Comparative Assessment of Catalytic Partial Oxidation and Steam

Ind. Eng. Chem. Res. 1995,34, 1037-1043

1037

Comparative Assessment of Catalytic Partial Oxidation and Steam Reforming for the Production of Methanol from Natural Gas Narendra Dave CSIRO Division of Coal and Energy Technology, PMB 7, Menai, New South Wales, 2234, Australia

Gary A. Foulds* Chemical Engineering & Industrial Chemistry, Department of Molecular Sciences, James Cook University, Townsville, Queensland 481 1, Australia

A commercial-scale (1825tons/day) methanol plant, operating a t steady state, that uses catalytic partial oxidation (CPO) as a primary route for synthesis gas production has been simulated using the ASPEN/SP process simulator. The overall energy and utility requirements for such a plant have been calculated and compared over different operating conditions. It has been shown that the overall energy demand per ton of methanol produced is significantly affected by CPO reactor pressure, steam-to-carbon ratio in the feed, and oxygen-to-carbon ratio in the feed. The performance of this plant is then further compared with similar scale methanol plants using conventional combined reforming, Imperial Chemical Industries (ICI) gas heated reforming, and Exxon's all-in-one type technologies for syngas production. Results show that CPO rates equivalently with these competing technologies on the basis of equivalent conversion.

Introduction

CH4

Many natural gas deposits, including gas recovered together with crude oil, marginal gas fields in remote locations including off-shore sites, and gas drained from coal deposits, are geographically' located so as to preclude conveying the gas by pipeline, which can be considerably more costly than transporting liquid product to suitable processing facilities. Thus, the challenge to convert remote natural gas into liquid product such as methanol is of great importance and creates an ideal opportunity for on-site conversion. Methanol, apart from being an ideal liquid product for transportation from remote sites, has the potential to be used as a neat fuel or fuel additive for automotive use (Morton et a1 1990). Environmental controls are already in place in most of the industrialised nations and emission control legislation has already stimulated the demand for methanol, principally for the production of methyl tert-butyl ether (MTBE) and methanol fuel blends. It has been estimated that if methanol makes a modest 10% penetration into the automotive fuel market in the USA, the increase in methanol demand would be of the order of 25 billion gallons per annum (Dautzenberg, 1990). The most common and industrially favoured method for converting natural gas to methanol is via the intermediate production of synthesis gas (syngas). Syngas is usually produced by steam reforming or combined reforming, although other novel ways of syngas production such as the all-in-one Exxon process (Goetsch and Say, 19901, ICI's gas-heated reforming (Lywood, 1989, Abbott et al., 1989), and catalytic partial oxidation (CPO)as depicted by Korchnak and Dunster (1990)have been proposed. Although conventional steam reforming is a reliable and established technology, it involves a highly endothermic reaction (1) and as a result is an energy-

* To whom correspondence should be addressed. E-mail: gary [email protected]. Q888-5885I95/2634-1Q37$09.QOfQ

+ H 2 0- CO + 3H2

AH = +49.3 kcaVmol (1)

intensive process which usually utilizes a nickelcontaining catalyst operating at temperatures in the range 650-1000 "C and at pressures of approximately 10-30 atm. Other disadvantages are that it requires a separate fuel stream and a C 0 2 import if a syngas stoichiometric number {SN = (H2 - CO2)/(CO C02)) close to 2 is required. The net result is a large furnace chamber, with associated high plant capital costs which, together with concomitant metallurgical constraints, contribute to the unsuitability of conventional steam reforming for off-shore application. Apart from conventional steam reforming, syngas can be produced from natural gas in a number of ways, including autothermal reforming, combined primary and secondary reforming and, more recently, catalytic partial oxidation. While most syngas is produced by steam reforming, with the other routes still being essentially at the research development stage, some of these other routes may be more attractive, depending on factors such as H2:CO ratio, i.e., downstream use, product purity, i.e., the presence of N2, COS,HzO, and CH4, plant capacity, feedstock availability, purity and cost (including0 2 , C02, HzO), and byproduct credit (Gaff and Wang, 1987; Michel, 1989). Syngas generation by combined reforming involves a primary steam reformer and a secondary oxygen reformer in series. This arrangement shifts a proportion of the reforming duty away from the primary reformer, reducing the required size of the primary reformer and the severity under which it operates. This route, typified by the Lurgi process (Farina and Supp, 1992), produces stoichiometric syngas for methanol production but requires an oxygen plant. IC1 (Lywood, 1989; Abbot et al., 1990) have recently patented a new version of the combined reforming process, in which heat required for the primary steamreforming state is supplied directly by heat exchange with the reformed gas from the secondary oxygen reformer. It is estimated that this novel way of combin-

