418 Table 5 Component Wt % AftPr M _inq Wt % After Ouench/Sec 0.067 0.2Q2 1.028 1.375 H2 5.225.599 5.673 5.769 5.784 CO 39.6339.631 39.631 39.631 39.631 C2 5.675.673 5.673 5.673 5.673 H2O 12.4812.480 12.480 12.480 12.480 CH4 0.672.133 2.512 3.094 3.201 10 C2H4 2.253 2.716 3.336 3.440 C2H6 --0.355 0.566 0.981 1.069 C3H6 8.342 10.191 12.767 13.209 C3H8 36.3323.476 20.485 16.175 15.414 n-c4Hlo --0.022 0.032 0.047 0.050 15 iso-c4Hlo 0.035 0.041 0.Q48 0.049 Example 4 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 250 kg/hr of 100% propane at a temperature of 25C and 553 kg/hr of water at 25C.
Temperatures of approximately 730C are achieved after a residence time of 1.25 sec. The reactor inlet composition after mixing and the quench effluent composition at 0.062, 0.185, 0.939 and 1.253 seconds are set forth below in Table 6.
Table 6 .Com~onent Wt % After Mixinq Wt % After Ouench~Sec 0.062 0.186 0.939 1.253 H2 5.855.993 6.024 6.050 6.052 30 CO 44.4144.410 44.410 44.410 44.410 C2 6.366.357 6.357 6.357 6.357 H2O 33.7133.714 33.714 33.714 33.714 CH4 0.75 1.384 1.630 1.989 2.046 C2H4 - 1.072 1.464 2.064 2.168 35 C2H6 0.081 0.125 0.159 0.156 C3H6 3.041 3.740 4.319 4.360 C3H8 8.92 3.923 2.495 0.873 0.668 n-C4H10 -- 0.013 0.025 0.045 0.048 iso-C4H10 0.012 0.016 0.020 0.021 Example 5 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 250 kg/hr of 100% ethane at a temperature of 25C and 364 kg/hr of water at 25C. The 27,418 reactor inlet composition after mixing and the quench effluent composition after residence times of 0.060, 0.181, 0.917 and 1.225 (final), respectively, are set forth below in Table 7.
Table 7 5 Component Wt % After Mixinq Wt % After Quench/Sec 0.060 0.181 0.gl7 1.225 H2 6.276.546 6.616 6.706 6.720 CO 47.6247.621 47.621 47.621 47.621 C2 6.8260817 ~.817 6.817 6.817 10 H2O 28.9228.921 28.921 28.321 28.921 CH4 0.80 0.914 0.963 1.052 1.070 C2H4 3.790 4.760 6.014 6.211 C2H6 9.56 5.194 4.019 2.431 2.171 C3H6 -- 0.002 0.003 0.005 0.005 n-C4H10 -- 0.190 0.276 0.430 0.462 i s o- C4 H 10 Example 6 20 Following the process of Example 1, 2000 kg/hr of effluent A at 1300C is quenched with 250 kg/hr of 100% propane at a temperature of 25C and 553 kg/hr of water at 25C and a pressure of 30 arm. Temperatures of approximately 730C are achieved after a residence time of 0.68 sec. The reactor inlet composition after mixing and the quench effluent composition at residence times of 0.033, 0.101, 0.510 and 0.681 seconds (final), respectively, are set forth below in Table 8.
Table 8 Component Wt % After Mixinq Wt % After Ouench/Sec 30 0.033 0.101 0.510 0.681 H2 5.855.968 6.007 6.045 6.049 CO 44.4144.410 44.410 44.410 44.410 C2 6.366.357 6.357 6.357 6.357 H2O 33.7133.714 33.714 33.714 33.714 35 CH4 0.75 1.223 1.461 1.832 1.896 C2H4 -- 0.841 1.247 1.881 1.995 C2H6 -- 0.043 0.073 0.111 0.114 C3H6 2.454 3.287 4.104 4.188 C3H8 8.92 4.977 3.422 1.505 1.233 40 n-C4H10 0.005 0.011 0.024 0.027 isO-c4Hlo 0.008 0.012 0.016 0.017 27,418 Example 7 Following the process of Example 1, 2000 kg/hr of effluent B at 1300C is quenched with 578 kg/hr of 100% propane at a temperature of 25C and 0 kg/hr of water. The reactor inlet composition after mixing and the quench effluent composition after residence times of 0.091, 0.368, 1.393 and 1.860 sec., respectively, are set forth below in Table 9. The quench effluent temperature after each time interval is 747, 726, 711 and 708C, respectively.
