Methane Oxidative Coupling: Technical and Economic Evaluation of a

Jan 16, 1995 - The economics are also compared to those of a conventional oxidative coupling chemical plant. The return on investment after taxes (ROI...
1 downloads 0 Views 1MB Size
Download PDF
Energy & Fuels 1996,9,794-801

794

Methane Oxidative Coupling: Technical and Economic Evaluation of a Chemical Cogenerative Fuel Cell Douglas Eng and Po-Hung Chiang Chemical Engineering Department, Tufts University, Medford, Massachusetts 02155

Michael Stoukides* Chemical Engineering Department and Chemical Process Engineering Research Institute, University Box 453, Aristotle University of Thessaloniki, 54006, Greece Received January 16, 1 9 9 9

A technical and economic evaluation of a chemical cogenerative solid oxide fuel cell plant that uses methane (natural gas) as he1 is presented. The 100 MW electrical power plant produces simultaneously 0.21 x lo9lb of ethylene from the methane oxidative coupling reaction that takes place at the anodic electrode-catalyst of yttria-stabilized zirconia solid electrolyte cells. The technoeconomic evaluation gives an acceptable return on investment before taxes of about 21.5%. The economics are also compared to those of a conventional oxidative coupling chemical plant. The return on investment after taxes (ROUT) is 13% for the fuel cell plant while for the equivalent conventional chemical plant it is about 14%. The most important contributors to the total cost are identified. Areas of possible cost reduction that could significantly increase the competitiveness of the fuel cell plant are also identified and discussed.

Introduction In the past 10 years, numerous reports and publications have focused on a very challenging problem, the direct conversion of methane to Cp hydrocarbons, preferably to ethylene.1,2 What makes this reaction system very interesting is the fact that the C2 yield (the overall methane conversion times the C2 selectivity) is limited with the best catalysts giving yields of the order of 25%.2 Recently, several innovative design approaches have been p r ~ p o s e din ~ , which ~ yields exceeding 85% have been achieved. Nevertheless, the problem of low per pass yields is still there. And, although the C2 yield is not the only criterion for the feasibility of a scaled-up p r o ~ e s sit , ~is generally understood as the main reason for the lack of competitiveness of this direct route of methane activation. As with most catalytic oxidations, it was realized that a key in achieving colhmercially acceptable yields of products of partial oxidation is the type and state of oxidant.1,2,6 In the oxidative coupling reaction, for example, researchers tested various alternatives including sequential feeding of oxygen and methane in the reactor or using the lattice oxygen of a metal oxide catalyst in which the metal can easily switch between two oxidation states.2 Along these lines, the use of solid

* To whom correspondence should be addressed.

Abstract published in Advance ACS Abstracts, July 15,1995. (1)Wolf, E. E. Methane Conversion by Oxidative Processes; Van Nostrand Reinhold New York, 1992. (2) Amenomiya, Y.; Birss, V. I.; Goledzinowski, M.; Galuszka, J.; Sanger, A. Cut. Rev.-Sci. Eng. 1990,32 (3), 163. (3)Tonkovich, A. L.;Carr, R. W.; A r i s , R. Science 1993,262,221. (4) Jiang, Y.; Yentekakis, I. V.; Vayenas, C. G. Science 1994,264, 1563. ( 5 ) Matherne, J. L.; Culp, G. L. In Methane Conversion by Oxidative Processes; Wolf, E. E., Ed.; Van Nostrand Reinhold: New York, 1992; p 463. (6)Rtchai, R.;Klier, K. Cutul. Rev.-Sci. Eng. 1986,28,13. @

0887-0624/95/2509-0794$09.00/0

feed

'

air

effluent

t

Figure 1. Schematic diagram of a methane solid electrolyte cell-reactor.

electrolyte cells in which oxygen is electrochemically has also been examined and results supplied as 02-, have been recently r e v i e ~ e d . ~ In addition to the different type of oxygen used, there are further advantages for carrying out a catalytic oxidation in such a reactor cell. Figure 1is a schematic diagram of a methane solid electrolyte cell. In this electrochemical reactor, the yttria-stabilized zirconia CySZ, the most commonly used solid state 02conductor) wall serves not only as an electrolyte but also as a separator of nitrogen from oxygen of the air. Thus, the cell is a natural separator of ambient air's two principal components. Also, the concentration profile is the inverse of what you typically get in the conventional reactor. Ionic oxygen is consumed on the electrode surface and unreacted oxygen diffises to the gas phase and forms a profile with decreasing 0 2 concentration from the surface to the bulk. In practice, with a highly active electrode, the gas-phase concentration of oxygen remains insignificant. This is important in cases where (7) Eng, D.; Stoukides, M. Cutul. Rev.-Sci. Eng. 1991,33,375.

