Design and Control of the Dry Methane Reforming ... - ACS Publications

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Design and Control of the Dry Methane Reforming Process William L. Luyben* Department of Chemical Engineering, Lehigh University, Bethlehem, Pennsylvania 18015, United States ABSTRACT: The much lower cost of natural gas and the desire to reduce carbon dioxide emissions have stimulated interest in the dry methane reforming process in which these two gases react to produce synthesis gas. The ratio of hydrogen to carbon monoxide in the resulting synthesis gas is close to unity, which makes it suitable for feeding to the Fischer−Tropsch process to produce liquid transportation fuel. The dry methane reforming reaction is favored by low pressure, but high pressure synthesis gas is required in the downstream process. Higher reactor pressures reduce compression costs of the synthesis gas. However, higher pressures also require a higher reactor temperature to achieve high methane conversion. When a maximum reactor temperature limitation is encountered, high methane conversion can be maintained by operating with a higher than stoichiometric CO2-to-CH4 ratio. But this means an excess of carbon dioxide must be fed to the reactor and subsequently recycled, which imposes additional compression costs. These competing effects produce interesting design trade-offs, which are explored in this paper. The dynamic control of the process is also studied. The dominant issue is the modeling of the reactor, which is a fired furnace with combustion of fuel providing the endothermic heat of the syngas reaction. Dynamic modeling is achieved by using a tubular process reactor in which the heat flux is determined by the heat generated in a fuel−air combustion reactor. CH4 + CO2 ⇔ 2H 2 + 2CO

1. INTRODUCTION The recent surge in the supply of natural gas as a result of improved recovery methods (fracking) has been a “game changer” in the energy and chemical processing fields. More extensive use of natural gas has resulted from the lower cost. One major change is the shift from coal to natural gas for electric power generation, which has the additional advantage of reducing CO2 emissions because of the smaller amount of carbon in natural gas compared to that in coal. Natural gas as fuel also permits the use of a combined-cycle combustion turbine power generation system that is almost twice as efficient as a conventional boiler system. Another potential use of natural gas is the production of liquid transportation fuels in gas-to-liquid plants (Fischer−Tropsch or methanol synthesis) as explored by Baliban et al.1 Of course, natural gas has been widely used for many years to produce hydrogen via the steam (wet) methane reforming process. Water and methane react to form a synthesis gas that is rich in hydrogen.

Note that the hydrogen-to-carbon monoxide ratio is unity in this dry methane reforming reaction. Four moles of products are formed from two moles of reactants, so the reaction is favored by low pressure. It is endothermic, so it is favored by high temperatures. These two reforming reactions have been investigated for many years. For example, in 1948 Reitmeyer et al.2 studied the effects of temperature, pressure, and reactant ratio on conversion and coke formation. They demonstrated theoretically and experimentally that the hydrogen-to-carbon monoxide ratio in the synthesis gas can be adjusted by varying the ratios of water and carbon dioxide to methane in the feed. They pointed out that the prevention of coking by adjustment of temperature is required for different feed ratios. A state-of-the-art survey was presented in 1996 by Wang et al.3 Olsbye et al.4 presented an interesting reaction engineering study in which a fluidized-bed dry methane reforming reactor was designed. The effects of particle size and contact time were explored. The issue of heat transfer to provide the required heat addition to the reactor was considered. A kinetic analysis study was reported by Quiroga and Luna5 to determine reaction rate equations. Lim et al.6 explored a combined steam and dry methane reforming process to reduce carbon dioxide emissions. Yang et al.7 investigated feeding both natural gas and coal to reduce CO2 emissions in a coal-to-olefins process. A recent study by Noureldin et al.8 suggests that a combination of the three types of reforming processes to produce synthesis gas (steam, dry, and partial oxidation) may have some advantages.