Table 9 Component Wt % After Mixing Wt % After ~uench/Sec 0.091 0.368 1.393 1.860 H2 2.923.327 3.367 3.391 3.394 Co 46.6146.610 46.610 46.610 46.~10 15 CO2 13.5513.5S3 13.553 13.553 13.553 H20 14.4714.469 14.469 14.469 14.469 CH4 0.032.263 2.753 3.205 3.298 C2H4 3.472 4.116 4.764 4.913 C2H6 --0.543 0.731 0.839 0.842 20 C3H6 9.208 10.329 11.005 11.089 C3H8 22.40 6.398 3.862 1.902 1.558 n-C4H10 -- 0.103 0.150 0.199 0.209 isO-c4Hlo 0.054 0.060 0.064 0.065 Example 8 Following the process of Example 1, 2000 kg/hr of effluent B at 1300C is quenched with 250 kg/hr of 100% propane at a temperature of 25C and 225 kg/hr of water at 25C. The reactor inlet composition after mixing and the final quench effluent composition for residence times of 0.129, 0.345, 1.297 and 1.729 seconds, respectively, are set forth below in Table 10. Temperatures at each residence time are 799, 794, 792 and 791C, respectively.
27,418 Table 10 Component Wt % After Mixinq Wt % After Ouench~Sec _ 129 0.345 1.297 1.729 H2 3.043.266 3.270 3.271 3.271 5 CO 48.5548.549 48.549 48.549 48.549 C2 14.1214.117 14.117 14.117 14.117 H2O 24.1824.162 24.162 24.162 24.162 CH4 0.03 1.535 1.686 1.780 1.790 C2H4 2.735 3.066 3.321 3.362 10 C2H6 0.150 0.121 0.069 0.059 C3H6 ~~ 4.698 4.702 4.608 4.589 C3H8 10.10 0.700 0.230 0.025 0.014 n-C4H10 0.055 0.061 0.063 0.063 isO-c4Hlo ~~ 0.032 0.034 0.035 0.035 Example 9 Following the process of Example 1, 2000 kg/hr of effluent B at 1300C is quenched with 842 kg/hr of 100% propane at a temperature of 25C and 0 kg/hr of water. The reactor inlet composition after mixing and the quench effluent composition for residence times of 0.141, 0.331, 1.438 and 1.922 seconds, respectively, are set forth below in Table 11.
Temperatures at each of the above residence times are 685, 667, 641 and 636~C, respectively.
table 11 Component Wt % After Mixina Wt % After Quench/Sec 0.141 0.331 1.438 1.922 H2 2.652.997 3.041 3.108 3.120 CO 42.2842.280 42.280 42.280 42.280 30 C2 12.29~2.294 12.294 12.294 12.294 H2O 13.1213.125 13.125 13.125 13.125 CH4 0.031.436 1.681 2.120 2.209 C2H4 2.101 2.381 2.827 2.911 C2H6 --0.406 0.561 0.896 0.972 35 C3H6 ~~7.856 8.999 10.874 11.231 C3H8 29.6317.450 15.573 12.394 1~.774 n-C4H10 --0.023 0.030 0.041 0.044 iso-C4H10 0.032 0.035 0.040 0.041 27,418 Example 10 Following the process of Example 1, 2000 kg/hr of effluent B at 1300~C is quenched with 250 kg~hr of 100% propane at a temperature of 25C and 368 kg/hr of water at 25C. The reactor inlet composition after mixing and the final quench effluent composition for residence times of 0.131, 0.362, 1.328 and 1.774 seconds, respectively, are set forth below in Table 12. Temperatures at each residence time are 734, 722, 710 and 708C, respectively.