+

0 1995 American Chemical Society

1038 Ind. Eng. Chem. Res., Vol. 34, No. 4,1995 ing the two reformers reduces the size of the syngas generation unit to approximately one-quarter of an equivalent conventional steam-reforming unit and should thus reduce capital costs further. This process has been proven commercially for ammonia production using air as oxidant, but awaits commercial demonstration in Australia (The Melbourne Age, 1991), for methanol production using oxygen as the oxidant. An all-in-one combined process, wherein steam reforming is combined with partial oxidation of natural gas in a single fluidized bed reactor, has been recently developed by Exxon (Goetsch and Say, 1990). In this process, natural gas and steam enter at the bottom of the fluidized bed reactor, while oxygen is separately sparged part way up the reactor into the fluidized catalyst bed of nickel supported on a-alumina. It has been suggested that reactor operation at steam to carbon ratios of as low as 0.5 is possible, with deposited carbon on the circulating catalyst being burned off by oxygen. In addition, the problem of avoiding the reverse methanation reaction over entrained catalyst fines appears to have been solved by rapid quenching of the syngas below the reaction threshold temperature. This process is at the process demonstration unit stage. Partial oxidation of methane to syngas, using oxygen as oxidant, represents an alternative to steam reforming. Unlike steam reforming, the partial oxidation of natural gas to syngas is exothermic, more selective, and theoretically yields a syngas stoichiometric number of 2 (2). CH,

+ 1/202 - CO + 2H,

AH

= -8.5 kcal/mol

(0)

Most of the partial oxidation processes that have been employed commercially are noncatalytic processes, but the high temperatures (1250- 1500 "C) and oxygen-tocarbon ratios used to date, have restricted its general application (Kirk and Othmer, 1980). However, as reported in a recent review of catalytic partial oxidation (Foulds and Lapszewicz, 19941, a number of studies have been published recently on the conversion of natural gas to syngas via the CPO route, using highly active catalysts based on Ni, Co, and Pt group metals (Green, 1990; Vernon et al., 1992,1990;Ashcroft et al., 1990; Kunimori et al., 1992; Choudhary et al., 1992a, 199213, 1992c, 1992d, 1992e; Lapszewicz and Jiang, 1992; Vermeiren et al., 1992). In addition, the review also reports the work of Hickman and Schmidt (1992a, 1992b, 19931, Hickman et al. (19931, and Hochmuth (19921, who have examined the CPO reaction using monolith-supported catalysts. The review also suggests that steam reforming could be replaced by CPO, using these new catalysts, which convert natural gas to syngas in the presence of steam and oxygen, using process conditions similar to those employed in this simulation study. The CPO process is exemplified by Davy McKee's fixed bed reactor (Korchnak and Dunster 1990), which utilizes a series of monolithic supported catalyst beds. As described in their patent, reactions are essentially restricted to partial oxidation and water gas shift, with little or no steam reforming occurring at the high space velocities employed (20 000 h-l). Steam is used mainly as a diluent in this process, in which natural gas and oxygen are fed separately to the reactor, well mixed in a mixing zone, before being fed to the monolith catalyst zones at high superficial linear velocities, using a distributor to avoid flashback.