0 1995 American Chemical Society

Methane Oxidative Coupling undesirable gas-phase reactions should be avoided or when the oxygen to methane ratio must be kept outside the explosion limits. Finally, the solid electrolyte cell can operate (a) as a conventional fuel cell in which reaction products (usually H20 and/or C o d need not be collected but only electrical energy is generated, (b) as an electrolyzer when electrical energy is consumed to produce the desired chemicals, and (c) as a chemical cogenerator, in which valuable chemicals and electrical energy are produced simultaneously.8 Considering all these advantages of solid electrolyte aided processes, a reasonable question is why the chemical industry has not yet adopted these systems on a large scale. Just the idea of simultaneous production of electricity and valuable chemicals alone looks quite attractive and commercially profitable. This question has been discussed and answered in various ways in previous communications.~-~~ Spillman, Spotnitz, and Lundquistg presented a model by which one could decide which of the three modes of operation of a solid electrolyte cell, fuel cell (producingelectrical power only), electrolytic cell (producing chemicals at the expense of electricity), or cogenerative cell (producing both electricity and chemicals), is economically preferable. Their analysis showed that in order for chemical cogeneration t o be the preferable choise, the difference in values of reactants and products should be low, the reaction AG should be high, and the molecular weight of the product should be Vayenas, Bebelis, and Kyriazis developed a simple model in which capital costs were also included and tested for the economic feasibility of various candidate cogeneration processes.loJ1 Results were promising for HzS04 and possibly for HNo3 production while the predictions were discouraging for formaldehyde and for ethylene oxide.1° The above authors found that good candidate reactions are those that are exothermic, with inexpensive raw materials and products, and are carried out at temperatures of 800 "C or higher. In general, their conclusions were in agreement with those of Spillman et al. with the exception of the addition of one very promising parameter: the possibility of non-Faradaic modification of catalytic activity (NEMCA) via which the economic analysis could be altered.ll This is because if a strong NEMCA effect appears, one could get a huge increase in the reaction rate by consuming a very small amount of electrical power.11J2 The reaction of methane oxidative coupling seems t o be a good candidate for chemical cogeneration since it fulfills the requirements set by the two previously mentioned research groups, i.e. (a) reactants and products are quite cheap, (b) the reaction is highly exothermic, and (c) the optimum operating temperature is about 800-850 0C.1,2 Vayenas et al. suggested that a more detailed economic analysis is needed in case a reaction satisfies the promising characteristics of their model.1° This is the topic of the present communication (8) Stoukides, M. Ind. Eng. Chem. Res. 1988,27, 1745 . (9)Spillman,R. W.; Spotnitz, R. M.; Lundquist, J. T., Jr. CHEMTECH 1984, 14,176. (10)Vayenas, C. G.; Bebelis, S. I.; Kyriazis, C. C. CHEMTECH 1991, 21, 422. (11)Vayenas, C. G.; Bebelis, S. I.; Kyriazis, C. C. CHEMTECH 1991, 21, 500.

(12)Vayenas, C. G.; Bebelis, S. I.; Yentekakis, I. V.; Lintz, H.-G. Catal. Today 1992, 11, 303.

Energy & Fuels, Vol. 9,No. 5, 1995 795 in which the technoeconomic aspects of a methane-toethylene cogenerative fuel cell are discussed.

The Ethylene Cogenerative Fuel Cell Plant To test the economic feasibility of a new chemical process, a reasonable approach would be to identify first the differences between an existing process and the one under consideration, then see how big is the economic impact of each of these differences, and finally sum them up to determine which of the two processes is financially favorable. It is not possible, however, to follow exactly this procedure in the case of methane oxidative coupling because there is no such industrial process in operation today. This has both positive and negative aspects: negative, because there is no basis of comparison with a real process, and positive, because in case the analysis provides promising results, it is easier for industry to adopt a totally new process than to replace one that is already in operation. The basic economics of the methane oxidative coupling plant were studied by Matherne and C ~ l p The .~ economic impact and sensitivity of various alternatives (cofeed vs sequential feed, high conversion vs high selectivity, effect of diluent) were examined. Their analysis also included a comparison between gasoline production from methane via the oxidative coupling or via the syngas-methanol route.5 The Matherne-Culp analysis served as a basis for the present study as well. The feasibility of cogenerating electricity and useful chemicals has been examined in the laboratory scale for several reaction systems and reviews of these works can be found in the literature.8J2 To upscale the cogenerative solid oxide fuel cell (SOFC),the market and location must be assessed. Presently, expectations for fuel cells call for medium-sized plants (200 k W to 20 MW) in residential and commercial areas. For the noncogenerative SOFC, this predicted market is highly realistic since SOFCs are environmentally safer than existing power plant designs. The small commercial SOFC can be safely sited in a well-populated area. The plants would burn methane cleanly and can be configured within compact areas.13 A chemical cogenerative SOFC plant with its auxiliary structures, however, faces strict limitations in wellpopulated sites since public ordinances and interests may restrict chemical production. Presently, the greatest awareness of SOFCs remains in Europe and in the United States, where public and environmental commitments are strong. In a remote region or industrial zone, a chemical cogenerative SOFC plant appears to be more feasible. In addition, the plant would be closer to other chemical facilities which may require ethylene or synthesis gas feedstock. In determining the possible location of the cogenerative plant, there are considerations implied for sizing: (a) a remote or industrial zone makes it possible t o upscale the plant to the medium to large level rather than the small plant size and b) basic economics implies that larger upscaling lowers the unit product price. The size arbitrarily chosen in the present study was a medium-large plant of 100 M W . Hence, in the technoeconomic evaluation which follows, an ethylene-cogen(13)Eng, D.; Chiang, P. H.; Stoukides, M. Technical Report to National Renewable Energy Laboratory of DOE; Contract XAR-313237-01-107811, February 1994.