Table 12 Component Wt After Mixing Wt % After Ouench/Sec 0.131 0.362 1.328 1.774 H2 2.873.036 3.059 3.079 3.082 CO 45.9045.898 4~.898 45.898 45.898 15 CO2 13.3513.346 13.346 13.346 13~346 H2O 28.3028.304 28.304 28.304 28.304 C~4 0.030.800 1.006 1.288 1.346 C2H4 ~~1.242 1.547 1.980 2.078 C2H6 -- 0O135 0.189 0.240 0.242 20 C3H6 3.575 4.132 4.629 4.692 C3~8 9 553.629 2.467 1.162 0.993 n-C4H10 -- 0.021 0.034 0.054 0.058 iso-C4H10 0.015 0.017 0.020 0.021 Example 11 Following the process of Example 1, 2000 kg/hr of effluent B at 1300C is quenched with 250 kg/hr of 100% ethane at a temperature of 25C and 222 kg/hr of water at 25DC. The reactor inlet composition after mixing and the quench effluent composition for residence times of 0.127, 0.341, 1.295 and 1.732 seconds, respectively, are set forth in Table 13.
Temperatures at each residence time are 808, 790, 769, and 765DC, respectively.
27,418 Table 13 COmDOnent Wt After Mixina Wt % After Ouench/Sec 0.127 0.341 1.295 1.732 H2 3 . 043 . 326 3. 381 3 . 448 3 . 461 5 CO 48.6148.608 48.608 48.608 48.608 C2 14.1314.134 14.134 14 . 134 14.134 H2O 24.0724.070 24.070 24.070 24.070 CH4 0.03 0.170 0.214 0.284 0.301 C2H4 3.959 4.712 5.646 5.835 10 C2H6 10.11 5.487 4.687 3.362 3.113 C3H6 - 0.002 0.003 0.004 0.004 C3H8 0.003 0.003 0.002 0.002 n-C4H10 -- 0.239 0.318 0.442 0.472 15 iso-C4H10 -_ 0.000 0.000 0.000 0.000 In a representative partial oxidation process which does not employ an alkane quench, natural gas at an initial pressure of 300 psig is compressed, preheated and injected into a gasifier, operated at a pressure of about 800 psi, with steam, where it is partially oxidized at 1300-1400 C by combustion with 99.5~ oxygen from an air separation plant. The hot gases enter the water heat recovery section where they are cooled by heat exchange to produce 1500 psi saturated steam, then further cooled in a water scrubbing section. The gases then flow to an amine scrubbing unit for removal of CO2. The C2 content of the untreated gas is about 3% and it may not be necessary to remove in some applications, such as methanol synthesis.
27,418 Typical design and performance parameters for the partial oxidation of natural gas are as follows:
Oxygen/natural gas, ~wt/wt) 1.29 Steam/natural gas, (wt/wt) 0.20 Exit gas pressure, psi 800 Exit gas temperature, F 2500 Dry gas produced, scf/lb natural gas 62.0 CO + H2 produced, scf/lb natural gas 59.4 CO + H2 produced, mol/mol oxygen 3.89 CO + H2 produced, mol/mol steam 14.1 The feed preheat system includes a zinc oxide sulfur guard bed to assure that any catalysts in downstream synthesis gas processing steps are protected from sulfur contaminants. A
portion of the 1500 psi saturated steam is used to preheat the oxygen feed and to inject Examples 12-15 Use of Alkane Ouench with Partial Oxidation of 1762 Pounds of Methane Following the general procedure described above, the partial oxidation of methane with no alkane quench (Example 12); 785 kg propane quench at 25C per 2000 kg raw syngas (Example 13); 1141 kg propane quench at 25C per 2000 kg raw syngas (Example 14); and 250 kg propane + 268 kg water quench at 25C per 2000 kg raw syngas (example 15) produced the results set forth in Table 14.