Several investigators have offered economic analyses of these new developments and suggest that this new route to methanol requires 10-15% less energy and approximately 2 5 3 0 % less capital investment (Korchnak and Dunster, 1987; Farina and Supp, 1992). Unfortunately, while some details on process economics are supplied, sensitivity analyses for the overall plant have not been reported. In this paper the advanced process engineering simulator, ASPENISP, has been used to simulate a methanol plant based on catalytic partial oxidation (CPO) of natural gas to syngas and its further conversion to methanol using a steam-raising converter. It is the aim of this study to assess the energy and utility requirements of a commercial scale methanol plant (1825 tondday of 99.5% pure methanol) that uses CPO of natural gas for syngas generation and to compare them over a range of different operating conditions including reactor pressure, oxygen-to-carbon ratio, and steam-to-carbon ratio. To complete the assessment of the catalytic partial oxidation route, the performance of this plant is compared with the same scale methanol plants employing combined reforming (Farina and Supp 1992, Schneider and LeBlanc 19921, ICI's gas heated reforming (Lywood, 1989; Abbott et al., 19901, and Exxon's all-in-one fluidized bed combined process (Goetsch and Say, 1990)technologies for syngas production. Noncatalytic partial oxidation was not included for comparison, mainly due to the high reaction temperatures employed. In addition, Farina and Supp (1992)have indicated that the total energy requirement (GJ/t of methanol) for noncatalytic partial oxidation is approximately 8% higher than that for combined reforming. The process data used in these simulations were retrieved from the listed references or generated mathematically. Part of this work was presented at the Natural Gas Conversion Symposium in Sydney Australia (Dave and Foulds, 1993).

Basis for Evaluation For the purposes of the simulation, the following assumptions were made: Natural gas (CHI (90.7 mol %}, C2Hs (5.89 mol %}, C3Hs (0.7 mol %}, COz (1.81 mol %}, Nz (0.9 mol %}) is available at 2800 kPa and ambient temperature, while oxygen (99.5%)is available at ambient temperature and CPO reactor pressure. Steam participates in the CPO reaction, and the reaction achieves thermodynamic equilibrium instantly. Saturated steam at CPO reactor pressure is available from the heat recovery steam generation unit within the plant. The optimum steam to carbon ratio for the CPO reaction is that which gives a syngas stoichiometric number as close as possible to 2 for the syngas. There is no carbon deposition on the CPO catalyst due to Boudouard type reactions. The methanol synthesis reactions approach equilibrium at 8100 kPa and at temperatures between 252 and 257 "C, and the methanol synthesis catalyst can tolerate up to 15 mol % CO2 in the syngas. The heat released during methanol synthesis is utilized in raising saturated steam at 4000 kPa. Raw methanol is purified to 99.5% purity level in a double-distillation column system. All the drivers for the compressors, oxygen plant, and boiler feed water pumps are steam driven turbines. Energy recovery from the high-pressure purge gas leaving the methanol loop by gas expansion prior to combustion is not relevant.

Ind. Eng. Chem. Res., Vol. 34, No. 4, 1995 1039

Natural Gas

b

FEED COMPRESSION & PRETREATMENT

T*

HEAT RECOVERY + & S T E A M .

CPO REACTOR

SYNQAS COMPRESSION

QENERATION

A Air

AIR

b

Oxygen

T

b

SEPARATION UNIT

z

8

3

i

METHANOL

F

REACTOR

u)

Methanol 4

The simulation flow sheet for the methanol plant was essentially divided into four main sections, i.e., syngas production, methanol production, methanol purification, and the heat recovery steam generation unit (HRSG). Simulations were carried out for a methanol plant using CPO syngas generation at various pressures and oxygento-carbon and steam-to-carbon ratios. The optimized CPO/methanol system was then compared with simulations of methanol plants using conventional combined reforming, ICI's gas-heated reformer process, and Exxon's all-in-one combined reforming process. Methanol Plant with CPO. Figure 1is a simplified process block diagram for an entire methanol plant, in which CPO is used as the syngas generation step. In the proposed flow scheme, natural gas is compressed to 3500 kPa and then preheated to 377 "C to remove sulfur compounds in a zinc oxide bed. The desulfurized gas is mixed with the required quantity of saturated steam, available from the HRSG unit at the CPO reactor pressure. The gas mixture is then further heated to 500 "C before entering the CPO reactor where 99.5% purity oxygen, preheated to 200 "C, is injected separately. The CPO reactor is simulated as an adiabatic equilibrium reactor. Heat is recovered from the reactor exit stream by raising high pressure steam and by preheating the feed streams to the reactor, as well as boiler feed water. After removing the liquid condensate, syngas is further compressed to the methanol synthesis pressure, preheated to the methanol reactor temperature, and passed through the methanol reactor. The methanol reactor is simulated as an isothermal equilibrium reactor using ASPEN/SP's REQUIL model, in which the following reactions are assumed to take place at 8000 kPa and 255 "C:

+ 2H, = CH,OH

(3)