Eng et al.

796 Energy & Fuels, Vol. 9, No. 5, 1995 N a n d Gar Feed

N n W gas treat"

cl

I

Air circulation

Quenching

Rimary SOFC

Exchanger

v Recycle Carbon dioxide

Waler

Methane gas

Carbon dioxide

recovery

treatment

Compression

U I

I c-3

Air Circulation

L4

I A

v

t

Products

Figure 3. Cross-flow multi-tubular design of the YSZ cells. Air flows inside and reaction takes place outside the cylindrical

tubes.

Heavy hydrocarbons

(a+)

C2H,

C=C (ethylene)

Figure 2. Basic design of the ethylene cogenerative fuel cell

plant. erative plant with an electricity power output of 100 MW was considered. The basic design for the ethylene-producing fuel cell plant is shown in Figure 2. Only the major equipment units of the plant are shown schematically. The process uses natural gas (or methane) as fuel to generate electricity by oxidizing methane to ethylene and water; ethylene is subsequently separated from the product stream. Equipment necessary for ethylene separation and methane recovery are also included in the design. Connections for processing units and principle processing material (excluding the SOFC itself) are stainless steel t o resist corrosion. A natural gas treatment unit to remove impurities from the feed is necessary if natural gas from the producing field is directly used as feed. Methane content in natural gas varies from 50 to 98%. Treatment is most probably desired if the feed is to enter with at least 97-98% purity. The remaining 2-3% is mostly ethane, carbon dioxide, and nitrogen gas. The treatment unit consists of a low-temperature compression tower for higher hydrocarbon removal and a standard molecular sieve adsorption unit for sulfur compound removal. The primary reactor consists of YSZ cells. Some possible designs are multiple-stacked, planar, monolithic, tubular, and mu1titub~lar.l~ Figure 3 shows the chosen cross-flow multitubular design which has been previously modeled.15J6 It is possible, however, that (14)Singhal, S., Iwahara, H., Eds. Proceedings ofthe 3rd Znternational Symposium on Solid Oxide Fuel Cells, The Electrochemical Society: Pennington, NJ, 1993. (15)McKenna, E. A.; Stoukides,M. Chem. Eng. Sci. 1992,47, 2951. (16) McKenna, E. A.; Othoneos, A.; Kiratzis, N.; Stoukides, M. Znd. Eng. Chem. Res. 1993,32, 1904.

planar multistack cells may be the least expensive to manufacture for a plant design. Whichever design may prove to be most economical will not make a significant difference in the present analysis since a fixed but high cost for electrolyte material and processing was assumed.13 Thus, the reactor considered in the present analysis consists of YSZ tubes with electrodes connected in series in thermally insulated housing. The dimensions for the YSZ tubes are 1.0 cm o.d., 1 mm wall thickness, and 200 cm in length, arbitrary sizing near those currently used in pilot plants. Electrodes consist of electronically conductive perovskite-type oxide, most likely strontia-ceria based. These promising materials have been shown to exhibit excellent coupling activity and can be made either ionically or electronically conductive depending on the degree of doping.'J7 As the anodic electrode, the catalyst must be electronically conductive. Each YSZ tube has approximately 625 cm2 of electrode area. There are approximately 2.15 x lo5 tubes,13 placed in series in modular units of ca. 1000. Each module will, therefore, have a total electrode surface area of 625 000 cm2. There are 215 of these units. Currently, in SOFC pilot plant up to 25 kW, cells have been constructed with in excess of 500 tubes per module.ls Certainly, there is a risk in scaling up from 25 kW to 100 MW and in the calculations that follow, it is assumed that fuel cells of that size will be constructed in the future. The cross-flowmultitubular design was chosen somewhat arbitrarily since most advancements regarding scaling up to the kilowatt level have used multitubular designs. It may well be, however, that recent advances in the planar stack and other nontubular designs offer brighter future prospects than the cross-flow multitu~

(17)Stoukides, M. J.Appl. Electrochem., in press. (18)Singhal, S. In Proceedings ofthe 3rd International Symposium on Solid Oxide Fuel Cells; Singhal, S., Iwahara, H., Eds.; The Electrochemical Society: Pennington, NJ, 1993; p 665.

