BACKGROUND OF THE INVENTION

Field of the Invention The present invention relates to the production methanol from hydrocarbonaceous feedstock by a process whi includes the partial oxidation of hydrocarbonaceous feedstock to produce a synthesis gas containing hydroyen, carbon monoxide and carbon dioxide, which is further processed and fed to a methanol synthesis loop. Description of the Prior Art

Methanol has long been produced by reacting hydrogen with carbon monoxide and/or carbon dioxide in the presence of a catalyst according to the equation: CO + 2H ₂ < > CH ₃ OH (1)

Hydrocarbonaceous feedstock, such as natural gases recover

from sites near petroleum deposits, are convenient startin materials for the production of methanol. Typically, natural gases contain, as their principal constituent, methane, with minor amounts of ethane, propane and butane. Also included in the conversion in some instances, may be low-boiling liquid hydrocarbons.

In order to convert a hydrocarbonaceous feedstoc into a feedstream suitable for introduction into a methano synthesis reactor, the feedstock is first converted into a synthesis gas containing hydrogen, carbon monoxide and carbon dioxide. The synthesis gas can be treated to adjus the ratio of hydrogen to carbon monoxide and carbon dioxid in order to obtain the proper stoichiometric proportions f methanol synthesis. The treated gas stream is compressed and fed to the methanol synthesis loop, where the carbon monoxide, carbon dioxide and hydrogen are reacted in conta with a catalyst -to- produce methanol. The methanol is then purified by any of a number of conventional techniques .

Methods that have been employed to convert hydrocarbonaceous feedstock to synthesis gas in prior art methanol production processes include steam reforming, combined steam reforming/autother al reforming, and partial oxidation. Steam reforming involves an endother ic reactio exemplified for methane by the equations

CH _* + H ₂ 0 > CO + 3H ₂ (2)

Partial oxidation involves an exothermic reaction exemplified for methane by the equation:

The product of both reactions (2) and (3) can be modified the exothermic water gas shift reaction according to the equation:

CO + H ₂ 0 > C0 ₂ + H ₂ (4) In steam reforming, desulfurized hydrocarbonaceo feedstock is mixed with between two and four moles of stea per mole of carbon and introduced into catalyst-filled tub in a primary reforming furnace, where it is converted to synthesis gas containing mainly hydrogen, carbon monoxide, carbon dioxide and residual methane and steam. The composition of the synthesis gas at the exit of the reforming furnace is dependent on the steam-to-feedstock ratio at the inlet and the temperature and. ressure at the outlet of the reforming furnace. Employing a high steam-to-gas ratio and a high temperature to increase production, however, also increases reformer fuel requirements, and low pressure operation increases the compression power requirement of the synthesis gas compressor which delivers the relatively low pressure synthesis gas into the higher pressure methanol synthesis loop. Typically, steam-to-carbon (in the feedstock) molar ratios in the range from 2.8:1 to 3.5:1 have been used wit a reformer tube exit temperature in the range from 850°C t 888°C and an exit pressure in the range from 15 to 25 bar.

Under these operating conditions, the residual methane content in the synthesis gas is approximately 3-4 mole% ( basis). Waste heat in the hot synthesis gas is recovered raising steam and preheating boiler feedwater. The endother ic heat of reaction is supplied by firing burners adjacent the catalyst-filled tubes in the refractory lined reforming furnace. Waste heat in the fl gas is recovered by raising and superheating steam and preheating combustion air. After heat recovery and final cooling, the synthesis gas is compressed in a centrifugal compressor to between 50 and 100 bar.

The fresh synthesis gas joins the gases circulating in the methanol synthesis loop at the suction the circulation compressor. From the discharge of the circulation compressor, the bulk of the circulating gases are preheated to the reaction temperature of 210° to 270° and fed to a catalytic methanol synthesis reactor. Formation of methanol at the operating conditions of the synthesis reactor is low, typically only 3.0-7.0%, depend upon the selected pressure of operation. This leads to t requirement for a circulating loop system where the reactants pass over the catalyst a number of times.

Gases leaving the reactor are used to preheat t circulating gas being fed to the reactor before crude methanol is condensed in a cooler and separated in a knock-out drum. The remaining gases return to the circulating compressor after removal of a purge gas strea by which the level of inert methane and excess hydrogen i the synthesis loop is controlled. The purge stream is utilized as fuel in the reforming furnace. Crude methano is discharged via a let down valve to a low pressure separator where dissolved gases flash off and are passed the reformer fuel system. The crude methanol is purified a distillation system.

Conventional steam reforming processes produce a synthesis gas which is hydrogen rich for methanol synthesis. This excess hydrogen has to be compressed to methanol loop pressure and then purged from the loop to be used as fuel in the reforming furnace.

Steam reforming can be supplemented by autother reforming to produce a stoichiometric synthesis gas according to the following reactions:

CO + 3H _a (2) COa + H ₂ (4)

HaO (6

Desulfurized natural gas feedstock is split into two streams. The first stream is mixed with steam and introduced into the primary steam reformer. As excess methane is oxidized in the downstream secondary reformer, the primary reformer can be operated at higher pressure, lower temperature and lower steam to gas ratios than in t case of the single primary reforming furnace route. The synthesis gas from the primary steam reformer is mixed wi the second stream of natural gas feedstock and introduced into the secondary reformer with preheated oxygen. Heat reaction in the secondary reformer is supplied by combust of methane, hydrogen and carbon monoxide. The exit temperature of the secondary reformer is typically 950-1000°C. Waste heat in the hot synthesis gas is recovered by raising steam, preheating boiler feedwater a providing distillation reboiler heat. Waste heat in the flue gas from the primary steam reformer is recovered by superheating steam, preheating feedstock and combustion air. The reduced flow rate of make-up synthesis gas is a higher pressure than in the conventional route and theref requires significantly less compression power. The metha

synthesis process is similar to that described above in conjunction with the steaπ. reforming route.

Since reforming is carried out in the primary a secondary reformers, the size of the primary roforming furnace can be reduced by as much as 75%. The total natu gas usage (feedstock plus fuel) is reduced by ppro iinate 6% due to reduced fuel requirement in the primary reformi furnace and reduced feedstock requirements as a result of more efficient utilizatior. of the carbon contained in the feedstock. The capital ccst is also reduced when compare to the conventional route as a result of the reduction in primary reforming furnace size, gas volumetric flow rates and compression power.

The autothermal reforming step requires a relat large volume of catalyst. Typically, space velocity requirements are between E,000 and 12,000 hr. ^-1 As used herein "space velocity" means the volumetric hourly throughput per volume of catalyst and the figures quoted herein refer to standard conditions of pressure and temperature.

Partial oxidaticn of hydrocarbonaceous feedstoc represents one alternative to steam reforming in the production of synthesis gas. The partial oxidation processes that have been used in connection with methanol production have been non-catalytic processes. Non-cataly partial oxidation reactions, however, are relatively inefficient. They operate at high temperatures, i.e., in the range of 2,200°F to 2,800°F (1205°C to 1340°C) and require large amounts of cxygen. Furthermore, free carbo is produced which is removed in a later step.

In methanol production processes of the prior a that employ partial oxidation, the feedstock is compresse to approximately 30-80 bar, heated and introduced to a partial oxidation generator. Preheated oxygen is injecte

- 6 -

into the generator burner. It has been reported that the feedstock is converted to carbon rich synthesis gas according to the following reactions:

CH _* + 20a > C0 ₂ + 2Ha0 (7) CIU + HaO > CO + 3H ₂ (2)

CIU + COa > 2C0 + 2H ₂ (8)

The endothermic heat of reaction for the reforming reactio (2) and (8) is supplied by the combustion of some methane, reaction (7). This combustion reaction is highly exothermic. Heat contained in the synthesis gas leaving t generator at approximately 1400°C is recovered by raising steam before the gas is passed to a carbon scrubber where free carbon is removed.