* purge G~~

+ 3H, = CH,OH + H,O CO + H,O = CO, + H,

CO,

ASPEN Simulation of Processes

CO

RAW METHANOL SEPARATION

METHANOL DISTILIATION

(4) (5)

The methanol reactor temperature is controlled by raising high-pressure steam at approximately 4000 kPa. Methanol is condensed from the product gas mixture, collected, and sent to the purification unit where unreacted gas is recycled. A small quantity of gas is purged as a fuel gas to prevent buildup of inerts such as nitrogen, argon, etc., in the system. Further details of the CPO reactor and methanol synthesis loop have been described elsewhere (Korchnak and Dunster, 1990; Supp, 1989). The air separation unit that provides 99.5% pure oxygen to the CPO reactor was considered to be operationally independent, except that its power requirement is drawn from the same source as the CPO reactor and methanol synthesis sections. This energy requirement was found to be a function of supply pressure as well as the installed capacity of the oxygen plant, and was calculated using correlations provided by an industrial supplier of oxygen plants (Saunders 1992). Methanol Plant with Combined Reforming. In a combined reforming process of the Lurgi type (Farina and Supp, 1992), part of the desulfurised natural gas is fed to a smaller primary reformer operating at approximately 780 "C. The partially reformed gas leaving the primary reformer is mixed with the remainder of the natural gas feed, and the mixture is then fed along with 99.5% pure oxygen t o a secondary reformer opeating at approximately 950 "C. Further details of the process have been described elsewhere (Farina and Supp, 1992; Schneider and LeBlanc, 1992). The primary and secondary reformers for this process were simulated as nonadiabatic equilibrium and adiabatic equilibrium reactors respectively using ASPEN/ SP's RGIBBS model. The heat duty for the primary

1040 Ind. Eng. Chem. Res., Vol. 34, No. 4, 1995

reformer was supplied via the complete combustion reaction of natural gas used as fuel. The composition of this fuel gas was assumed to be the same as the natural gas feed. The ASPEN simulations for methanol synthesis and the oxygen supply sections for syngas generation are essentially the same as those described for the methanol plant using CPO. Methanol Plant with ICI's Gas Heated Reformer Process. In this process, the desulfurised natural gas is first saturated with water to satisfy process steam requirements and then preheated to the primary reformer inlet conditions by contacting with hot syngas (approximately 500 "C) leaving the shell side of the primary reformer. The preheated feed gas is then partially steam reformed under relatively mild conditions of about 730 "C on the tube side of the primary reformer, the heat being supplied by hot reformed gas on the shell side of the primary reformer, which comes from the exit of the secondary reformer, which is operating under more severe conditions of just over 1000 "C. The feed to the secondary reformer consists of partially reformed gas from the primary reformer and oxygen from an oxygen plant which is supplied at approximately 200 "C. The amount of oxygen fed is sufficient to complete the conversion of the partially reformed natural gas and to adjust the syngas stoichiometric number of the final syngas as close as possible to 2. The main advantages of this process are negligible pressure differential across steam reformer tubes, highpressure operation, and control of the HdCO ratio, all suitable for methanol production. Further details of the process conditions for syngas generation using this system are given by Lywood (1989). The ASPEN simulations for methanol synthesis and purification and the oxygen supply sections for syngas generation are essentially the same as those described for the methanol plant using CPO. Methanol Plant with Exxon's All-in-One Combined Reforming Process. The operating parameters for this process were obtained from the open literature (Goetsch and Say, 1990). In this case, the desulfurised natural gas is mixed with steam and preheated to 500 "C, before being fed to a fluidized bed reactor operating at 980 "C and 2500 kPa. The reactor temperature is controlled by moderating the pure oxygen feed at 200 "C, which is fed into the fluidized bed partway up the reactor. The effluent from the reactor is rapidly quenched below 650 "C, by raising steam, to avoid the reverse methanation reaction in the presence of catalyst fines. The ASPEN simulations for methanol synthesis and purification, and the oxygen supply sections for syngas generation are essentially the same as those described for the methanol plant using CPO.