Table 14 Time Temp Temp (H2-C2L _2H4 Eauiv- Lb. C3H8 Onch. Lb (Sec) mix C ex C H2~ (CO+C021 Svnaas Alkane Net Gross Net Examule 12 0.065 1300 1300 1.8262 1.3900 1255 ---- 1255 Base No Qnch Example 13 35 0.065 900 752 1.9810 1.7321 1361 939 2300 1731 1180 0.197 900 734 1.~950 1.7449 1371 1054 2425 1731 1342 0.996 900 713 2.009 1.7577 1378 1191 2569 1731 1545 1.331 900 710 2.0105 1.7591 1378 1211 2589 1731 1575 27,418 Example 14 0.067 800 698 1.9583 1.7113 1346 734 2080 2515 890 0.202 800 676 l.g842 1.7351 1363 8~4 2257 2515 1097 1.028 800 647 2.0178 1.7658 13~2 1115 2497 2515 1395 1.375 800 642 2.0230 1.7706 1385 1153 2538 2515 1448 Example 15 0.060 9Q0 824 1.8891 1.6479 ~2~8399 1697 551 500 10 0.182 900 818 1.8911 1.6497 1299424 1723 551 536 0.911 900 816 1.8914 1.6500 1300435 1735 551 550 1.125 900 815 1.8914 ~.6500 1300436 1736 551 551 In a representative reforming process not employing 15 an alkane quench, natural gas feed and steam are heated and passed through catalyst-filled tubes in the reformer furnace. Natural gas is burned in the furnace to supply the endothermic heat of reaction and the sensible heat to the hot product gases. A heat recovery system which 20 recovers heat from the furnace flue gases and the reformed gases to preheat the feed and generate high pressure superheated steam is employed. A zinc oxide guard bed is included in the feed preheat system to ensure that the reformer catalyst is protected from sulfur contamination.
The gases are further cooled to condense the unreacted steam, processed through an amine unit to remove CO2, and compressed to 750 psig for delivery.
The high pressure, superheated steam is expanded through a turbine to supply power for the gas compressors 30 and intermediate pressure steam for the reformer furnace.
Low pressure steam is generated by heat recovery from product gas and flue gas to provide reboil steam for the C2 removal unit.
The synthesis gas has a H2/CO ratio of 5.4 and 3S contains 3.37% of methane and 0.1% of nitrogen.
27,41~
2 g Typical design and performance parameters are as follows:
Steam/carbon, moles/atom3.30 Steam/natural gas feed, wt/wt 3.68 Reformer exit pressure, psia 275.00 Reformer exit temperature, F 1560.00 Dry gas produced, scf/lb feed 91.9 CO + H2 produced, scf/lb feed 80.9 CO + H2 produced, mol/mol steam 1.04 Examples 16-20 Use of Alkane Ouench with Reforming 936 Pounds Methane Employing a natural gas which is low in CO2 and nitrogen content, and following the general procedure described above, reforming of 936 pounds of methane with no alkane quench (Example 16); 200 kg propane quench at 25C per 2000 kg raw syngas (Example 17); 500 kg propane quench at 25C per 2000 kg raw syngas (Example 18); 500 kg propane quench at 600C per 2000 kg raw syngas (Example 19); and 200 kg propane quench at 600C per 2000 kg raw syngas (Example 20) produced the results set forth in Table 15.
Table 15 Time Temp Temp (H2-C2L C2H4 Equiv- Lb- C3Hg Onch. Lb (Sec) mix,C ex,~C H2/CO ~CO+CO21 Synqas Alkane Net Gross Net Example 16 -- 857 857 5.3751 2.6931 918 ---- 918 Base No Qch Example 17 0.654 779 720 5.5232 3.0004 918 284 1202 441 348 Example 18 0.682 690 741 5.5340 3.0070 918 270 1188 1102 317 27,418 Example 19 0.633 793 694 5.6920 3.1039 982 594 1576 1102 730 5 Example 20 0.628 827 759 5.5377 3.0092 960 329 1289 441 409 In a typical autothermal reforming process, using the typical design and performance parameters for the autothermal reforming of natural gas set forth in Table 16 below, the natural gas feed and steam, preheated to about 1000F, and 99.5% oxygen preheated to about 500F, are fed into catalyst filled reactors where reforming is carried out. The feed preheat system includes a zinc oxide guard bed to insure that downstream catalysts are protected from sulfur contamination.