In order to adjust the carbon/hydrogen ratio in the synthesis gas to that required for methanol synthesis, it is necessary to reduce carbon oxides such as by shiftin carbon monoxide to carbon dioxide according to the reactio

CO + HaO > COa + Ha (4)

and then removing carbon dioxide by any of the convention or proprietary acid gas removal processes, or by removing the carbon monoxide directly by pressure swing adsorption. Following additional heat recovery and adjustment of the carbon/hydrogen ratio, the synthesis gas is compressed to synthesis loop pressure. Since the gas is stoichiometric and at considerably higher pressure, the power required f synthesis gas compression is reduced. The saving in synthesis gas compression, however, does not necessarily result in overall cost savings. As a result of the highe generator operating pressure, oxygen compression costs ar increased, and it is often necessary to include a nattiral gas compressor. This results in a net increase in power requirements, since substantial power, although reduced, still required to run the methanol process gas compressor

The additional power requirement and the less efficient utilization of carbon contained in the feedstock result an increase in specific natural gas requirements (i.e. natural gas consumed per unit of methanol produced) when compared to the conventional steam reforming process.

Conversion efficiency of oxidation processes ca generally be improved by the use of catalysts; but where oxidation process in only partial, i.e. with insufficien oxygen to completely oxidize the hydrocarbon, then the catalyst is subject to carbon deposit and blockage. Car deposits can be avoided by using expensive catalyst materials in generally uneconomical processes. For exam U.S. Patent 4,087,259, issued to Fujitani et al., descri employment of a rhodium catalyst in a process wherein li hydrocarbonaceous feedstock is vaporized and then partial oxidized in contact with the rhodium catalyst at a temperature in the range of 690° to 900°C with optional steam added as a coolant at rate not more than 0.5 by vol relative to the volume of the liquid hydrocarbon in terms the equivalent amount of water. The rhodium catalyst enables partial oxidation without causing deposition of carbon, but at temperatures greater than 900°C, thermal decomposition ensues producing ethylene or acetylene impurities. When steam is added, the quantity of hydrog produced is increased while the yield of carbon monoxide remains constant due to catalytic decomposition of the s to hydrogen gas and oxygen. A "LHSV" (Liquid Hourly Spa Velocity) from 0.5 to 25 1/hour is disclosed; particular a high yield from partial oxidation of gasoline vapor, without steam, is produced at a temperature of 725°C and a LHSV of 20, and with steam, is produced at temperature 700°C and 800°C and at a LHSV of 2.

The use of catalysts in partial oxidation processes requires that the reaction be carried out with

specific range of space velocity. In order to obtain acceptable levels of conversion using catalytic partial oxidation processes of the prior art it has been necessar to use space velocities below about 12,000 hr. ^-1 For example, U.S. Patent No. 4,522,894, issued to Hwang et al describes the production of a hydrogen-rich gas to be use as fuel for a fuel cell. The process reacts a hydrocarbo feed with steam and an oxidant in an autothermal reformer using two catalyst zones. The total hourly space velocit is between 1,960 hr. ^-1 and 18,000 hr. ^-1 . Because the pri art processes must be carried out at low space velocity, catalytic partial oxidation reactors of the prior art hav had to have large catalyst beds in order to achieve the throughput desired in commercial operation. This increas the size and cost of the partial oxidation reactor.

It is an object of the present invention to provide a process for the production of methanol from hydrocarbonaceous feedstock which is energy efficient, is capable of using low cost catalysts and employs relativel small, inexpensive equipment volume to achieve commercial acceptable throughput.

It is a further object of the invention to prov a process for the production of methanol from hydrocarbonaceous feedstock with a relatively low oxygen demand, thereby increasing throughput of hydrocarbonaceou feed.

These and other objects of the invention are achieved by a process which is described below.

Summary of the Invention The invention provides a process for producing methanol from hydrocarbonaceous feedstock in a manner wh uses relatively smaller and less costly equipment and operates at a relatively higher level of efficiency, in

terms of feedstock conversion, than prior art processes fo methanol production.

This invention provides a process for the production of methanol in which synthesis gas is generated by the catalytic partial oxidation of a hydrocarbonaceous feedstock, such as natural gas, with an oxidant stream und temperature and steam conditions producing essential no fr carbon at a space velocity in the range from 20,000 hour ^-1 - to 500,000 hour- ; hydrogen, carbon monoxide and carbon dioxide in the synthesis gas are reacted under methanol-producing conditions; and methanol is recovered. If necessary, the ratio of hydrogen to carbon monoxide and carbon dioxide in the synthesis gas is adjusted by removal of carbon monoxide and/or carbon dioxide to provide the proper stoichiometric amounts of reactants for the methano production reaction.

In one embodiment, the invention provides a process for producing methanol from hydrocarbonaceous feedstock which comprises: (a) introducing to a catalytic partial oxidatio zone a gaseous mixture of a hydrocarbonaceous feedstock, oxygen or an oxygen-containing gas and, optionally, steam which the steam-to-carbon molar ratio is from 0:1 to 3.0:1 and the oxygen-to-carbon molar ratio is from 0.4:1 to 0.8: said mixture being introduced to the catalytic partial oxidation zone at a temperature not lower than 200°F (93°C below its autoignition temperature and preferably at or above its autoignition temperature;

(b) partially oxidizing the hydrocarbonaceous feedstock in the catalytic partial oxidation zone to produ a gas consisting essentially of methane, carbon oxides, hydrogen and steam by passing the mixture througli a cataly capable of catalyzing the oxidation of the hydrocarbons, said catalyst having a ratio of surface area to volume rat

of at least 5 cm ² /cm ³ and a volume sufficient to produce a space velocity in the range from 20,000 hour-" ¹ to 500,000 hour ^-3 -;

(c) reacting the hydrogen, carbon dioxide and carbon monoxide under methanol-producing conditions, therel producing a product stream containing methanol; a d

(d) recovering the methanol. BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is an elevated cross-section view of a catalytic partial oxidation reactor having at its input a mixer and distributor suitable for introducing the reactan to the catalyst bed.

FIG. 2 is an enlarged elevational cross-section view of broken-away portion of the mixer and distributor o FIG. 1.

FIG. 3 is a top view of a broken-away quarter section of the mixer and distributor of FIG. 1.

FIG. 4 is a bottom view of a broken-away quarter section of the mixer and distributor of FIG. 1. FIG. 5 is a diagrammatic elevational cross-sectional illustration of a broken-away portion of mixer and feeder of FIGS. 1 and 2 showing critical dimensions.

FIG. 6 is a graph plotting σxygen-to-carbon mol ratio vs. steam-to-carbon molar ratio in the catalytic partial oxidation reaction for three different operating temperatures at an operating pressure of 400 psig (2760 KPa) .

FIG. 7 is a graph plotting the hydrogen-to-carb monoxide molar ratio in the catalytic partial oxidation reaction product vs. the steam-to-carbon molar ratio Cor three different operating temperatures at an operating pressure of 400 psig (2760 KPa).

FIG. 8 is the graph plotting the volume % methan in the catalytic partial oxidation product vs. the steam-to-carbon molar ratio for three different operating temperatures at an operating pressure of 400 psig (2760 KPa) .

FIG. 9 is a graph plotting the volume % carbon dioxide in the catalytic partial oxidation product vs. steam-to-carbon molar ratio for three different operating temperatures at an operating pressure of 400 psig (2760 KPa).

FIG. 10 is a graph plotting the molar ratio of total hydrogen and carbon monoxide in the catalytic partia oxidation product to total hydrogen and carbon vs. steam- to-carbon molar ratio in the feedstock for three different: operating temperatures at an operating pressure of 400 psi (2760 KPa) .

FIG. 11 is a flow diagram of a process for producing methanol from a hydrocarbonaceous feedstock in accordance with the invention FIG. 12 is a sectional view of a tube cooled converter for use in producing methanol in accordance with the invention.

FIG. 13 is a process flow diagram of a first portion of a large methanol plant in accordance with the invention.

FIG. 14 is a process flow diagram of a second portion of the large methanol plant of FIG. 13.

FIG. 15 is a process flow diagram of a third portion of the large methanol plant of FIG. 13. FIG. 16 is a process flow diagram of a first portion of a small methanol plant in accordance with the invention.

FIG. 17 is a process flow diagram of a second portion of the small methanol plant of FIG. 16.