Results and Discussion Effect of CPO Reactor Pressure. Considering the catalytic partial oxidation step in isolation (see eq 2), the reaction is exothermic and also results in a molar volume increase. It is clear that for greater conversion, this reaction should be conducted at low pressure and low adiabatic temperature rise across the reactor. The quality of syngas, as indicated by the rising syngas stoichiometric number, also improves at lower pressure. However, the methanol synthesis reactions (see eqs 3 and 4) are favored by operation at higher pressure,

Table 1. Effect of Catalytic Partial Oxidation Pressure CPO reactor pressure (kPa) CPO reactor inlet temp ("C) CPO reactor exit temp ("C) oxygen-to-carbon ratio steam-to-carbon ratio conversion of natural gas (96 by mol) selectivity t o carbon monoxide (% by mol) syngas stoichiometric no. HdCO molar ratio of syngas COZmolar % in syngas (dry basis) CPO compressor load (MW) methanol synthesis pressure (kPa) methanol synthesis temp ("C) recycle ratio purity of raw methanol (% by wt) methanol plant compressor load (MW) oxygen plant power load (kW) oxygen per unit of 99.5% methanol natural gas per unit of 99.5% methanol demineralized water per unit of 99.5% methanol net energy required (GJm of methanol)

3200 500 842 0.465 2.0 80.75 60.64 1.726 4.014 13.38 0.457 8100 255 3.19 79.9 19.6 26.03 0.71 0.80 0.597

2150 500 818 0.465 2.0 83.25 60.0 1.760 4.115 13.48 -0.381 8100 255 3.18 79.2 24.59 24.13 0.674 0.761 0.607

42.79 41.19

1125 500 781 0.465 2.0 86.98 58.81 1.808 4.287 13.70 0.0 8100 255 3.08 77.1 37.63 21.92 0.629 0.71 0.765 45.85

indicating a conflict between syngas generation by CPO and methanol synthesis. Table 1 lists data relating the effect of CPO reactor pressure on the overall performance of a proposed 1825 tondday methanol plant. A number of features are evident. For a constant conversion of natural gas, the oxygen requirement increases steadily with an increase in CPO reactor pressure, when the steam-to-carbon ratio is held steady. In turn, higher oxygen input increases the adiabatic temperature rise due to the increased contribution of complete combustion reactions of natural gas to COn and H20. This is reflected by decreasing HdCO ratio and syngas stoichiometric number. The increase in adiabatic temperature rise with pressure becomes excessive at pressures above 3200 kPa, where the reactor exit temperature exceeds 1000 "C, even with the high steam-to-carbon ratio of 2. This could cause problems in terms of materials of construction for the CPO reactor and other downstream processing equipment. In addition, it should also be noted that as a result of decreasing syngas stoichiometric number, the natural gas requirement per ton of methanol also increases, despite conversions of over 98% being achieved at high pressure. Against the above-mentioned drawbacks of increased oxygen plant capacity, natural gas requirement, and mechanical problems, the major advantage of highpressure operation is that the syngas compression load in the methanol synthesis loop decreases rapidly with an increase in pressure. In fact, the energy requirement within the methanol plant, in terms of syngas and recycle gas compression load, increases rapidly below 2100 kPa, and significantly raises the net energy requirement per ton of methanol produced. Thus, the overall effect of increasing the CPO reactor pressure is a reduction in energy for the entire plant, per ton of methanol produced. Effect of Oxygen Addition at Constant Steamto-Carbon Ratio. Table 2 lists data which illustrate the effects of increasing the oxygen to carbon ratio at a constant steam-to-carbon ratio of 2, constant CPO reactor pressure of 3200 kPa, and constant reactor inlet temperature of 500 "C. Clearly, as the oxygen to carbon ratio is increased, both natural gas conversion and reactor outlet temperature increase. CO selectivity also increases, but the H2/ CO ratio and syngas stoichiometric number decrease. Overall, the net energy requirement per ton of methanol

Ind. Eng. Chem. Res., Vol. 34,No. 4, 1995 1041 Table 2. Effect of Oxygen Addition at Constant Steam-to-CarbonRatio CPO reactor pressure (kPa) CPO reactor inlet temp ("C) CPO reactor exit temp ("C) oxygen-to-carbon ratio steam-to-carbon ratio conversion of natural gas (% by mol) selectivity to CO (% by mol) syngas stoichiometric no. HdCO molar ratio of syngas C02 molar % in syngas (dry basis) CPO compressor load ( M W ) methanol synthesis pressure (@a) methanol synthesis temp ("0 recycle ratio purity of raw methanol (% by wt) methanol plant compressor load (MW) oxygen plant power load (kW) oxygen per unit of 99.5% methanol natural gas per unit of 99.5% methanol deminerliazed water per unit of 99.5% methanol net energy required (GJm of methanol)