The exit gases at about 1750F enter the heat recovery section where they are first cooled by generating high pressure unsaturated steam. The gases are further cooled to condense unreacted steam, fed to an amine unit to remove CO2,. and compressed to a pressure of 750 psig. The saturated steam, superheated in a natural gas fired unit, is expanded through turbines to provide power before being injected in the gasifier or used for stripping steam in the CO2 removal unit. The amount of steam generated is in balance with the steam demands, although some electric power is purchased.
Table 16 Desiqn and Performance Parameters Steam/carbon, mols/atom 2.10 Steam/natural gas, wt/wt 2.34 Oxygen/natural gas, wt/wt 1.10 Reformer exit pressure, psia 290 Reformer exit temperature, F 1750 Dry gas production, scf/lb natural gas 73.9 CO + H2 production, scf/lb natural gas 65.1 27,418 3 !~
Co + H2 production, mol/mol oxygen 4.98 CO+ H2 production, mol/mol steam 1.32 Examples 21-22 Use of Alkane Ouench in Autothermal Reforminq of 984 Pounds of Methane Following the general procedure described above, autothermal reforming of 984 pounds of methane with no alkane quench (Example 22); and as modified by the process of the 10 present invention using 500 kg of propane quench at 600C per 2000 kg raw syngas to first cool the exit gases (Example 23) produced the results set forth in Table 17.
Table 17 Time Temp Temp (H2-C2L C2~4 Equiv- Lb. C3H~ Onch. Lb (Sec) mix C ex.C H2/CQ rCO+CO21 Syngas Alkane Net Gross Net Example 22 --- 940 940 3.3781 1.7988 793 Base 793 Base No qch Example 23 0.069 848 746 3.6138 1.9496 836 539 1375 1102 657 25 0.208 848 729 3.6476 1.9712 842 634 1476 1102 782 0.348 848 720 3.6610 1.9800 844 674 1518 1102 836 0.686 848 712 3.6751 2.0000 8~7 722 1569 1102 902 It will be understood by those skilled in the art that the optimum amount and type of quench for the particular situation, within the parameters of this invention, will be determined by local feedstock availability, markets, end products desired, and the like.
The above discussion is intended to be only illustrative of the invention. The full spirit and scope of the invention should be determined by reference to the following claims.
27,418
Claims (22)
1. An improved method for producing syngas and olefins comprising the steps of: reacting a hydrocarbon containing feedstock with an oxygen-containing material in a reaction zone of a reactor at a temperature of from 1000°-2000°C to produce an effluent stream comprising reaction products consisting essentially of hydrogen, COx wherein x is 1 or 2 and water; and introducing an alkane quench material having a temperature of up to 1000°C into said effluent stream to produce a chemically quenched effluent stream having a reduced temperature.
2. The process of claim 1 wherein said chemically quenched effluent stream has a temperature of from 600-800°C.
3. The process of claim 1 wherein said chemically quenched effluent stream has a temperature of from 600-750°C.
4. The method of Claim 1 wherein the temperature of said chemically quenched stream is further reduced by thermal means to retard retrograde reactions and produce a product stream.
5. The method of Claim 4 wherein said thermal means comprises a heat exchanger, water or recirculating refractory oils.
6. The method of Claim 4 additionally comprising the step of recycling a recycle stream comprising carbon dioxide formed during said chemical quenching and downstream product synthesis to said reaction zone.
7. The method of Claim 1 wherein said alkane quench material comprises a C3-C20 alkane, and wherein chemical quenching said effluent stream with said alkane quench material results in the conversion of at least a portion of the C3-C20 alkane quench material to alkenes with associated production of hydrogen.
8. The method of Claim 1 wherein said reactor is a partial oxidation reactor, and the temperature is maintained at from about 1000 to 1450°C.
9. The method of Claim 1 wherein said reactor is a classical reforming reactor, and the temperature is maintained at from about 800 to 900°C.
10. The method of Claim 1 wherein said reactor is an autothermal reforming reactor, and the temperature is maintained at from about 800 to 1250°C.