- 12 -

DESCRIPTION OF THE PREFERRED EMBODIMENTS Essentially, the process of the invention involv three steps: conversion of the hydrocarbonaceous feedstoc into synthesis gas containing hydrogen, carbon dioxide and carbon monoxide by catalytic partial oxidation under temperature and steam conditions avoiding production of fi carbon; reaction of the hydrogen, carbon dioxide and carb monoxide under methanol producing conditions; and recover of methanol. Additionally, it may be necessary to adjust the molar ratio of hydrogen to carbon dioxide and carbon monoxide in the gas undergoing conversion to methanol, in order to provide stoichiometric amounts of these reactant in the methanol synthesis step. This may be done at any three points in the process. Catalytic Partial Oxidation

The catalytic partial oxidation of hydrocarbonaceous feedstock is carried out according to a process described in copending, commonly assigned U.S. application Serial No. 085,160 filed August 14, 1987 in t names of M. Dunster and J.D. Korchnak.

One particular aspect of the invention is the substantial capital cost savings and/or advantageous operating economy resulting from the employment of cataly partial oxidation to produce the raw synthesis gas employ in the methanol producing process. This is made possible the discovery that catalytic partial oxidation performed a temperature, as measured at the exit of the catalytic reaction zone, equal to or greater than a minimum non- carbon-forming temperature equal to or-greater than a minimum temperature selected as a linear function which includes a range from 1600°F (870°C) to 1900°F (1040°C) corresponding to a range of the steam-to-carbon molar ra from 0.4:1 to 0:1 and at a space velocity in the range f

20,000 hour ^-1 to 500,000 hour- ¹ produces essentially no f carbon deposits on the catalyst. Further, it is found th products of the partial catalytic oxidation in the proces of the invention consist essentially of hydrogen, carbon monoxide and carbon dioxide at oxidation temperatures equ to or greater than the minimum temperature, rhodium catalysts are not required to prevent carbon formation. example in FIG. 6, dotted line 25 represents a generally linear function which, at a steam/carbon ratio of 0, corresponds to a minimum partial oxidation temperature of about 1900°F (1040°C), and at a steam/carbon ratio of 0.4 corresponds to a minimum partial oxidation temperature of about 1600°F (870°C); favorable catalytic partial oxidati without producing free carbon occurs at temperatures and steam/carbon ratios equal to or greater than points on th line. Further, lower minimum temperatures at correspondi steam/carbon ratios greater than 0.4 can be extrapolated from the linear function represented by line 25.

Generally the catalytic partial oxidation is performed at a temperature, as measured at the exit of th catalyst, in the range from 1400°F (760°C) to 2300°F (1260°C). Preferably, the catalytic partial oxidation temperature, as measured at the exit, is in the range fro 1600°F (870°C) to 2000°F (1090°C). At temperatures below about 1400°F (760°C), uneconomic quantities of methane ar left unconverted, and at temperatures above 2300°F (1260° excessive amounts of oxygen are used.

The pressure at which the partial oxidation tak place is generally above 150 psig (1030 KPa) and preferab above 300 psig (2060 KPa). The pressure can be up to or above the methanol synthesis loop pressure. Preferably, pressure does not exceed the methanol synthesis loop pressure by more than that necessary to provide synthesis

gas flow through the processing equipment into the methanol synthesis loop.

Essentially little or no reforming reactions are employed in the process of the invention; that is, the process of the invention relies essentially solely on partial oxidation and the water gas shift reaction (equati 4) to convert hydrocarbonaceous feedstock to synthesis gas. Catalytic partial oxidation of uniformly premixed feedstoc and oxygen does not require any reforming reactions to tak place. The catalyst is selected to promote the partial oxidation reaction, and not necessarily any reforming reaction. The steam reforming reaction (equation 3) generally requires a low space velocity, i.e. generally below about 12,000 hour ^-1 , and the employment of space velocities above 20,000 hour ^-1 in the present process precludes efficient steam reforming of the feedstock. It believed that increased hydrogen production, above that attributable solely to partial oxidation, is due more to t water gas shift reaction (equation 4) than to the steam reforming reaction (equation 2 ) .

The process of the invention can be employed to convert hydrocarbonaceous feedstock to synthesis gas with very low levels of hydrocarbon slippage (unreacted feedstock), i.e. as low as 2% or lower, if desired. Becau the rate of reaction in the partial oxidation reactor is mass transfer controlled, the process of the invention can be carried out efficiently using relatively small volumes and relatively inexpensive catalyst materials, provided t the surface area-to-volume requirements of the invention met. In accordance with the process of- the invention the reactant gases are introduced to the reaction zone, i.e. catalyst bed, at an inlet temperature not lower than 200° (93°C) less than the autoignition temperature of the feed mixture. The autoignition temperature of the feed depend

on the composition and conditions of the feed mixture and the catalyst employed. Preferably the reactant gases are introduced at a temperature at or above the autoignition temperature of the mixture. A further essential feature the catalytic partial oxidation reaction is that the reactants shall be completely mixed prior to the reaction taking place. Introducing the thoroughly mixed reactant gases at the proper temperature ensures that the partial oxidation reactions will be mass transfer controlled. Consequently, the reaction rate is relatively independent catalyst activity, but dependent on the surface-area-to-volume ratio of the catalyst. It is possible to use any of a wide variety of materials as a catalyst, provided that the catalyst has the desired surface-area-to-volume ratio. It is not necessary that t catalyst have specific catalytic activity for the steam reforming reaction. Even materials normally considered t be non-catalytic can promote the production of synthesis g under the reaction conditions specified herein when used a a catalyst in the proper configuration. The term

"catalyst", as used herein, is intended to encompass such materials.

The catalytic partial oxidation step can be understood with reference to the figures. The catalytic partial oxidation zone is typically the catalyst bed of a reactor such as that illustrated in FIG. 1. Λs shown in FIG. 1, a reactor for partially oxidizing a gaseous feedstock includes an input mixing and distributor section indicated generally at 30. The mixer and distributor 30 mixes the feedstock with an oxidant and. distributes the mixture to the entrance of a catalytic reactor section indicated generally at 32 wherein the feedstock is partia oxidized to produce a product which is then passed throug the exit section indicated generally at 34.

The reactor includes an outer shell 40 of structural metal such as carbon steel with a top 42 secured thereon by bolts (not shown) or the like. Λ layer 43 of insulation, such as 2300°F (1260°C) BPCF ceramic fiber insulation, is secured to the inside of the upper portion o the shell 40 including the top 42. In the lower portion ol the mixing section 30 in the reactor section 32 and outlet section 34, there are secured layers 46, 48 and 50 on the inside of the shell. The layer 46 is a castable or equivalent insulation such as 2000°F (1090°C) ceramic insulation. The layer 48 is also a castable or equivalent layer of insulation but containing 60% alumina for withstanding 3000°F (1650°C). The internal layer 50 is a refractory or equivalent layer such as 97% alumina with ceramic anchors or 97% alumina brick for withstanding the interior environment of the reactor section.

The catalytic reactor section 32 contains one or more catalyst discs 54. Λs shown, the reactor contains a sequence of discs 54 separated by high alumina rings 58 between each adjacent pair of discs. The stack is support by a grill with high alumina bars 56. Λ sample port 60 is formed in the lower end of the reaction section and has a tube, such as type 309 stainless steel tube 62 extending below the bottom refractory disc 54 for withdrawing sample of the product.

The outlet section 34 is suitably formed for bei connected to a downstream heat recovery boiler (not shown) and/or other processing equipment.

The catalyst comprises a high surface area material capable of catalyzing the partial oxidation of th hydrocarbonaceous feedstock. The catalyst is in a configuration that provides a surface area to volume ratio of at least 5 cmVcm ³ . Preferably, the catalyst has a geometric surface area to volume ratio of at least 20

cm cm ³ . While there is no strict upper limit of surface area to volume ratio, it normally does not exceed about 4 cm ² /cm ³ . A wide variety of materials can be used in the construction of the catalyst including materials not normally considered to have catalytic activity, provided that the catalyst configuration has the desired surface a to volume ratio.