3200 3200 3200 3200 500 500 500 500 842 867 913 983 0.465 0.500 0.550 0.600 2.0 2.0 2.0 2.0 80.75 86.42 93.50 98.14 60.64 63.38 67.07 70.58 1.726 1.723 1.706 1.662 4.014 3.759 3.437 3.129 13.38 12.94 12.30 11.66 0.457 0.421 0.382 0.364 8100 8100 8100 8100 255 255 255 255 3.19 2.85 2.36 1.91 79.9 80.9 82.2 83.6 19.60 16.71 13.54 11.60 26.03 25.75 25.70 26.71 0.710 0.701 0.7 0.726 0.801 0.736 0.669 0.635 0.597 0.517 0.426 0.344 42.79 39.34 36.91 36.30

decreases, since the increased natural gas conversion reduces the gas compression and methanol refining loads. However, this decrease becomes less pronounced a t oxygen-to-carbonratios above 0.55. In addition, the teactor exit temperature approaches 1000 "C at oxygen to carbon ratios above 0.55. This, together with an operating pressure of 3200 kPa introduces materials of construction considerations, similar to those encountered in conventional steam reforming. Effect of Steam Addition at Constant Oxygento-Carbon Ratio. Table 3 lists data obtained when the steam flow rate is increased, using a constant oxygento-carbon ratio of 0.465, a constant CPO reactor inlet temperature of 500 "C, and a constant CPO reactor pressure of 3200 kPa. Increasing the steam flow rate moderates the adiabatic temperature rise across the reactor, while having a small detrimental effect on conversion. In addition, when catalysts that are prone to carbide formation are used, e.g., supported Ni, the steam cofeed performs an additional function in limiting the formation of carbon on the catalyst surface. Data relating 02/CH4 ratio to steadCH4 ratio have been presented by Korchnak and Dunster (1990),in which the regions of carbon formation are shown for a nickel catalyst. Thus, for example, at 950 "C with a feed having O&H4 = 0.5, the minimum steam/C& ratio is 0.5. However, it has a negative effect overall, primarily due to increased water gas shift reaction, with selectivity to CO, quality of syngas, i.e., syngas stoichiometric number, and purity of raw methanol, all suffering as the steam-to-carbon ratio is increased. This results in greater gas compression and methanol refining loads and, as a result, more energy required per ton of methanol produced. In addition, an excessive steam-to-carbon ratio results in the COS concentration of the syngas increasing beyond the limits of the industrial methanol synthesis catalysts. Literature data (Ladebeck, 1991; Lee, 1990; Supp, 1989) suggests that the conventional CdZnO catalyst gives a reduced methanol yield when the C 0 2 concentration in the syngas exceeds 12 vol %. Hence, in practice, the steam-to-carbon ratio should not exceed more than the amount necessary to prevent carbon deposition. Comparison of CPO with Alternative Reforming Technologies in Methanol Synthesis. Comparison of the various syngas generation technologies in isolation has limited meaning, since downstream processing

can interact considerably with the syngas production "front end" of the plant. This is supported by the results obtained so far for a 1825tondday methanol plant using the CPO route for syngas generation, which indicate that the overall energy demand per ton of methanol produced is significantly affected by partial oxidation reactor pressure, oxygen-to-carbonratio, and steam-tocarbon ratio. As a result we have compared the various syngas technologies within the context of methanol production using conventional methanol synthesis technology. Table 4 lists data comparing the performance data of 1825 todday methanol plants based on the CPO, combined reforming (COMB. REF.), ICI's gas heated reforming, and Exxon's all-in-one fluidized bed reforming processes. The process data for these processes were retrieved from the relevant references and used as such for ASPEN simulation of the respective methanol plants. The results clearly indicate that the net energy demand for the overall methanol synthesis process, when using the CPO route as the syngas production step, is essentially on par with the competing technologies. However, oxygen demand per ton of methanol produced is highest for the CPO process. If we combine reactions 2 and 3, the stoichiometric oxygen requirement per ton of methanol produced could be calculated as 0.5 ton for an ideal methanol synthesis process where oxygen reforming is a part of the overall process. Thus, if oxygen consumption per ton of methanol production is used as a criteria, the conventional combined reforming and ICI's gas-heated reforming technologies seem to be the most attractive options for natural gas utilization. In addition, since the latter process offers reduced overall size of the syngas generation unit (Foulds and Lapszewicz, 1994), one may select the IC1 gas heated reforming type process configurationfor remote gas field and offshore application. Of course, the final decision should involve capital and operating cost benefits, ease of operation, safety, and overall reliability of the process. Finally, it should be noted that the reactor exit temperature for completely reformed gas in the case of the CPO process is essentially the same as that encountered in the other competing technologies. This implies that mechanical constraints due to pressure and thermal conditions, in the CPO process could be similar to those encountered using the competitive syngas technologies.