Methane Oxidative Coupling

Energy & Fuels, Vol. 9, No. 5, 1995 797

Table 1. Material and Energy Balances for the Cs Cogenerative Fuel Cell Plant chemical

millions of lb/year

methane feedstock ethylene production carbon dioxide carbon monoxide water

475.0 212.9 523.0 83.2 762.0

Assumed Reactions and Selectivities CH4 202 COz 2H20 Scol = 40% CH4 3/202 CO H2O SCO = 10% 2CH4 V 2 0 2 C2H6 H2O S c 2 ~=@15% 2CH4 + 0 2 C2H4 2Hz0 S C ~=H 35% ~

+

+ +

--

4

+

+

+

+

overall methane conversion on a single pass (without recycle) = 50% Energy Balance waste heat produced per year steam (316 "C, 3 atm) generated per year

through the SOFC reactor.13 The whole plant operates autothermally and produces 100 MW of electricity. The conversion of 50% with 35% selectivity t o ethylene was found to be sustainable in a auto-operated SOFC with an assumed 60% efficiency he., free energy converted t o electricity).13 In addition, it was checked that excess heat did not accumulate within the SOFC to interfer with autothermal operation a t 850 "C. The stoichiometric reaction can be written as CH,

0.175C,H4 -l-1.425HZO (1)

For each tube, the oxygen flux through the electrolyte wall t o the catalyst is

5.41 x 1014cal 7.56 x lo8 kg

bular design. In the latter, the anode-catalyst is deposited on the outside peripheral area of the tubes and reactants flow between the outer walls of the tubes. Inside the tubes is the cathode in contact with air which is slightly compressed into the tubes by centrifugal or rotary blowers. This design was chosen for two principal reasons: (1)The anode on the outside wall allows a greater catalyst surface area. On the inside wall, which has a smaller surface area, is the cathode which has a much higher conductivity due t o recent advances in technology14and which would not limit the ionic flux. (2) If the coupling reaction takes place inside the 200 cm long tubes, mixing may not be proper. In a crossflow reactor, mixing is more easily facilitated by the tube matrix. The tubes should be fairly close packed since the contribution of homogeneous gas-phase reactions must be minimized. Natural gas and oxygen are separately fed into the reactor. The mean reactor temperature is 850 "C a t atmospheric pressure. During the reaction, methane is partially oxidized to form ethane, ethylene, water, and carbon oxides. Electricity is generated by ionic oxygen diffusing through the solid electrolyte membrane. The cell voltage for a single electrolyte tube is assumed t o be 0.5 V with current density of 1.5 Ncm2 based on the anode-catalyst surface area. Current density is based on the anode rather than the cathode since the former is the limiting factor for conductivity. Currently, SOFC pilot plants utilize cell voltages from 0.4 t o 0.8 at 900 "C and 0.6 to 0.9 a t 1000 "C.19 At 850 "C, it can be expected that 0.5 V or slightly higher is reasonable. Methane conversion is assumed to be 50% per pass. Since at these elevated temperatures the non-Faradaic enhancement is negligible,7 the effect of NEMCA on conversion and selectivity was not considered in the calculations. Product selectivity for carbon is assumed as follows: COZ,40%; CO, 10%; C2H6,15%; CZH4,35%. The mass balance for the plant based on these values is shown in Table 1. These values are realistically achievable numbers based on the reported result^^,^ and make the overall reaction exothermic enough to maintain the reactor at the given temperature. The heat of reaction was checked for several different product compositions and temperatures based on a single pass (19)Bossel, U. G. In Proceedings of the 3rd International Symposium on Solid Oxide Fuel Cells; Singhal, S., Iwahara, H., Eds.; The Electrochemical Society: Pennington, NJ, 1993; p 833.

+ 1.16250, - 0.4c0, + 0.1Co + 0.075C& +

= i/4F = (1.5)(625>/(4)(96487> =

2.43 x

mol/(s.tube) (2)

The required methane feed and product yields may be calculated based on eq 1. vCH4= (vO2)(2>/(2.325> = 2.09 x

mol/(s.tube) (3)

(7)

Based on the above calculations, the amount of individual chemical production per year is summarized in Table 1. Annual CZproduction is 2.13 x lo8 lbs. The ethane produced from the primary reactor is fed t o the secondary reactor to be converted to ethylene. The secondary reactor consists of strontia-ceria-ytterbia (SCY) tubes, having the same dimensions as the YSZ tubes in the primary reactor. Operating conditions are 850 "C and atmospheric pressure. The SCY is a hightemperature proton-conducting solid electrolyte and, similar to the YSZ cells, it is used here as a wall of the ethane dehydrogenation fuel cell reactor. Thus, electricity is produced by oxidizing hydrogen which is transported as H+ through the solid e l e c t r ~ l y t e . ~ ~ ~ ~ , ~ ~ Electrode area per tube is also assumed to be 625 cm2 and current density is again assumed t o be 1.5 Ncm2 at 0.5 V. The conversion of ethane is assumed to be 50%, and 50% of Hz produced inside the reactor permeates through electrolyte wall to react with oxygen in the air at the cathodic electrode to form water. The reaction is

The total current generated by this reactor is (20) Iwahara, H.; Esaka, T.; Uchida, H.; Maeda, N. Solid State Ionics 1981,3/4,359. (21)Hamakawa, S.; Hibino, T.; Iwahara, H. J. Electrochem. SOC. 1993, 140, 459.

Eng et al.