CH4 + H 2O ⇔ 3H 2 + CO

Note that the hydrogen-to-carbon dioxide ratio is 3 in this steam methane reforming reaction. Four moles of products are formed from two moles of reactants, so the reaction is favored by low pressure. It is endothermic, so it is favored by high temperatures. Additional hydrogen can be produced by reacting the carbon monoxide with more water via the water−gas shift reaction. CO + H 2O ⇔ H 2 + CO2

The carbon dioxide is typically removed using pressure swing adsorption to generate a high-purity hydrogen stream for industrial uses. Hydrotreating of petroleum products for sulfur removal is a major consumer of hydrogen. An alternative to steam methane reforming is “dry methane reforming” in which methane is reacted with carbon dioxide. © 2014 American Chemical Society

Received: Revised: Accepted: Published: 14423

June 13, 2014 July 26, 2014 August 28, 2014 August 28, 2014 dx.doi.org/10.1021/ie5023942 | Ind. Eng. Chem. Res. 2014, 53, 14423−14439

Industrial & Engineering Chemistry Research

Article

Figure 1. Effect of temperature and pressure on conversion with FR = 1.

Figure 2. Effect of feed ratio and pressure on conversion at 1000 °C.

The hydrogen-to-carbon monoxide ratio in the produced synthesis gas varies with its final use. They point out that “the prospect of utilizing two green-house gases to produce a useful product makes dry reforming an important option to consider.” These studies have not explored the effect of pressure on the overall process in which both feed and product compression costs (both capital and power) can be very significant. This paper considers the inherent trade-offs involved in designing a dry

methane reforming process whose objective is to produce a synthesis gas product stream at a final pressure of 30 bar.

2. EFFECTS OF DESIGN VARIABLES ON CONVERSION Before getting into the details of the flowsheet, it is instructive to explore how several important variables affect the conversion of methane in the dry methane reforming. An Aspen RGIBBS reactor model is used that calculates reactor effluent conditions at chemical equilibrium given reactor temperature, reactor 14424

dx.doi.org/10.1021/ie5023942 | Ind. Eng. Chem. Res. 2014, 53, 14423−14439


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Figure 3. Effect of pressure and specified conversion on required temperature.

means that a reactor operating at this pressure and temperature would not require an excess of CO2 to attain 98% conversion and no CO2 recycle would be required. However, the synthesis gas would have to be compressed from 1 to 30 bar in an expensive multistage compressor train. To achieve the same 98% conversion with a reactor operating at 5 bar and 1000 °C, a CO2/CH4 ratio of 1.17 would be required. This means that the unreacted excess of CO2 leaving the reactor would have to be compressed from the low pressure of the stripper in the amine CO2 removal process (about 2 bar) up to 5 bar so that it can be recycled back to the reactor. However, now the compression of the synthesis gas would be less expensive (5 to 30 bar). 2.3. Conditions for Fixed Conversion. Figure 3 gives these results in a different and insightful format for the case when the process involves only reaction and compression with no CO2 recycle because the CO2/CH4 ratio is unity. The capital investments in a CO2 absorber and stripper units are avoided, as is the energy duty in a stripper reboiler. If the desired methane conversion is 99% and the maximum temperature limitation is 1000 °C, the reactor must be operated under vacuum conditions (0.763 bar). As we demonstrate in the next section, an expensive 4-stage compression train with intermediate cooling between compression stages is required to raise the synthesis gas up to the required 30 bar. In the numerical example considered in the next section, the compressor power consumption of this configuration is 18.3 MW with a compressor capital investment of $37,900,000. We will demonstrate that a process with CO2 recycle is less expensive in this situation (conversion 99% and temperature limit 1000 °C). If the desired methane conversion is only 98% with the same maximum temperature limitation, the reactor pressure can be raised to 1.553 bar. Now a less expensive 3-stage compressor train is needed with a lower power consumption (11.4 MW) and a reduced compressor capital investment ($30,600,000).