The catalyst disc 54 can be, for example, a monolithic structure having a honeycomb type cross-sectio configuration. Suitable monolithic structures of this ty are produced commercially, in sizes smaller than those us in the process of the invention, as structural substrates for use in the catalytic conversion of automobile exhaust and as catalytic combustion chambers of gas turbines or f catalytic oxidation of waste streams. Typically, the monolithic structure is an extruded material containing a plurality of closely packed channels running through the length of the structure to form a honeycomb structure. T channels are typically square and may be packed in a dens as high as 1,200 per square inch of cross section. The monolithic structure can be constructed of any of a varie of materials, including cordierite (MgO/Λla0 ₃ /Si0 ₂ ) , Mn/M cordierite (Mn-MgO/Λl ₂ 0 ₃ /Si0 ₂ ) , ullite (Λl ₂ 0 ₃ /SiOa), mullite aluminum titanate (Ala0 ₃ /Si0 ₂ -(Al,Fe)a0 ₃ /TiOa), zirconia spinel (Zr0 ₂ /MgO/Λl ₂ 0 ₃ ), spinel (MgO/Λl ₂ 0 ₃ ), alumina (Λla0 ₃ ) and high nickel alloys. The monolithic catalyst may consist solely of any of these structural materials, even though these materials are not normally considered to have catalytic activity by themselves . Using honeycombed substrates, surface area to volume rati up to 40 cm ² /cm ³ or higher can be obtained. Alternativel the monolithic substrate can be coated with any of the metals or metal oxides known to have activity as oxidatio catalysts. These include, for example, palladium, platin

rhodium, iridium, osmium, ruthenium, nickel, chromium, cobalt, cerium, lanthanum and mixtures thereof. Other metals which can be used tc coat the catalyst disc 54 include noble metals and metals of groups IΛ, IIΛ, III, I VB, VIB, or VIIB of the periodic table of elements.

The catalyst discs 54 may also consist of structural packing materials, such as that used in packin absorption columns. These packing materials generally comprise thin sheets of corrugated metal tightly packed together to form elongate channels running therethrough.

The structural packing materials may consist of corrugate sheets of metals such as high temperature alloys, stainle steels, chromium, manganese, molybdenum and refractory materials. These materials can, if desired, be coated wi metals or metal oxides known to have catalytic activity f the oxidation reaction, such as palladium, platinum, rhodium, iridium, osmium, ruthenium, nickel, chromium, cobalt, cerium, lanthanum and mixtures thereof.

The catalyst discs 54 can also consist of dense wire mesh, such as high temperature alloys or platinum mesh. If desired, the wire mesh can also be coated with metal or metal oxide having catalytic activity for the oxidation reaction, including palladium, platinum, rhodiu iridium, osmium, ruthenium, nickel, chromium cobalt, cesi lanthanum and mixtures thereof.

The surface area to volume ratio of any of the aforementioned catalyst configurations can be increased coating the surfaces thereof with an aqueous slurry containing about 1% or less by weight of particulate meta or metal oxide such as alumina, or metals of groups 1Λ, III, IV VB, VIB and VIIB and firing the coated surface a high temperature to adhere the particulate metal to the surface, but not so high as to cause sintering of the surface. The particles employed should have a BET

(Brunnauer-Emmett-Teller) surface area greater than about m ² /gram, preferably greater than about 200 m ² /yram.

A gaseous mixture of hydrocarbonaceous feedstoc oxygen or an oxygen-containing gas such as air, and, optionally, steam is introduced into the catalytic partia] oxidation zone at a temperature not lower than 200°F (93° below its autoignition temperature. Preferably, the gase mixture enters the catalytic partial oxidation zone at a temperature equal to or greater than its autoignition temperature. It is possible to operate the reactor in a mass transfer controlled mode with the reactants entering the reaction zone at a temperature somewhat below the autoignition temperature since the heat of reaction will provide the necessary energy to raise the reactant temperature within the reaction zone. In such a case, however, it will generally be necessary to provide heat input at the entrance to the reaction zone, for example b sparking device, or by preheating the contents of the reactor, including the catalyst, to α temperature in exce of the autoignition temperature prior to the introducing the reactants in order to initiate the reaction. If the reactant temperature at the input to the reaction zone is lower than the autoignition temperature by more than abou 200°F (93°C), the reaction becomes unstable. When the reactant mixture enters the catalytic partial oxidation zone at a temperature exceeding its autoignition temperature, it is necessary to introduce th mixture to the catalyst bed immediately after mixing; that is, the mixture of hydrocarbonaceous feedstock and oxidan should preferably be introduced to the catalyst bed befor the autoignition delay time elapses. It is also essentia that the gaseous reactants be thoroughly mixed. Failure l mix the reactants thoroughly reduces the quality of the product and can lead to overheating. Λ suitabJe apparatu

for mixing and distributing the hydrocarbonaceous feedstoc and oxygen or oxygen-containing gas so as to pi ovide thorough mixing and to introduce? the heated reactants into the reaction zone in a sufficiently short time is illustrated in Figs. 1-5 and described in more detaj J in cσpending commonly assigned U.S. patent application Serial No. 085,159 filed August 14, 1987 in the names of J.D. Korchnak, M. Dunster and J.H. Marten.

Referring to Fig. 1, one of the feed gases, i.e. hydrocarbonaceous gas or oxygen-containing gas, is introduced into the input section 30 through a first inlet port 66 through the top 42 which communicates to an upper feed cone 68 which forms a first chamber. The cone 68 is fastened by supports 69 in the top 42. The other feed gas is introduced into the input section 30 through second inlets 70 extending through side ports of the shell 40 and communicating to a second chamber 72 which is interposed between the upper chamber 68 and the inlet of the catalyst reaction section 32. A ring 73 mounted on the central portion of an upper wall 75 of the chamber 72 sealingly engages the lower edge of the cone 68 so that the wall 82 forms a common wall between the upper chamber 68 and lower chamber 72. The chamber 72 has an upper outer annular portion 74, see also Figs. 2 and 3, which is supported on the top surface of the refractory layer 50. Λ lower portj of the chamber 72 has a tubular wall 76 which extends downward in the refractory sleeve 50. The bottom of the chamber 76 is formed by a cast member 78.

Optionally, steam can be introduced into either both of the hydrocarbonaceous feedstock and oxygen ot oxygen-containing gas. The gases are fed to the reactor i relative proportions such that the steam-to-carbon molar ratio is in the range from 0:1 to 3.0:1, preferably f.rom 0.8:1 to 2.0:1. The oxygen-to-carbon ratio is from 0.4:1

0.8:1, preferably from 0.45 to 0.65. Although air may be used as the oxidant in the process of the invention, it is preferred to use oxygen or an oxygen-rich gas, in order to minimize the inert ingredients such as nitrogen that must carried through the system. by "oxygen-rich" is meant a g. containing at least 70 mole.% oxygen, preferabJy at l αst ' mole.% oxygen.

The reactant mixture preferably enters the catalytic reactor section 32 at a temperature at or above its autoignition temperature. Depending on the parti.cu.lar proportions of reactant gases, the reactor operating pressure and the catalyst used, this will generally be between about 550°F (290°C) and 1,100°F (590°C). Preferably, hydrocarbonaceous feedstock and steam are admixed and heated to a temperature from 650°F (340°C) to

1,200°F (650°C) prior to passage through inlet port(s) 70 < 66. Oxygen or oxygen-containing gas, such as air, is heat to a temperature from 150°F (65°C) to 1200°F (650°C) and passes through the other inlet port(s) 66 or 70. Referring again to FIG. 1, the mixing and distributing means comprises a plurality of elongated tube 80 having upper ends mounted in the upper wall 75 of the chamber 72. The lumens of the tubes at the upper end communicate with the upper chamber 68. The bottom ends of the tubes 80 are secured to the member 78 with the lumens ■ the tubes communicating with the upper ends of passageways 84 formed vertically through the member 78. Orifices 86 a formed in the walls of the tvibes 80 for directing streams gas from the chamber 72 into the lumens of the tubes 80. The inlets 66 and 70, the cone 68, the supports 69 n vG formed from a conventional corrosion and heat resistant metal while the chamber 72, tubes 00 and member 78 n m formed from a conventional high temperature alloy or refractory type material.