798 Energy & Fuels, Vol. 9, No. 5, 1995

i = NH22F= 6.45

x

lo6 A

(10)

The number of tubes required to convert ethane to ethylene is calculated by

n = i/(As)(j) = 6882 tubes

(11)

where A, is the electrode area and j is the current density. Therefore, the power generated by the secondary fuel cell is 3.2 MW. The contribution of the secondary reactor is negligible in terms of electrical power generated. It is thus possible that the SCY cells be replaced by a conventional dehydrogenative refractory-lined steel reactor which further oxidizes ethane to ethylene. The SCY secondary reactor is largely hypothetical compared to the YSZ primary reactor at this time simply because of the novel technology involving proton conductors. Whether proton conductors prove to be versatile or not is a test of time. In addition, currently, proton conductors are relatively expensive to construct because of the more costly raw materials. In this study, the secondary reactor is relatively expensive and is not as economical as the primary reactor in returning investment. The poor investment factors regarding the SCY cells make it quite probable t o drop this unit from the plant design and replace it with a conventional reactor. The exit product gas stream at 850 "C is cooled to 537 "C in the quenching unit to produce high-temperature supersaturated steam at 316 "C and 3 atm pressure. By energy balance calculation, 705 million kg of steam can be generated each year.13 The product gas stream is cooled further to 75 "C in a counter flow water-cooled heat exchanger to preheat methane and air feed to 450 "C. The cooled product stream is compressed in the compression system to remove most of the water. The compression unit is similar to those described in the l i t e r a t ~ r e . ~The , ~ ~compression -~~ unit consists of centrifugal motor compressors with an incremental fourstage pressure process from 2 to 20 atm. The carbon dioxide or acid gas treatment system is an amine-scrubbing system that removes carbon dioxide and the remaining water from the compressed product stream. The product stream is subsequently dried in a standard fixed particle bed.5,24 The methane gas recovery system consists of a series of refrigerations which condense the heavier hydrocarbons (Cp and higher if formed). The minimal temperature of this refrigeration series is below 169 K, that of the boiling point of ethylene. In these stages, methane is separated as a gas from liquified ethane, ethylene, and higher hydrocarbons. The recovered methane gas is recycled back to the reactor feed stream. The heavy stream, containing non-methane hydrocarbons, continues to the distillation towers. The methane and higher hydrocarbon streams are heated back to near ambient temperature by heat exchangers with the gas entering the refrigeration cycle.

The heavy stream from the methane recovery system enters the distillation tower where heavy hydrocarbons are separated from the Cp hydrocarbons. Heavy hydrocarbons are separated from CZ'S at the bottom of the first tower. The Cp stream from the top of first column enters a second tower for the separation of ethane from ethylene. The purity of ethylene can be expected to reach near 99.8% depending on the number of distillation plates. The ethane is then fed to the secondary fuel cell reactor (or a conventional reactor) where it is oxidized to ethylene.

Cost Calculations

b.y Oxidative Processes; Wolf, E. E., Ed.; Van Nostrand Reinhold: New York, 1992; p 429.

Current state-of-the-art SOFC pilot plants run 3000 h and produce up to 25 kW. Presently, mature market costs are estimated as $1000 to $1200 per k W for a pure SOFC plant.25 It has been shown in current tests that SOFCs can be operated continuously for 25 000 h.25The estimated continuous operation of our plant is 56 000 h (or 6.7 years) before some major maintenance would be required. The auxiliary plant facilities would operate considerably longer. Both operational spans are incorporated into the depreciation rates for the economic evaluation. Laboratory studies of methane oxidative coupling have shown that, in addition to the problem in maintaining high CZ yields, it is considerably difficult to maintain the catalyst activity for a long time. Since sufficient data on catalyst lifetime and required frequency of regeneration are not widely available, it was decided to assume a lifetime of 6.7 years. This is an arbitrary number which, however, does affect the cost calculations. Nevertheless, for the purpose of comparing a catalytic process with the present chemical cogenerative approach,as it is done below (Table 31, it is fair enough at this stage to consider equal lifetimes for the catalyst and the electrode respectively. Operation of the plant is year round (ca. 350 days, 24 h per day) and can vary in capacity depending on seasonal requirements. If there is increasing demand for Cp's, then the plant should be flexible enough to allow greater production of chemicals at the probable expense of electricity. Normal operations are assumed at 80% of maximum Cp processing production. If electricity is in heavy demand, the plant can be modified as a fuel cell only. Normally, the plant is operated as a cogenerative system since spontaneous generation of electricity is carried by 02-which oxidizes methane. Table 2 summarizes the economics for the proposed plant. The major investment cost for the plant is that for the primary SOFC reactor. The price of YSZ tubes is based on (1)an estimated large order price assumed to be 1' 3 of the current quoted YSZ tube price by Zircoa Products13and (2) the cost of raw materials of ceramicgrade (99+%) zirconium oxide and yttrium oxide for large quantity orders for several chemical companies. Electrode price is based on perovskite-type oxide price as quoted by Seattle Specialty Ceramics.13 Density of electrode material is taken to be the near that of SCY powder, which is a perovskite-type oxide. The amount of electrode material required is calculated by assuming that electrodes have an average thickness of 100 pm.