pressure, and the ratio of methane to carbon dioxide in the feed. Peng−Robinson physical properties are used. The components specified in the RGIBBS reactor are the reactants and products. Carbon is not considered. 2.1. Effect of Temperature and Pressure. Figure 1 gives a plot of methane conversion over a range of temperatures for several values of pressure. The feed to the reactor is the stoichiometric ratio of unity for the carbon dioxide-to-methane ratio (CO2/CH4 = 1). Under these conditions, there is no need for the recycling of CO2, so no recovery system is required (CO2 absorber and stripper with its energy requirements). Higher pressures reduce conversion at a constant temperature because of the increase in moles that occur during the reaction (Le Chatelier’s principle). Higher temperatures increase conversion at a constant pressure because the reaction is endothermic (higher chemical equilibrium constant at higher temperature). For example, if a methane conversion of 95% is desired, a reactor temperature of 863 °C is required when the pressure is 1 bar. A reactor temperature of 928 °C is required when the pressure is 2 bar. At 5 bar, the temperature must be above 1000 °C. Of course these higher reactor temperatures imply higher energy consumption for preheating the reactor feed. A maximum temperature limit of 1000 °C has been reported by Wang et al.3 in order to prevent the formation of nickel carbide on the surface of the catalyst. The maximum temperature limit decreases as the CO2/CH4 ratio is increased. They also report a minimum temperature limit of 800 °C to prevent the formation of coke for reactant ratios around unity. The minimum temperature limit increases with increasing pressure. 2.2. Effect of Feed Ratio. The results shown in Figure 1 display the effects of temperature and pressure when the CO2/ CH4 ratio is the stoichiometric value of unity. In Figure 2 the temperature is fixed at 1000 °C, and the effect of feed ratio on methane conversion is shown for several pressures. Higher CO2/ CH4 ratios increase conversion for the same pressure. At 1 bar pressure and 1000 °C, the chemical equilibrium conversion of methane is greater than 98% with a CO2/CH4 ratio of unity. This 14425



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Figure 4. No recycle; 99% conversion; 0.763 bar.

Figure 5. CO2 recycle; 99% conversion; 4 bar.

process. All the other costs associated with providing off-site facilities and services would typically increase the total plant investment by a factor of 3 (Julian9). Therefore, the effective annualization period is closer to 10 years.

The interplay among these process parameters indicates that there should be an optimum reactor pressure, which balances reactor performance with compression and CO2 recovery costs. In the following section, these trade-offs are quantitatively evaluated by finding the pressure that minimizes the total annual cost, which accounts for both energy consumption (compressor work and stripper reboiler duty) and capital investment (reactor, compressors, and heat exchangers). The payback period for capital is assumed to be three years, which may seem too small and put too much emphasis on capital. It should be noted, however, that the investment costs considered are just the installed cost of equipment used in the

3. DESIGN BASIS The dry methane reforming process consists of several units in series with the potential need for a recycle stream of carbon dioxide. Figure 4 shows a flowsheet in which there is no CO2 recycle. Figure 5 gives a flowsheet in which CO2 recycle is used. On the basis of the discussion in the previous sections, the following conditions and design parameters are assumed: 14426



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Table 1. 99% Conversion, 1000 °C Cases no recycle reactor pressure recycle lean reflux capital

energy

TAC

(bar) (kmol/h) (kmol/h) (kmol/h) (106 $) absorber stripper preheater FEHE reactor furnace KCO21 KCO22 K1 K2 K3 K4 cooler condenser total HX drums

CO2 recycle

0.763 0 0 0

2 26.79 900 157.3

3 54.60 2100 237.3

4 98.03 2500 297.3

5 144.8 4500 309.0

0 0 0 1.370 2.884 0 0 9.233 9.856 9.472 9.371 0.3227 0

0.2969 0.0556 0.1372 1.462 2.894 2.404 0 9.359 9.942 9.681 0 0.3415 0.1405 2.674 0.4079

0.2594 0.09054 0.2422 1.599 2.907 3.635 0 8.138 8.161 8.034 0 0.3536 0.2564 3.166 0.4079

0.2455 0.1228 0.3107 1.650 2.890 4.732 0 7.239 7.067 6.982 0 0.3730 0.3688 3.451 0.4079

0.2416 0.1433 0.3642 1.771 2.934 2.997 3.134 9.472 9.104 0 0 0.4096 0.4507 3.600 0.2190

total (106 $/yr) reactor furnace compression stripper QR

43.50

37.58

34.71

33.02

33.12

38.12 8.698 0

37.66 7.895 0.4097

38.50 6.764 0.9566

38.21 6.145 1.598

39.00 5.957 2.059

total (106 $/yr)