The number of tubes 80, the internal diameter 90 (see FIG. 5) of the tubes 80, the size and numb i; of the orifices 86 in each tube are selected relative to the gas input velocities and pressures through inlets 66 and 70 so as to produce turbulent flow within the tubes 80 at a velocity exceeding the flashback velocity of the mixture. The minimum distance 92 of the orifices 86 from the bottom end of the tube 80 at the opening into the diverging passageways 84 is selected to be equal to or greater than that required for providing substantially complete mixing o the gas streams from chambers 68 and 72 under the condition of turbulence therein. The size of the internal diameter 9 of the tubes 80 as well as the length 94 of the tubes is designed to produce a sufficient pressure drop in the gas passing from the chamber 68 to the reaction chamber so as t provide for substantially uniform gas flow through the tub 80 from the chamber 68. Likewise the size of the orifices 86 is selected to provide sufficient pressure drop between the chamber 72 and the interior of the tubes 80 relative t the velocity and pressures of the gas entering through inlets 70 so as to provide substantially uniform volumes o gas flows through the orifices 86 into the tubes 80.

The diverging passageways in the member 78 are formed in a manner to provide for reduction of the velocit of the gas to produce uniform gas distribution over the inlet of the catalyst. The rate of increase of the cross-section of the passageway 84 as i 'proceeds downw rd i.e., the angle 98 that the wall of the passageway 04 mnk". with the straight wall of the tubes 80, must genera] ly r? equal to or less than about 15° and preferably equal to υr less than 7° in order to minimize or avoid creating v r ic within the passageways 84. This assures that the essentially completely mixed gases, at a temperature near or exceeding the autoignition temperature, will pass into

the catalyst bed in a time preferably less than autoigniti delay time. The configuration of the bottom end of the passageways, as shown in FIG. 4, is circular.

The gas exiting the catalytic partial oxidation reactor contains hydrogen, carbon dioxide, carbon monoxide and some methane. The synthesis gas leaving the catalyst zone is first cooled by heat exchange, either by heating t hydrocarbon and steam feed stock, by heating the oxi ant stream, by super heating steam, by raising steam in a boiler, by preheating boiler feed water or any combination thereof .

Adjustment of Hydrogen to Carbon Oxide Ratio

The gas stream that is converted to methanol in the methanol synthesis loop contains hydrogen and carbon oxides, i.e. carbon dioxide and carbon monoxide. The mola ratio of hydrogen to carbon oxides in the gas to be converted can be expressed as the methanol stoichiometric synthesis gas ratio (MSSGR). The MSSGR is defined as the following molar ratio:

MSSGR = Ha

2 CO + 3COa

The MSSGR should have a value of at least 0.8 and preferab from about 0.95 to 1.1 for methanol synthesis. Normally, the synthesis gas from the catalytic partial oxidation reactor will be slightly carbon-rich, that is, the MSSGK i somewhat lower than that ideally desired. In order to correct the stoichiometry, carbon is removed, either from the synthesis gas passing to the methanol synthesis loop, the methanol loop or the purge gas from the methanol synthesis loop.

Carbon dioxide can be removed from the synthesis gas stream by any know method. For example, the gas stre

,__„ _n^ 06297

- 24 -

can be passed through a countercurrent liquid stream of α carbon dioxide absorbing medium. Commercial processing units for carbon dioxide removal a re available, for exa p] under the trademarks Selexol, Λmine Guard, and Benfieid 5 Since the amount of carbon dioxide j n the synthesis g s is relatively small it may be necessary to remove a portion of the carbon monoxide in ordor to achiev the desired MSSGR. Any knowij method for carbon monoxide removal can be employed. One suitable method involves 0 converting at least a portion of the carbon monoxide to carbon dioxide by water gas shift reaction and then remov the carbon dioxide from the gas stream. The water gas sh reaction is known, and suitable equipment for carrying ou the reaction is commercially available. 5 An alternative method for removing carbon oxide is pressure swing absorption. This procedure can be employed not only to remove the desired amount of carbon monoxide and/or carbon dioxide, but also to remove components of the gas stream that are not required for 0 methanol production such as methane and nitrogen. Pressu swing adsorption involves the adsorption of components to removed at high pressure followed by their desorption at pressure. The process operates on a repeated cycle havin two basic steps, adsorption and regeneration. Not all th 5 hydrogen is recovered as some is lost in the waste gas during the regeneration stage, but by careful selection o the frequency and sequence of steps within the cycle the recovery of hydrogen is maximized.

Regeneration of the adsorbent is carried out in 0 three basic steps.

(a) The adsorber is depressurized to the lnw pressure. Some of the waste components are desorbed during this step.

- 25 -

(b) The adsorbent is purged at low pressure, the product hydrogen removing the remaining wa components .

(c) The adsorber is repressurized to adsorpti pressure ready for service.

Pressure swing adsorption is most effectively employed to remove carbon oxides and inert matnriaJs, l example, methane and nitrogen, from the methanol synthes loop. Methanol Synthesis

The hydrogen, carbon dioxide and carbon monσxi are reacted under methanol-producing conditions. Any kn procedure for reacting hydrogen, carbon dioxide and carb monoxide to produce methanol can be used. Preferably, t are reacted in a medium pressure process in a circulatin catalytic reactor at a pressure from about 50 αtm (5070 to 120 atin (12160 KPa), more preferably from about 70 at (7090 KPa) to 100 atm (10130 KPa). This type of synthes and equipment for carrying it out are known in the art. The synthesis gas is compressed to between 50

(5070 KPa) and 120 atm (12160 (KPa), preferably between atm (7090 KPa) and 100 atm (10130 KPa). The gas enters methanol synthesis vessel in which it is admixed with recycle gas. The gas then passes into a catalytic metha converter. The methanol converter typically comprises a pressure vessel containing a catalyst bed and facilities moderating the exothermic reaction of hydrogen with the carbon oxides to produce methanol, for example by in ject cold gas at intervals within the catalyst bed. Any commercially available catalyst which is capable of catalyzing the reaction of hydrogen, carbon monoxide and carbon dioxide to produce methanol can be employed. Such catalysts are manufactured, for example, Imperial Chemical Industries, Katalco, and Haldor Topsoe

Inc. Preferred catalysts for methanol synthesis are composed of zinc oxide and copper oxide. The methanol synthesis reaction generally takes place at a temperature the range from about 41()°F (210°C) to 570°F (300°C), depending on the activity of the particular catalyst employed.

The exit gas from the methanol converter is pas; through a condenser in which it is cooled with water and then through a separator. The bottoms product of the separator contains methanol condensate. An amount of gas purged from the methanol synthesis loop in order to maint the concentration of inert gases circulating within the l at acceptable levels. The remaining gas is admixed with incoming synthesis gas to be recycled to the methanol converter.

As previously indicated, carbon monoxide and/or carbon dioxide can be removed either from the circulating gas in the methanol synthesis loop or from the purge gas order to maintain the desired MSSGR, preferably between 0 and 1.10, entering the converter. When carbon oxides are removed from the purge gas, the remaining hydrogen-rich g is recycled to the methanol synthesis loop.

The crude methanol containing water is then purified by conventional means to obtain essentially pure methanol. Preferably, the methanol condensate is purifie in one or more distillation columns.

The process of the invention can be understood further by reference to the flow diagram of FIG. 11.

When compared to present day commercial processes, the? catalytic partial oxidation process offers the folJ wing advantages .

(1) The high cost steam reforming furnace i eliminated.

(2) Low catalyst volume when compared to c?i.th stea reforming or autothermal reforming.

(3) Low oxygen consump ion when compared to conventional non-cafαlytic partial oxidation processes.

(4) Low water consumption when compared to st reforming or steam reforming plus autothermal reforming. The process is therefore parti culα suitable either when water is not availabJe (e. in desert locations) or is expensive (e.y. mus produced by desalination).

(5) Reduced area requirement when compared to steam reforming route to methanol (particularl suitable for offshore application). (6) High efficiency in terms of feedstock conversion when compared to present processes (either reforming or partial oxidation). (7) Lower in capital cost then all present commercial processes. A process flow diagram of a methanol plant designed for maximum efficiency is shown in FIGS. 13-15. Hydrocarbonaceous feedstock, such as natural gas stream FIG. 14, is optionally desulfurized using conventional methods. Desulfurization may, for example, be convenient carried out by preheating the hydrocarbonaceous feedstoc a temperature between about 250°F (120°C) and 750°F (400 and absorbing the sulfur compounds into zinc oxide contai in one or several desulfurization vessels 202. The desulfurization vessel 202 is located upstream of a feedstock saturator 204 in the embodiment shown in FIG. alternatively, the desul urization vessel 202 can be loc downstream of the synthesis gas compressor, as shown in embσdiment in FIG. 16.