Conclusions ASPEN simulations of a methanol plant, based on the catalytic partial oxidation of natural gas, and operating under various process scenarios, clearly indicate that the overall energy demand per ton of methanol produced is significantly affected by partial oxidation reactor pressure, oxygen-to-carbon ratio, and steam-to-carbon ratio. It has been shown that CPO rates poorly against competing technologies when compared on the basis of identical steam-to-carbon and oxygen-to-carbon ratios but rates comparably on the basis of equivalent conversion. Ideally, the CPO reactor should operate a t low pressure for greater gas conversion and syngas quality. However, since the downstream use of syngas usually requires high pressure, as is the case in conventional methanol synthesis, the differential benefits of greater conversion and syngas quality are rapidly lost to high gas compression cost. The loss in conversion due to

1042 Ind. Eng. Chem. Res., Vol. 34, No. 4, 1995 Table 3. Effect of Steam Addition at Constant Oxygen-to-CarbonRatio 3200 3200 CPO reactor pressure P a ) 500 948 0.465 0.5 81.29 93.21 1.732 2.268 5.26 0.433 8100 255 2.81 92.9 16.83 24.64 0.671 0.758 0.178 40.49

CPO reactor inlet temp ("C) CPO reactor exit temp ("C) oxygen-to-carbon ratio steam-to-carbon ratio conversion of natural gas (8 by mol) selectivity to CO (% by mol) syngas stoichiometric no. HdCO molar ratio of syngas COz molar % in syngas (dry basis) CPO compressor load ( M W ) methanol synthesis pressure ( P a ) methanol synthesis temp ("C) recycle ratio purity of raw methanol (% by wt) methanol plant compressor load (MW) oxygen plant power load (kW) oxygen per unit of 99.5% methanol natural gas per unit of 99.5% methanol demineralized water per unit of 99.5% methanol net energy required (GJD of methanol)

Table 4. Comparison of Alternative Reforming Technologies on Equal Conversion Basis" COMB. CPO

REF.

reactor pressure (kPa)

3200

reactor inlet temp ("C)

500'

reactor exit temp ("C)

983

41OOb 390Oe 500' 58lC 750b 974c 0.465 3.0b 95.34 74.68 1.904 3.268 9.72 0.878 8000 255 2.26 84.0 10.65 22.03 0.537 0.607 0.467 35.88

reforming technologies

oxygen-to-carbonratio 0.600 steam-to-carbonratio 2.0 natural gas conversion (% by mol) 98.14 selectivity to CO (% by mol) 70.58 1.662 syngas stoichiometric no. (SN) HdCO molar ratio of syngas 3.129 11.66 COz molar % in syngas (dry basis) natural gas compressor load (MW) 0.364 methanol synthesis pressure (kPa) 8000 methanol synthesis temp ("C) 255 recycle ratio 1.91 punty of raw methanol (% by wt) 83.6 methanol plant compressor load (MW) 11.6 oxygen plant power load (kW) 26.71 oxygen per unit of 99.5% methanol 0.716 natural gadunit of 99.5% methanol 0.635 process waterhnit of 99.5%methanol 0.344 net energy required (GJfl' methanol) 35.93

IC1 Exxon 4100* 3900' 42fjC 64@ 640b 98Oe 0.480 3.0 97.94 65.50 1.901 3.852 9.72 0.849 8000 255 2.12 80.5 9.59 22.29 0.543 0.594 0.586 35.68