47.82 62.32

46.59 59.15

46.22 57.79

45.94 56.96

47.01 58.05

1. The reactor achieves chemical equilibrium. An Aspen RGIBBS reactor model is used in the simulations during the steady-state design. A different reactor model is used in later dynamic control studies so that realistic heat-transfer issues can be addressed between the combustion of fuel and the endothermic reaction process stream. 2. The reactor temperature is fixed at 1000 °C in most cases, but its impact on design is examined. The 1000 °C reactor temperature is assumed to keep carbon production low while preventing nickel carbide formation. 3. Methane conversion is fixed for each individual case and is varied to see its impact on the process design. 4. The flow rate of the fresh methane feed is 1000 kmol/h in the base case but will be varied in some cases considered. The methane feed pressure is high enough to not require compression even at the highest process pressure studied. 5. The flow rate of fresh carbon dioxide is 1000 kmol/h in the base case for those systems with no CO2 recycle. It changes slightly when there is recycle because the concentration of CO2 affects the small amount of water formed in the reactor because of the water−gas shift reaction. 6. The reactor capital cost is calculated as the cost of a fired furnace (Turton et al.10) with the heat duty necessary to provide the large endothermic heat of reaction. Reactor energy is assumed to cost $16.8 per GJ. 7. Multistage compressors are used for synthesis gas compression and for carbon dioxide compression (both fresh feed and recycle). The same compression ratio is used in each stage. The number of stages is determined by

8.

9.

10.

11.

14427

keeping compressor discharge temperatures below a maximum of 200 °C (Walas11). Polytropic compression ASME efficiency is 80%. There are two process−process heat exchangers: the feedeffluent heat exchanger on the reactor (FEHE) and the economizer heat exchanger (Preheater) between the absorber and the stripper on the lean and rich solvent streams. These are designed with a 20 °C minimum approach temperature. The gas−gas FEHE has an overall heat-transfer coefficient of 0.17 kW/(m2 K). The liquid− liquid Preheater in the amine unit has an overall heattransfer coefficient of 0.57 kW/(m2 K). The water-cooled gas heat exchangers (cooler after the reactor FEHE, the stripper condenser, and the coolers between compression stages) have overall heat-transfer coefficients of 0.20 kW/(m2 K) with cooling water inlet and outlet temperatures of 32 and 43 °C, respectively. Carbon dioxide is removed from the synthesis gas in an 11stage absorber using a lean amine solvent (8 mol % monoethanol amine, 92 mol % water). The operating pressure in the absorber is 0.6 bar lower than the reactor pressure. In the process with no recycle and no absorber, the suction pressure of the first compressor is 0.2 bar lower than the reactor pressure because of the 0.1 bar pressure drops over the FEHE and cooler. Solvent is fed to an 11-stage stripper on stage 3, which operates at 2 bar. Low-pressure steam ($7.78 per GJ) is used in the stripper reboiler. Overhead vapor is partially condensed at 50 °C. The gas stream from the stripper reflux drum is combined with the fresh carbon dioxide and dx.doi.org/10.1021/ie5023942 | Ind. Eng. Chem. Res. 2014, 53, 14423−14439


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required to achieve 99% methane conversion. Note that the fresh CO2 feed flow rate is 1118 kmol/h. The total CO2 entering the reactor is 1216 kmol/h, which gives an entering concentration of 54.6 mol % CO2. The resulting CO2/CH4 ratio is sufficient to achieve the specified methane conversion of 99% at the 1000 °C reactor temperature and 4 bar reactor pressure. The fresh CO2 feed and the CO2 recycle streams are combined and compressed from 1 to 4.2 bar, requiring 1.96 MW of work. Thus, the recycle process requires a CO2 compressor in addition to synthesis gas compressors. Fresh methane feed is added, and the total stream is fed to a large feed-effluent heat exchanger (4188 m2) to raise its temperature to 980 °C. The reactor heat input from the combustion of fuel is 73.12 MW. The flow rate into the reactor is 2216 kmol/h. The flow rate leaving the reactor is 4195 kmol/h. The reactor effluent composition is 44.1 mol % H2, 50.2 mol % CO, 2.2 mol % CO2, 0.2 mol % CH4, 36 ppm of CO2, and 3.2 mol % H2O. Table 2 compares the reactor inlet and outlet conditions for the nonrecycle and recycle processes. There is very little water in

compressed up to a pressure 0.1 bar greater than the reactor pressure. 12. Liquid from the reflux drum is combined with a water makeup stream (to account for the water removed from the lean solvent in the absorber) and pumped back to the top of the stripper.