The desulfurized hydrocarbonaceous feedstock i saturated with water vapor in the forced film saturator 20 The forced film saturator 204 is a distinctive feature of the process and results in reduced capital cost and power requirements when compared with the more conventional pacJ. tower type of saturator presently employed in coniiiierciaJ installations. The forced film saturator 204 consists of vertical shell and tube exchanger and a water circulation system 206. The desulfurized hydrocarbonaceous feedstock and water enter the top head 208 of the exchanger and flo vertically downward through the tubes 210. The feedstock heated as it flows though the tubes, for example to a temperature of about 400°F (206°C), and as it is heated, water vapor content increases due to vaporization of the circulating water. The heating medium, in this case methanol reactor effluent gas stream 330, passes to the shell side of the tubes 210 in the forced film saturator the bottom and flows countercurrent to the feedstock flow emerge at a lower temperature at the top outlet of the shell. The desulfurized hydrocarbonaceous feedstock and unvaporized water emerge from the exchanger tubes 210 at bottom and are separated in the bottom head 212. Recover water is recirculated by the pump 206 to the top 208 of t forced film saturator and the saturated hydrocarbonaceous feedstock passes to saturated feedstock line 220. Makeup water 216 is added to the recycle water stream, and blowd line 218 for the saturator water recycle stream is ptovid The major advantages of the forced film sal.mat 204 over the conventional packed bed saturator are: (1 ) forced film saturator has a considerably lower water circulation rate and hence lower power consumption. (7.) The forced film saturator has a simpler design and consequently improved operability and reliability. (3)

capital cost of the forced film saturator is less than t of the packed bed type.

Λs shown in FIG. 1J, the saturated feedstock passes through line 220 to coil 222 of a fired heater 224 where the saturated feedstock is preheated to a temperati in the range from 650°F (340°C) to 1200°F (650°C), fυr example to 1100°F (595°C), by combustion of fuel in the fired heater with air. The fuel 225 can be one or more waste gas streams from the methanol plant, such as fusel stream 226 from distillation, flash gas stream 228 from ammonia synthesis loop, PSΛ purge gas stream 230, and li ends stream 232 from distillation which are combusted wi air 227.

The heated saturated feedstock on line 234 fro the heater 224 is fed to the catalytic partial oxidation

(CPO) reactor 28 where it is mixed with an oxygen or oxy containing stream 236 as has been described hereinbefore. The oxygen stream 236 is obtained from an air separation plant 230 and preheated to a temperature in the range fr 150°F (65°C) to 1200°F (650°C), for example, 300°F befor being mixed with the natural gas and steam feedstock pas to the CPO catalyst. High pressure steam is used to pre the oxygen in a heat exchanger 240. The. ain overall reactions taking place within the CPO reactor 28 are the partial oxidation reactions:

C„Ha„ _* a + "/a0 ₂ > nCO + (n + 1)H ₂ ( and the water gas shift reaction:

CO + HaO > C0 ₃ + H ₂ (

FIG. 6 shows oxygen consumption for the above process, as a function of the steam-to-carbon molar rati natural gas feedstock, for reaction temperatures of 1,60 (870°C), 1,750°F (950°C) and 1,900°F (1040°C) and an operating pressure of 400 psig (2760 KPa). It can be se from the graph that oxygen consumption, expressed as

oxygen-to-carbon molar ratio, is relatively low for the process of the invention as compared with present commerci partial oxidation processes. The dashed line 25 in FIG. represents the minimum temperature and steam conditions, i.e. the minimum temperature is α linear function of the steam-to-carbon ratio, which have been discovered to prev formation of carbon deposits on the catalyst. Generally, lower temperature of reaction is preferred, such as an ex temperature of 1700°F (925 ^C C) at a pressure of about 415 psig (2860 KPa). Conveniently, the saturator 204 provide the total quantity of steam necessary to achieve a steam- carbon molar ratio of approximately 1.3 to 1.0 and additional ma e-up steam is not required to enable the catalytic partial oxidation reaction at the desired temperature in accordance with the invention.

FIG. 7 shows the molar ratio of hydrogen, s H ₂ to carbon monoxide in the product as a function of steam-to-carbon ratio for reaction temperatures of 1,600° (870°C), 1,750°F (950°C) and 1,900°F (1040°C). FIGS. 8 and 9, respectively, show the amounts o methane and carbon dioxide, as volume %, in the product a function of steam-to-carbon ratio for reaction temperatur of 1,600°F (870°C), 1,750°F (950°C) and 1,900°F (1040°C). FIG. 10 shows the effective H ₂ production of th process, expressed as total moles of H ₂ and carbon monoxi in the product divided by total moles of Ha and carbon in the feedstock as a function of steam-to-carbon ratio for reaction temperatures of 1,600°F (870°C), 1,750°F (9 ^r >0"C) and 1,900°F (1040°C) . The reactor effluent 244 is first cooled by generating steam in a boiler 246 which receives water f i high pressure steam drum 248 and returns steam to the dr The drum 248 operates at a high pressure of, for example about 1550 psig (10700 KPa). Water supply 250 for the s

drum is first heated in coils 252 of the heater 224 and t fed to the drum through line 254. Steam output. 256 from drum 248 is further heated in coils 258 to produce superheated steam 260 which can be used to operate i.earn turbines or provide heating for process steps. Line 262 a blow down drum (not shown) provides for blow down of th drum 248.

The reactor effluent is passed by line 264 to distillation column reboiler 266, by line 268 to distillation column reboiler 270, by line 272 to demineralized water heater 274, and by line 276 to synthe gas cooler 278 for further cooling, for example, to 10ϋ°F (30°C), by recovering and utilizing the heat in the synthesis gas. Water in the process stream is condensed the stream passes to knockout drum 280 where the water 20 is separated from the synthesis gas. Alternative cooling and heat recovery schemes can be used.

After water removal, the synthesis gas in line is mixed with hydrogen 204 from PSA unit 286, FIG. 14, an passes to mαke-up compressor unit 288 where the synthesis gas is compressed to the methanol loop pressure, for example, to 1220 psig (0410 KPa). The make-up compressor unit of FIG. 13 includes compressor 290, heat exchanger 2 knockout drum 294, and compressor 296 to produce the make gas stream 290 which is passed to the methanol synthesis loop in FIG. 14. The compressors 290 and 296 can conveniently share a common steam turbine drive 300 with methanol loop circulator 302, FIG. 14.

The make-up synthesis gas 298 from the com ess unit 280 joins the gas circulating in the methanol synth loop at the discharge 304 of the circulator 302. The lo stream 306 is preheated in loop heat interchanyer 300 bef ^' passing in line 310 to tube cooled methanol converter 3.1 Start-up heater 314 is provided so that, upon start up o

the methanol converter, the incoming process stream can be heated to the reaction temperature until the methanol converter becomes hot enough due to the heat generated by the reaction to heat the incoming gases, whereupon oporati.* of the start-up heater 314 can be discontinued.

The tube cooled converter, shown in detaiJ in FI< 12, is a distinctive feature of the process and includes a simple gas-to-gas heat exchanger with no high differential pressure tubesheets. The vessel shell 314 is designed to retain the process pressure. Inlet 316 connects through distributor 310 to intermediate branch tubes 320 which, in turn, are connected through distributors 322 to a plurali of tubes 324 extending vertically upward through the catalyst bed 326. The catalyst bed is supported on a bed ceramic balls 320 which are separated from the outlet 330 screen support 332. The reactant gas is fed though inlet 316 into the bottom of the converter where it is distribut by distributors 310, intermediate branch tubes 320, and distributors 322 to the vertical tubes 324. The gas is heated as it flows upward by heat exchange through the tub and distributor walls. From the exits 334 of tubes 324 at the top of the vessel, the reactant gas pass downward, through the catalyst, which is packed in the space between the tubes 324. The gas flow through the tubes 324 is countercurrent to the gas flow through the catalyst bed. The temperature profile against methanol concentration is quasi-isothermal and is very favorable in terms of reactio rates and conversion.