2500 500

982 0.520 0.5 91.54 94.37 1.745 2.206 4.82 8000 255 1.97 93.0 15.31 23.58 0.647 0.654 0.132 35.53

a Figures were mathematically derived to achieve 98% conversion. Primary reformer. Secondary reformer.

increasing pressure can be compensated for, to a certain degree, by increasing the reactor temperature. However, high temperatures, coupled with high pressure, often necessitate the use of expensive materials of construction. It has also been shown that steam requirements add considerably to the overall energy demand and, as a result, the steam-to-carbon ratio should be kept to a minimum. This will of course depend on the propensity of the catalyst to form carbon deposits. In terms of catalyst development, catalysts thus developed for operation at pressure should not be prone to excessive carbon deposition, so that a low steam-to-carbon ratio can be used. Although the net energy requirement for a methanol plant based on the CPO process for syngas generation is essentially on par with the competing technologies, the oxygen demand per ton of methanol produced is highest. It is approximately 45% higher than the stoichiometric minimum requirement of 0.5 and close to 35%higher than that exhibited by the IC1 gas heated

500 902 0.465 1.0 81.29 80.51 1.733 2.784 8.60 0.442 8100 255 2.93 87.8 17.88 25.14 0.685 0.774 0.327 41.33

3200 500 869 0.465 1.5 81.06 69.74 1.730 3.365 11.26 0.456 8100 255 3.04 83.5 19.06 25.93 0.703 0.793 0.475 42.37

3200 500 842 0.465 2.0 80.75 60.64 1.726 4.014 13.38 0.457 8100 255 3.19 79.9 19.60 26.03 0.710 0.801 0.597 42.79

3200 500 819 0.465 2.5 80.43 52.96 1.722 4.734 15.10 0.457 8100 255 3.56 77.0 19.90 26.03 0.710 0.801 0.690 42.80

reformer technology. This necessitates the requirement of a larger air separation unit for the CPO technology, a disadvantage for offshore utilization. Indications are that the oxygen source will contribute substantially to the overall cost of the process. This, together with safety considerations, suggests that alternatives to cryogenic separation should be considered. The simulations utilize an adiabatic equilibrium reactor into which oxygen is injected separately, since introducing a premixed feed with the desired Oz/CH4 ratio of 0.5 a t pressure is fraught with danger, particularly if no or little steam cofeed is used, as it is well within the explosion limit at the reaction temperatures being considered. Very rapid transfer of the gas mixture to the catalytic zone is required if the possibility of flashback and the occurrence of gas-phase reactions are going to be minimized. In addition, the occurrence of hot spots in the reactor should be avoided for the same reasons. Fluidized bed and recirculating fluidized bed reactors emerge as the most promising in this regard. A fluidized bed system not only facilitates good mixing of the reactant gases but, with a suitable catalyst, also could facilitate operation with a steam to carbon ratio well below 2 and still minimize carbon deposition due to Boudouard reactions. The simulation results confirm that Exxon's fluidized bed process (Goetsch and Say, 1990) is on par with the alternative syngas technolgies, in terms of net energy requirement. Nevertheless, it appears that the reactor design and construction are areas which would probably contribute considerably to the capital cost of such a plant. To conclude, on the basis of net energy requirement, CPO is on par with the alternative syngas technologies. Notwithstanding, a final comparison should involve capital and operating costs, as well as reliability and ease of operation. However, it should be noted that a comparison of this nature would be difficult since the major differences in capital and operating costs would be associated with the reactor, associated oxygen plant, and heat recovery system. At present this is proprietry for existing systems such as ICI's gas-heated reformer and Exxon's fluidized bed reactor, while detailed design would be needed for any new CPO reactor if a meaningful comparison was to be made. It is worth noting that if the CPO reactor design can incorporate the high space velocities recently demonstrated in the open literature for laboratory trials, then a much more compact reactor

Ind. Eng. Chem. Res., Vol. 34, No. 4, 1995 1043 should be able to perform the same duty, thereby giving it the edge in productivity.

Acknowledgment The authors wish to thank J. Lapszewicz, V. N. Ravavarapu, P. Fleming, and P. Jackson, Energy Research & Development Corp., and B.H.P. Co. Ltd. for technical and financial assistance.

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Received for review November 10, 1994 Accepted December 2, 1994" IE940028H Abstract published in Advance ACS Abstracts, February 15, 1995. @