(23) Kuo, J. C. W. In Methane Conversion by Oxidative Processes; Wolf, E. E., Ed.; Van Nostrand Reinhold: New York, 1992; p 483. (24) Peters, M. S.; Timmerhaus, K. D. Plant Design and Economics for Chemical Engineers, 3rd ed.; McGraw-Hill: San Francisco, 1980.

(25) Hooie, D. T. In Proceedings of the 3rd International Symposium on Solid Oxide Fuel Cells; Singhal, S., Iwahara, H., Eds.; The Electrochemical Society: Pennington, NJ, 1993; p 3.

(22) Edwards, J. H.; Do, K. T.; Tyler, R. J. In Methane Conversion

Methane Oxidative Coupling

Energy & Fuels, Vol. 9,No. 5, 1995 799

Table 2. Economic Summary for CZCogenerative Fuel Cell Plant cost ($ millions)

items YSZ Reactor reactor (YSZ tubes @ $400/tube) electrodes, @ $l.O/g housing and insulation, 5% of YSZ electrodes supporting equipment (based on production capacity) SCY Reactor reactor (SCY tubes @ 150% of YSZ tubes) electrodes, @ $l.O/g housing and insulation supporting equipment

+

compression acid gas removal cold box refrigeration refining instrumentation electrical equipment total fixed investment working capital (15% of fmed capital) total capital investment (TCI) Operation Cost natural gas (@ $0.06/lb) operators (8 person shifts, 3 shifts) shift supervisor (1person per shift) overall supervisors (2 persons) labor overhead, 35% of direct labor total labor maintenance, 6% of fmed capital plant overhead, 4.5% of fxed capital depreciation, 12% of fmed capital total fixed cost

4.13 0.01 0.21 1.70 25.78 5.98 19.21 19.21 13.24 2.58 19.68 251.59 37.74 289.33 32.31 0.84 0.15 0.11 0.39 1.49 15.10 11.32 30.19 58.09 90.41

total production cost ethylene credit, @ $0.21/lb steam credit, @ $30/ton

(44.73) (21.13)

operation cost for electricity income from electricity @ $O.lOkWh income before taxes (IBT) rate of return on investment before tax (ROIBTP (%) a

85.33 0.26 4.28 50.0

24.55 86.69 62.14 21.5%

ROIBT = IBTPTCI.

Table 3. Comparison of Cogenerative Fuel Cell Plant and Conventional CZPlant" cost/$ millions items

case 1

case 2

total fmed investment working capital total utilized investment (TUI) operation cost fixed cost total cash cost (TCC) income (I) depreciation (D) taxes (TX)

530 60 590 97 40 137 320 53 48 82 14

1181 78 1259 108.6 17.4 186 563 118.1 96 163 13

NIAT R O U T (%)

a Case 1: methane diluent case of ref 5. Case 2: ethylene cogenerative fuel cell. Working capital: 10% of sales plus 20% of [raw materials utilities costs]. Depreciation, D = 10% of total fmed investment. Taxes, TX = 0.37 [ I - TCC - Dl. Net income after tax, NIAT = Z - TCC - D - TX. Return on Investment after tax, ROUT = NIATPmTI.

+

Cost of housing and insulation for the electrolyte tubes is assumed to be 5%of total cost of tubes and electrodes. This value is roughly estimated and may deviate from the actual price; however, it will have a minimal effect on the final analysis since it is a relatively small figure. The supporting equipment, service, and contingency

for the reactorlgenerator, based on the production capacity of the reactor, is estimated to be $50 million, rather than the standard 200% of delivered equipment.24 Using the standard 200% of delivered equipment would unrealistically inflate the fixed investment since the YSZ reactor is considerably more expensive than conventional reactors for the same production output. Price of a SCY tube is assumed to be 150% of a YSZ tube, since SCY is not yet commercially available but should command a higher price than YSZ. The price of compression, carbon dioxide treatment, cold box, refrigeration, and refining are based on the price of the same equipment for a CPproduction plant with capacity of one billion pounds per yeara5 Since our plant produces only 213 million pounds annually, the prices were scaled according t o the equation24 cost of equipment A = cost of equipment B x (capacity of equipment Ncapacity of equipment B)o.6 The Marshall and Swift equipment cost indexP6was used to adjust the originally quoted 1989 price for the plant with a lo9 lb annual capacity (index = 895.1) to a mid-1993 price (index = 966.6). Included in the equipment cost are the installation, supporting equipment, building cost, etc. The cost for electrical installation and instrumentation of the power plant is based on price given for similar equipment for a 1000 M w oil-fired power plant.27 Although the present power plant has 1/10 of the capacity of this oil-fired plant, we estimated our plant's electrical installation and instrumentation to be 80% of the price for the 1000 M w plant before adjustment for inflation. The rationale was that the cost of electrical equipment is only slightly affected by the magnitude of power The working capital is assumed to be the standard 15%of fixed capital. The number of operators per shift is estimated from typical labor requirements for process equipment.24 Salary per operator is assumed at $35 000 per operator per year a t 1993 with 3 shifts per day. Salary for shift supervisor is assumed to be $50 000 per person per year. Salary for overall supervisor is assumed to be $55 000 per year. Labor overhead is assumed to be the standard 35% of total fixed labor cost. Plant maintenance is estimated as 6% of fxed capital. Plant overhead is 4.5% of fixed capital. Both these values are standardeP4 The equipment annual depreciation is estimated a t 12%,which is the average of 15%for reactods) and 10% for other equipment. The figure for reactor is based on the assumption that it has a depreciation lifespan of 6.7 years whereas the remaining equipment spans about 10 years. The credit for ethylene is $0.21 per pound.13 Electricity is assumed to be priced at $0.10 per kwh which is based on the current mean rates set by utility companies. A credit of $30/ton of steam is assumed for the cogenerated steam which is near the average market value.'3