The main features of the tube cooled converter are: (1) The design is mechanically simple with no high differential pressure across the tube material and w> tubesheet construction problems. (2) Catalyst loading is simple. (3) Λ single reactor can be constructed capable producing methanol at high capacity of over 2000 tons

(1,815,000 kg) per day. (4) Methanol synthesis loop circulating gas rates are reduced when compared to a conventional quench type of reactor. (5) Heat recovery i simplified; boiler feedwater preheating, steam raising or feedstock saturation may be employed. (6) Control is simplified. (7) Catalyst volume is reduced.

From the tube cooled converter 312, the effluen gas 330 passes to the shell side of the forced film saturator 204, where the gas is cooled, for example to a temperature of about 340°F (170°C), by heat transfer to t circulating water and incoming natural gas stream flowing through the tubes 210. The converter effluent gas emerge from the saturator in line 336 and passes to the heat exchanger 300 where the effluent gas is further cooled by heat exchange with the loop feed for the converter. Then the gas from the converter is fed over line 338 to loop condenser 340 which is cooled by water to condense the methanol in the methanol loop stream. Condensed methanol separated from the methanol loop in separator 342, and passed over line 344 to pressure letdown vessel 346 befor being passed in crude methanol stream 340 to the distillation section 350 in FIG. 15.

The gas stream 354 from the separator is split into a recycle stream 356 passing to the suction inlet of the circulator 302 and a purge stream 350 taken through valve 360 to maintain the concentration of inert material the methanol loop at acceptable levels. Hydrogen is recovered from the purge gas stream 350 by the pressuj.e swing adsorption unit 286 to generate the hydrogen recycl stream 284 used to improve stoichiometry of the methanol loop stream. The flash gas 228 is taken through valve 26 and the purge gas 230 is taken from the pressure swing adsorption unit 286 to form a portion of the fuel to the fired heater 224.

In the distillation section 350, FIG. 15, the crude methanol 340 is passed to a distillation column 364 where light end materials such as absorbed g ses are removed. The overhead 366 om the column 364 is passed through water cooled condensers 368 and 370 with the condensate passing to drum 372 from which it is drawn by pump 374 to form a reflux s ream for the column. The noncondensed light ends 232 are passed as fuel to the fire heater. Bottom streams frcm the column 364 are recycled through reboilers 266 and 374 heated by the synthesis gas stream and steam, respectively. The methanol stream 370 i passed by pump 380 to distillation column 382 where the stream is separated into the fusel oil stream 226, a water stream 384, and the product methanol stream 386. The distillation column has retailers 270 and 388 heated by th synthesis gas stream and steam, respectively. The overhea 389 is withdrawn under vacuum produced by vacuum pump syst 390. The overhead vapors ere condensed by water cooled condenser 392 and separated in reflux drum 394 before bein passed by pump 396 back to the distillation column as a reflux stream. The vacuum pump system 390 includes pump 3 circulating water from collector 400 through water cooled heat exchanger 402. Excess water is passed by line 404 fr the collector 400 back to the bottom of the distillation column 382. The water stream 304, withdrawn by pump 406 from the bottoms of column 302, is further cooled by heat exchanger before discharge in line 410.

A process flow diagram of a methanol plant designed for low capital cost is shown in FIGS. 16 and 17. Hydrocarbonaceous feedstock, such as natural gas stream 4 FIG. 17, is passed to a feedstock saturator 422 which is similar to the saturator 204 in the embodiment shown in F 14. Water is recirculated by pump 424 to the top of the forced film saturator where it flows with the feedstock

stream through the tubes of the saturator. The saturated feedstock is separated from the water in the bottom of th saturator to form saturated feedstock stream 425 and the water recycle stream. Makeup water for the recycle water stream comes from synthesis gas condensate 426 and distillation bottoms 428. Methanol converter effluent stream is fed to the shell side of the saturator to heat water and feedstock stream. Line 430 provides Cor blowdo of the saturator water. Λs shown in FIG. 16, the saturated feedstock 42 passes through coil 434 of a fired heater 436 where the saturated feedstock is preheated by combustion of waste f 438 with air 439. The heated saturated feedstock on line 440 from the heater 436 is fed to the catalytic partial oxidation (CPO) reactor 28 where it is mixed with an oxyg or oxygen-containing stream 442 and subjected to catalyti partial oxidation to produce synthesis gas as has been described hereinbefore.

The reactor effluent 444 is first cooled by generating steam in a boiler 446 which receives water fro steam drum 448 and returns steam to the drum. Λ second steam drum 450 has water flow boiled in coils 452 of the fired heater 436. Steam outputs 454 and 456 of the steam drums 448 and 450 are combined to produce the steam 458 u in the process. Lines 460 and 462 provide for boiler blowdown from the drums 448 and 450. Boiler feedwater 464 branches into streams 466 and 468 feeding the respective drums 448 and 450. The branch 466 is preheated in heat exchanger 472 by the reactor effluent in line 470 from th boiler 446 before passing in line 474 to the drum 448. T synthesis gas in line 476 from the heat exchanger 477 is further cooled by water in condenser 478 and then passed through line 480 to synthesis gas separator 402 where wat

condensate 484 is removed by pump 486 to provide the condensate stream 426.

After water removal, the synthesis cjas in line 4' is compressed to the methanol loop pressure by compressor 492 to produce feed 494 to the desulfurization vessel 202. The output 496 of the desulfurization vessel 202 forms the make-up gas for the methanol loop of FIG. 17.

The make-up synthesis gas 496, in FIG. 17, joins the gas circulating in the methanol synthesis loop at the discharge 498 of the methanol loop circulator 500. The lo stream 502 is preheated in heat exchanger 504 by steam before passing in line 506 to tube cooled methanol convert 508 which substantially similar to tube cooled methanol converter 312 of FIG. 14. Line 510 from the loop heater 5 provides heated loop gas to the top of the converter 508 during initial start-up of the methanol converter.

From the tube cooled converter 508, the effluent gas 512 passes to the shell side of the forced film saturator 422, where the gas is cooled by heat transfer t the circulating water and incoming natural gas stream flowing through the saturator. The converter effluent gas emerges from the saturator in line 514 and passes to wate cooled condenser 516 to condense the methanol in the methanol loop stream line 518. Condensed methanol is separated from the methanol loop in separator 520 and pas over line 522 and valve 524 to pressure letdown vessel 57.

The gas stream 528 from the separator 520 is εp into a recycle stream 530 passing to the suction inlet of the circulator 500 and a purge stream 532 taken through valve 534 to maintain the concentration of inert material the methanol loop at acceptable levels. Flash gas 5 K» through valve 530 from the let-down vessel 526 is combine with the purge stream 532 to form the waste fuel stream 4 to the fired heater 436.

From the let-down vessel 526, crude methanol 54 is passed to a distillation column 542 where the stream i separated into a light ends stream 544, a water stream 54 and the product methanol stream 540. The distillation column has a reboiler 550 heated by steam. The? overhead is cooled by water cooled condenser 554 producing a condensed methanol stream 556, a portion of which is used a reflux stream through pump 558. The remaining portion the condensed methanol stream 556 is combined with the fu oil fraction 560 to form the product stream 548. The wat stream 546 from the bottoms of column 542, is cooled in cooler 562 and passed by pump 564 to the stream 428 fee i the saturator 422.

The following examples are intended to illustra further the invention described herein and are not intend to limit the scope of the invention in any way.

EXAMPLE I A methanol plant in accordance with FIGS. 13, 1 and 15 is operated to produce 2000 tons (1,800,000 kg) of methanol per day. The following TABLES I, II, and III se forth moles/hour, mole percent, and parameters of pressur temperature, water/steam, and heat transfer for the plant. The moles/hour are lb moles/hour (0.4536 kg moles/hour).