4. CASES WITH 99% CONVERSION In this section we explore two alternative process designs. Both are based on a desired methane conversion of 99%. In the first process, there is no carbon dioxide recycle, so there are no absorber and stripper. In the second process, recycle is used to achieve the same conversion at higher reactor pressures so that compression costs can be reduced. However, now there are capital and energy costs for recovering and recycling the excess carbon dioxide fed to the reactor. In both cases, reactor temperature is at its assumed maximum of 1000 °C. 4.1. No Recycle Process. Figure 4 gives the flowsheet in which a low reactor pressure of 0.763 bar is required to achieve 99% conversion. The fresh methane and carbon dioxide flow rates are both 1000 kmol/h. The FEHE provides the 28.82 MW of heat, which preheats the feed to the reactor to 980 °C and requires 3148 m2 of heat-transfer area. The reactor is a fired furnace that burns fuel to provide the large heat required to keep the temperature at 1000 o C with the very endothermic reforming reaction. This fuel cost is a major expense for the process ($38,120,000 per year in this example). Notice that the feed to the reactor is 2000 kmol/h. The reactor effluent is 3980 kmol/h of synthesis gas with a composition of 49.6 mol % H2, 0.3 mol % CH4 (1% unconverted methane), 49.9 mol % CO, 951 ppm (molar) CO2, and 0.2 mol % H2O. A heat exchanger using cooling water reduces the temperature of the synthesis gas leaving the FEHE from 163.4 to 50 °C before it enters the first compressor. The cooling duty is 3.67 MW. The pressure entering the compressor train is 0.563 bar (0.1 bar pressure drops in the two heat exchangers downstream of the reactor). The heuristic that the compression ratio should be the same in each stage is assumed. The total compression ratio is 30/0.563 = 53.28. A 4-stage compressor train is required if the compressor discharge temperatures must be lower than 200 °C. Thus, the compression ratio per stage is (55.28)0.25 = 2.702. The resulting compressor discharge pressures and temperatures are shown in Figure 4. The total compressor power is 18.3 MW. Detailed economic information is given for this system in the first column of Table 1. The capital investment for the reactor and FEHE is $4,260,000. Capital investment in the four compressors is almost an order of magnitude larger ($37,900,000). Fuel cost in the reactor is $38,120,000 per year. Power cost in the compressor train is $8,698,000 per year. There are no capital or energy costs for an absorber−stripper CO2 removal system because there is no recycle. The total annual cost of this no-recycle case with 99% conversion and 1000 °C reactor temperature is $62,320,000 per year, using a payback period of 3 years. In the following section we compare this process with one in which CO2 is recycled so that a higher reactor pressure can be used. 4.2. CO2 Recycle Case. Figure 5 gives the flowsheet for a process in which there is CO2 recycle and a recovery system (absorber and stripper) is required. With a 1000 °C reactor temperature and a reactor pressure of 4 bar, a recycle flow rate of 98.03 kmol/h (the stripper reflux drum vapor product) is

Table 2. Reactor Conditions with and without CO2 Recycle; 99% Conversion, 1000 °C Cases no recycle

recycle

R in

R out

R in

R out

molar flow rate (kmol/h) H2 CO CO2 CH4 H2O

0 0 1000 1000 0

1974 1986 3.78 10 6.21

0 0 1211 1000 5.59

1851 2108 92.5 10 134.1

total (kmol/h)