Discussion The early research works on the oxidative coupling of methane considered the per-pass CP yield as a key ( 2 6 )Chemical Engineering; McGraw-Hill: New York, 1993; p 150. (27) Hicks, T. G. Power Plant Evaluation and Design Reference Guide; McGraw-Hill: San Francisco, 1986.

800 Energy & Fuels, Vol. 9, No. 5, 1995

parameter determining the economic feasibility of the process. In one of these early technical and economic evaluations conducted by Jackson, Thompson, and Whitehead,28 a minimum yield of 25% was set as a requirement. The above authors compared the production of gasoline via methane coupling with the conventional technology of methane to gasoline via syngas and methanol. Their results showed that with the existing data, the indirect route via methanol was favored. Nevertheless, several potential cost saving areas in the oxidative coupling plant were identified which could lead t o a significant improvement in the competitiveness of the coupling process.28 Extending the above work, a more detailed analysis was prepared by Edwards, Do, and Tyler.22 Their experimental results showed that a single fluidized bed process that combined ethane pyrolysis and methane coupling was technically feasible. Their economic evaluation identified two key factors that mostly affect the process cost, (a) the need to achieve higher methane conversion at a given high C2 selectivity and (b) the need t o operate at the highest permissible pressure.22 Kuo conducted an engineering evaluation study of several direct processes of methane activation and screened them down t o the two most promising: the partial oxidation to methanol and the oxidative coupling, which he studied in more There is a number of interesting findings in that work including that the oxidative coupling becomes competitive if a 35% conversion is achieved with a simultaneous 88% or higher selectivity, i.e., a 31%+ yield per pass. Matherne and Culp examined the sensitivity of the methane coupling economics t o reactor diluent and methane to oxygen feed ratio.5 Three possibilities of dilution in the cofeed mode were considered, i.e., with nitrogen, with steam, and with excess methane. These three cases were also compared to a fourth one in which no diluent was used and the feed of oxygen and methane was sequential. The nitrogen- and steam-dilution cases could give yields as high as 22% while the methanedilution and the no-diluent cases had a 14% and 15% C2 yield, re~pectively.~ The relatively lower yields in the latter two cases were due t o the restriction in the methane to oxygen ratio determined by the explosion limits. It is interesting to point out that the return on investment after taxes (ROIAT) was not exclusively dependent on the C2 yield. For example, the methanedilution case had a 14% ROIAT while for the nitrogendilution case the ROUT was only 3%. Clearly, the CZ yield alone is not a safe predictor of the economics of the p r o ~ e s s . Matherne ~ and Culp also compared the economics of two methane to gasoline routes, one via methane coupling and one via syngas and methanol. Results showed that the methanol route was only slightly favored. This small difference was considered encouraging for the coupling process in view of the fact that the syngas-methanol route has been developed for many years so that there is no room for big improvements in the cost as opposed to the coupling route which is still at its early stages of de~elopment.~ Table 2 shows that rate of return on investment before tax (ROIBT) is 21.5%. The value of the ROIBT (28) Jackson, P. J.; Thompson S. C.; Whitehead I. G. Proc. 7th Australasian Chem. Eng., Conf: CHEMECA 89, Gold Coast, Queensland 1989, 719.

Eng et al.

1

0.41

0

400

200

600

800

1000

YSZ cost ($ I tube)

0.0

0.1

0.3

0.2

Electricity price ($ I kwh)

0.0

0.1

0.2

0.3

0.4

Ethylene price

0.5

0.6

0.7

(1/ Ib)

Figure 4. (a, top) Dependence of ROIBT on the cost of the solid electrolyte material. (b, middle) Dependence of ROIBT on the sale price of electricity. (c, bottom) Dependence of ROIBT on the sale price of ethylene.

should be at least 25% if this design is to be attractive to industry. The minimum rate of return should be at least 20% for a sound i n ~ e s t m e n t .It~ was ~ found that the ROIBT is strongly dependent on the following: the cost of the fuel cell materials, the credit for ethylene, and the credit for e1e~tricity.l~ The credit for ethylene was taken at a price of $0.21 per pound but this value has fluctuated to over $0.30 per pound in recent years.