TABLE I Moles / Hour SCRIPTION CO C02 H2 CH4 C2H6 N2 Ar 02 H20 CH30H C4F ^Q 0H CH30CH3 TOTALS tural Gas 3.7 6236.0 9.4 6.3 6255.4

(line 200) . Nat. Gas 3.7 6236.0 9.4 6.3 8303.7 14559.1

(line 220) ygen 16.0 3186,1 3202.1

(line 236) O out 4017.5 1652.9 12254.2 591.2 6,3 15.0 7360.2 25908.3

(line 244) op Make-Up 4019.6 1653.8 12369.0 591.2 6.3 16.0 30.8 19186.7

(line 298) nverter In 9803.5 10861.9 62624.& 35614.5 407,6 946.0 55.0 492.2 1.0 30.7 120836.3

(line 310) J CO nverter Out 5876.1 9451.9 50539.5 35614.5 407.6 946.0 1493.5 5737.0 8.1 37.8 110162.0

(line 330) p. Liquid 4.4 103.5 34.5 63,2 0.3 2.0 1463.9 5287.3 7.1 5.6 6977.8

(line 344) p. Gas Out 5871.7 9348.4 50505.0 35551.3 407.3 944.0 24.6 499.7 1.0 31.2 103134.2

(line 354) op Recycle 5783.9 9208.1 49755.4 35023.3 401.3 930.0 24.2 492.2 1.0 30.7 101650.1

(iine 355) rge Gas 87.8 140.3 749.5 528.0 6.0 14.0 0.4 7.5 0.5 1534,1

(line 358)

(line 223) ude Methanol 4 . 0.4 4.7 0.1 1453.5 523C.3 6.2 5310.5

(line 343) (M02)

TABLE I (CONTINUED)

Moles / Hour SCRIPTION CO tal Fuel 80.1

(line 225) mb. Air

(line 227) ater Flue

(233) ght Ends 0.8

(line 232) sel Oil

(line 226) stillation In

(line 378) lumn Bottoms

(line 410) oduct Methanol

(line 386)

TABLE I I

Mole Percent

C2H6 N2 Ar 02 H20 CH30H C4H90H CH30CH3 0.15 0.10 0.06 0,04 57.04

0.50 99.50

0.02 0.06

0.03 0.08

0.34 0.78 o

0.37 0.86

0.03

0.39 0.92

0.39 0.92

0.39 0.92

0.65 1.51 0.04 0.81 0.05

0.13 1.14 0.13 3.39 0,24

21.55 77.54 0.09

TABLE I 1 (CCNT I NUED)

Mole Percent SCRIPTION CO C02 H2 CH4 C2H6 tal Fuel 7.02 18.92 13.97 46.05

(line 225) mb. Air

(line 227) ater Flue 10.39

(233) ght Ends 0,75 52.77 0.50 5.90

(line 232) sel Oil

(line 226) stillation In

(line 378) lumn Bottoms

(line 410) oduct Methanol 100.00

(line 386)

TABLE I I I

Parameters ESCRIPTION Pressure Temperature Water/Steam Net Heat Transfer PSIG KPa DEG F DEG C LBS/HR KG/HR MMBTU/HR KCAL 10 ⁶ /HR atural Gas 200 450 3103 aturator 204 176.4 44.5 ake-up water 216 157040 71230 lowdown 218 7450 3380 turated Gas 220 444 3061 il 222 114.3 28.8 sel oil 226 ash gas 228 A purge 230 ght ends 232 ated Feedstock 234 ygen 236 535 _, 3689 ater 240 5.5 1.4 _e» O out 244 413 2348 t at exchanger 246 46.8 11.8 eam Drum 248 1550 10637 ater coils 252 65.5 16.5 ter 254 305530 138590 ater coils 258 68.7 17.3 eam 260 1500 10342 owdown 262 3030 1370 ocess stream 264 boiler 266 ocess stream 253 boiler 270 ocess stream 272 at exchanger 274 ocess stream 275 ndenser 273 densate 231 132040 59390 ocess stream 232 403 2779

TABLE ill (CCNTINUED)

Parameters SCRIPTION Net Heat Transfer MMBTU/HR KCAL 10 ⁶ /HR

drogen recycle 284 ke-up gas 293 rculator out 304 ocess stream 306 at exchanger 308 186.0 45.9 nverter feed 310 nverter out 330 ocess stream 336 ocess stream 338 ndenser 340 153.9 33.5 ude methanol 343 erhead 366 ndenser 368 37.5 9.5 J ndenser 370 7.0 1.3 boiler 376 noraaLiv nil densate stream 384 175 80 oduct methanol 386 105 41 boiler 388 31.9 3.0 erhead 389 34 104 40 denser 392 205.3 51.7 t exchanger 408 1.9 0.5 umn bottoms 410 100 38

EXAMPLE II Λ methanol plant in accordance with FIGS. 16 and 17 is operated to produce 102,000 lbs. (46,400 kg) of methanol per day. The following TABLES IV, V, and VI set forth moles/hour, mole percent, and parameters of pressure, temperature, water/steam, and heat transfer for the plant. The moles/hour are lb moles/hour (0.4536 kg moles/hour).

TABLE IV

Moles / Hour

Ar CH4 C2H6+ CH30H C2H50H+ 160.56 1.23 150.56 1.28

10.40

10.40

454.00 16.99

454.00 149.24 0.13

453.50 17.37

443.60 16.99

9.90 0.33

0.39 0.12

10.29 0.50

0.11 131.75 0.13

TABLE IV (CCNTINUED)

Moles / Hour ESCRIPTION CO C02 H2 N2 02 Ar CH4 C2H6+ CH30H C2H50H+ CH30CH3 H20 TOTALS omb. Air 166.70 44.72 2,03 6.27 219.72

(line 439) ngress Air 15.15 4.06 0.13 0.57 19.96

(line 441) eater Flue 29.44 193.60 8.13 2.21 58.94 292.32

(443) ight Ends 0.02 1.72 0.03 0.02 0.11 0.15 0.13 2.18

(line 544) usel Oil 2.25 0.06 0.95 3.26

(line 560) olumn Bottoms 0.08 0.05 24.92 25.05

(line 423) roduct Methanol 129.27 0.07 0.0?. 1.21 130.57

(line 557) lended Product 131.52 0.13 2.16 133,83

(line 543)

TABLE V

Mole Percent

Ar CH4 C2H6+ CH30H C2H50H+ CH3OCH3 H20 97.80 0.78 47.17 0.37 51.77

1.58 2.08 12.64 O.Λ7 -C 13,65 4.49 0.01 14.35 0.55 14.35 0.55 14.35 0.55 15.18 4.57 14.38 0.70 0.07 81.81 0.1]

TABLE V (CCNTINUED)

Mole Percent

CO CO 2 H2 N2 02 Ar CH4 C2H6+ CH30H C2H50H+ CH30CH3 H20

75.87 20.36 0,92 0.57

75.87 20.36 0.92 0.57

10.07 66.23 0.76 2.78 20.15

0.92 78.90 1.38 0.92 5.05

CO

TABLE V I

Parameters RIPTION Pressure Temperature Water/Steam Net Heat Transfer PSIG KPa DEG F DEG C L3S/HR KG/HR MM3TU/HR KCAL 10 ⁶ /HR ral Gas 420 500 3450 60 16 rator 422 4.12 1.04 ensate 426 2840 1290 oms water 428 455 206 down 430 121 55 rated Gas 425 430 2960 390 200 434 2.78 0.70 e Fuel 438 104 40 ed Feedstoc 1100 593 en 442 285 141 out 444 375 191 ____. er 446 6.36 1.5C s 452 3.81 0.95 m 454 491 255 m 455 491 255 er Blowdown er Blowdown er Feed Wate 275 135 er Feed 456 er Feed 468 ess Stream 470 550 288 Exchanger 472 1.17 0.29 er Feed 474 415 213 ess Stream 475 315 157 enser 478 3.93 C.99 ess Stream 480 104 40 hesis Gas 4?o 357 2590 104 40 ressed Syngas 4 904 6230 310 154 -Up Gas 496 394 515C 310 154

TABLE V I (CONT I NUED)

Parameters SCRIPTION Pressure Temperature Water/Steam Net Heat Transfer PSIG KPa DEG F DEG C LBS/HR KG/HR MMBTU/HR KCAL 10 ⁶ /HR rculator Output 498 894 6160 115 46 op Heater 504 4.85 1.22 nverter Feed 506 nverter Out 512 ocess Stream 514 ndenser 516 9.60 2.42 ndenser Output 513 ash Gas 536 ude Methanol 540 st. Column 542 ght Ends 544 σ ended Product 543 boiler 550 4.40 1,10 erhead 552 35 164 73 denser 554 4.20 sel Oil 560 194 90 ttoms Cooler 562 0.08 o.c;

Since many modifications, variations, and change in detail may be made to the above described embodiments without departing from the scope and spirit of. the invention, it is intended that all matter described above and shown in the accompanying drawings be interpreted as illustrative and not in a 1uni ing sense.