11. The method of Claim 1 wherein said reactor is a pyrolysis reactor, and the temperature is maintained at from about 1300 to 2000°C.
12. The method of Claim 1 wherein said alkane quench material is an alkane having from 3 to 5 carbon atoms.
13. A method for converting a hydrocarbon feedstock to syngas and olefins comprising: reacting the feedstock with an oxygen- containing material in a reaction zone of a partial oxidation reactor at a temperature of about 1000° to 1450°C to produce an effluent stream comprising reaction products consisting essentially of COx wherein x is 1 or 2, H2 and water; contacting said effluent stream with an alkane quench material having a temperature of up to 1000°C to produce a 27,418 chemically quenched effluent stream having a reduced temperature; further reducing the temperature of said chemically quenched stream by thermal means to retard retrograde reactions and produce a product stream; and removing water and hydrogen from said product stream.
14. The process of Claim 13 additionally comprising the steps of converting said product stream to a lower alkyl alcohol; separating and removing unreacted methane, hydrogen, water and inert materials; and recycling carbon dioxide to said reaction zone.
15. The process of Claim 13 wherein said alkane quench material is an alkane having from 3 to 5 carbons atoms.
16. The process of Claim 13 wherein said alkane quench material is propane.
17. A method for converting a hydrocarbon feedstock to syngas and olefins comprising: reacting the feedstock with an oxygen-containing material in a reforming reactor at a temperature of about 800 to 950°C from to produce an effluent stream comprising reaction products consisting essentially of COx wherein x is 1 or 2, H2 and water; contacting said effluent stream with an alkane quench material having a temperature of up to 950°C to produce a chemically quenched effluent stream having a reduced temperature; further reducing the temperature of said chemically quenched stream by thermal means to retard retrograde reactions and produce a product stream; and removing water and hydrogen from said product stream.
18. The process of Claim 17 wherein said alkane quench material is an alkane having from 3 to 5 carbons atoms.
19. The process of Claim 17 wherein said alkane quench material is propane.
20. A method for converting a hydrocarbon feedstock to syngas and olefins comprising: reacting the feedstock with an oxygen-containing material in an autothermal reforming reactor at a temperature of about 800 to 1250°C from to produce an effluent stream comprising reaction products consisting essentially of COx wherein x is 1 or 2, H2 and water;
contacting said effluent stream with an alkane quench material having a temperature of up to 1000°C to produce a chemically quenched effluent stream having a reduced temperature; further reducing the temperature of said chemically quenched stream by thermal means to retard retrograde reactions and produce a product stream; and removing water and hydrogen from said product stream.
contacting said effluent stream with an alkane quench material having a temperature of up to 1000°C to produce a chemically quenched effluent stream having a reduced temperature; further reducing the temperature of said chemically quenched stream by thermal means to retard retrograde reactions and produce a product stream; and removing water and hydrogen from said product stream.
21. The process of Claim 20 wherein said alkane quench material is an alkane having from 3 to 5 carbons atoms.
22. The process of Claim 21 wherein said alkane quench material is propane.
Applications Claiming Priority (2)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| US37630389A | 1989-07-06 | 1989-07-06 | |
| US376,303 | 1989-07-06 |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA2019001A1 true CA2019001A1 (en) | 1991-01-06 |
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Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA 2019001 Abandoned CA2019001A1 (en) | 1989-07-06 | 1990-06-14 | Process for manufacturing syngas |
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| CA (1) | CA2019001A1 (en) |
Cited By (1)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US11230679B2 (en) | 2017-09-14 | 2022-01-25 | Torrgas Technology B.V. | Process to prepare a char product and a syngas mixture |
-
1990
- 1990-06-14 CA CA 2019001 patent/CA2019001A1/en not_active Abandoned
Cited By (2)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US11230679B2 (en) | 2017-09-14 | 2022-01-25 | Torrgas Technology B.V. | Process to prepare a char product and a syngas mixture |
| US11667856B2 (en) | 2017-09-14 | 2023-06-06 | Torrgas Technology B.V. | Process to prepare a char product and a syngas mixture |
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