Energy & Fuels, Vol. 9, No. 5, 1995 801

Methane Oxidative Coupling

Also, the credit for electrical power was assumed to be $ O.lO/kWh and it is also expected to fluctuate quite a bit and probably toward lower values. Finally, it can be seen in Table 2 that the cost for the YSZ tubing is more than one-third of the total fmed investment and, therefore, the economics are clearly affected by the price of the solid electrolyte. Figure 4, a, b, and c, shows the dependence of ROIBT on the YSZ cost, the price of electricity, and the price of ethylene, respectively. Figure 4a shows that ROIBT decreases drastically with increasing the cost of the zirconia tubes. The calculated 21.5% is attained at an assumed cost of $400NSZ tube and at $600/tube it goes down t o 16.5% while for $200/tube it goes up t o 31%. In Figure 4b the electricity price is allowed t o vary from as low as 3 $/kwh to as high as 20 @/kWh.With YSZ cost and ethylene sale price maintained a t $400/tube and $0.21/lb, respectively, electricity must be sold a t a price of 16 g/kWh for ROIBT to exceed 30%. In Figure 4c, where ROBT is plotted vs the price for ethylene, one can see that in order to exceed the 30% ROIBT, ethylene must be sold a t $ 0.32Ab or higher. In general, the above results indicate that the economics for the chemical cogenerative fuel cell plant are encouraging but not very strongly. In order to make a more fair comparison, Table 3 has been prepared in which the basic cost and income calculations from the present work are compared to one of the four coupling processes of the Matherne-Culp a n a l y ~ i s .The ~ methane-diluent case was selected because it resembles the fuel cell design in which no diluent other then methane itself is fed in. In Table 3, case 1represents the above methane-diluent case and case 2 corresponds to the present study. The numbers for case 1were left exactly as r e p ~ r t e d and , ~ there were some necessary adjustments made in case 2. Both cases represent an annual production of lo9 lb of ethylene, sold a t a price of $0.32/ lb. No steam credit is given in either case but, in case 2, an additional credit for electricity ($241 millions) was given with an assumed electricity price of $0.06/kWh. The values for total fixed investment and fixed cost were increased from 251.59 and 16.59 (Table 2) t o 1,181and 77.4, respectively, i.e., multiplied by 110.213 which is the ratio of the capacity of the plant of Table 3 to that of Table 2. The operating cost is primarily the cost of methane and therefore should be essentially the same in both cases 1and 2. Nevertheless, it was assumed t o be 20% higher in case 2, just in order to be on the conservative side. Working capital, depreciation, taxes, and net income after tax (NIAT) were calculated the same way in both cases.5 Table 3 shows that the returns on investment after tax (ROIAT)are about the same, 14% and 13% for cases 1 and 2, respectively. The results of Table 3 indicate again that the ethylene cogenerative fuel cell does generate an acceptable return which, however, is not higher than those that could be generated by a well-designed conventional methane coupling plant. This looks somewhat surprising in view of the fact that the fuel cell plant provides an additional

income from electricity which is comparable (Table 3) to the income from ethylene. One partial explanation can be that, as previously mentioned, the fuel cell cost items were inflated to some extent in order to draw conclusions more safely. Also, several advantages of solid oxide fuel cells were poorly exploited as for example, the 0 2 - N ~separation. In the present design, air at the cathode is circulated because of the large oxygen fluxes required.13 Oxygen and nitrogen are separated but at the cost of losing heat to nitrogen which is heated up t o near reaction temperatures. Hence, some steam credit is lost and the advantage of free separation is underutilized. It should be pointed out, however, that even if the advantageous characteristics of these cells were utilized more properly, a reasonable but not outstanding increase in the return should be expected. There are additional reasons for this moderate success of chemical cogenerative fuel cells, as has already been Very briefly, these reasons include (a) immature testing of SOFC pilot plants, which makes the industrial sector reluctant to adopt them, (b) expensive SOFC technology, which will require further developments in materials science and ceramics technology in order to advance the electrolyte and electrode materials to a state where they are demonstrably practical, and (c) limited diversity in industry, which makes chemical companies t o need additional expertise before they enter two worlds, i.e., chemicals and electrical power generation.13 In summary, the present technoeconomic evaluation showed that, under certain conditions, the ethylene cogenerativeplant is a marginally favorable investment, comparable to a well-designed conventional coupling process. Since conventional processes of methane oxidative coupling have not been scaled-up either, it seems highly unlikely that the cogenerative ethylene fuel cell will be adopted by industry at the present time or in the very near future. It is encouraging, however, that there is a lot of room for economic improvement by either suppressing the basic costs of the process or by relative increase in the unit sale price of ethylene or of electrical energy. Finally, since the issues and policies of energy production and consumption are, among others, political processes in which public, corporate, and government groups are involved, it is reasonable to expect in the future state- or federal-funded incentives for industries to develop cogenerative fuel cells. Incentives of this kind could include tax breaks, subsidaries, and grants as well as support by environmentally concerned citizens.

Acknowledgment. We gratefully acknowledge the National Renewable Energy Laboratory of the U.S. Department of Energy for support of this work under contract XAR-3-13237-01-107811. Partial support by the JOULE I1 Program of the EEC is also kindly acknowledged. EF950011R