2000

3980

2216

4196

the reactor effluent of the nonrecycle process, but there is a small amount in the reactor effluent of the recycle process because of the higher CO2 concentration. But when this gas is fed into the absorber, it picks up a significant amount of water from the MEA/water solvent. This has a cooling effect in the absorber and requires that a water makeup stream be fed to the stripper reflux drum. The cooled gas is fed into the bottom of the absorber operating a 3.4 bar. Lean solvent (3500 kmol/h) is fed to the top. The gas leaving out the top of the absorber is at 57.1 °C, and the enriched solvent leaving the bottom has a temperature of 57.2 °C. The absorber off-gas is compressed in a 3-stage compression system with intermediate cooling to 50 °C between stages. Remember that the nonrecycle process required a 4-stage compression train. The pressure at the top of the absorber is assumed to be 0.6 bar lower than the reactor pressure because of pressure drops through the two heat exchangers and the absorber column. The total compressor work to raise the pressure from 3.4 bar to the desired 30 bar is 11.44 MW. This should be compared with the 18.3 MW required in the 4-stage compressors in the norecycle process. Some water is condensed and removed between compression stages. The final product is 4002 kmol/h of synthesis gas with a composition of 46.3 mol % H2, 52.7 mol % CO, 36 ppm of CO2, 0.3 mol % CH4, and 0.8 mol % H2O. The synthesis gas product streams of the two processes have slightly different compositions with less hydrogen and more CO2 in the recycle process. 14428



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Table 3. Conversion Cases (No Recycle; 1000 °C) conversion (% CH4) 90

93

95

97

99

8.994 1135 1135

5.948 1086 1086

4.096 1056 1056

2.372 1027 1027

0.763 1000 1000

1.609 2.979 2.222 2.270 6.256 6.308 0 0 0.3734 0.3325 0 0 0.1751

1.534 2.946 1.781 1.846 8.078 8.334 0 0 0.3545 0.3414 0 0 0.1596

1.532 2.924 0.9805 0 9.749 9.874 0 0 0.3319 0.3443 0 0 0

1.459 2.890 0.5782 0 8.554 8.329 8.264 0.3463 0.3318 0.3338 0 0

1.376 2.884 0 0 9.233 9.856 9.472 9.371 0.3227 0.3384 0.3131 0.3418 0

total (106 $/yr) reactor furnace compression

22.52

25.37

25.74

31.10

43.50

39.68 4.388

39.13 5.185

38.78 5.319

38.44 6.522

38.12 9.698

total

44.07

44.31

44.10

44.96

47.82

excess CH4

(106 $/yr)

4.605

2.944

1.913

0.9293

0

TAC

(106 $/yr)

56.18

55.72

54.79

56.26

62.32

reactor pressure fresh CH4 fresh CO2 capital

energy

(bar) (kmol/h) (kmol/h) (106 $) FEHE reactor furnace KCO21 KCO22 K1 K2 K3 K4 cooler HX1 HX2 HX3 HXCO2

5. OPTIMUM DESIGN CONVERSION WITH NO RECYCLE The specified conversion has a profound effect on the optimum reactor pressure selection. Higher conversion requires lower pressure or higher CO2 recycles. In the previous section we demonstrated that with a 99% conversion specification, a recycle process was better than a no-recycle process. The selection of 99% conversion was arbitrary but typical of many chemical processes. We would expect that high methane conversion would be desired because wasting a valuable feed stock is usually not economical. The value of the feeds and the value of the products are typically much more significant than energy or capital costs, as pointed out by Douglas12 several decades ago. For example, if we assume a natural gas price of $5 per 1000 SCF, a 1% loss of unconverted methane in the 1000 kmol/h fresh feed represents an additional raw material cost of about $340,000 per year. We need to have a fair comparison among the various cases. If the conversion is changed without any adjustment of the fresh feeds, the amount of product synthesis gas will change. The procedure used in this paper is to fix the amount of hydrogen produced. With the ratio of the two fresh feeds kept at unity (for the no-recycle process), their flow rates are adjusted in each case to produce the same molar flow rate of hydrogen. As conversion is reduced, more methane and CO2 must be fed above the basecase amounts of 1000 kmol/h. The cost of the addition methane above the 1000 kmol/h base case is considered an operating cost. The total annual cost is the sum of energy cost, excess methane cost, and annual capital cost (total capital divided by a three year payback period).

The process shown in Figure 5 gives the minimum total annual cost for a process with recycle and 99% conversion at 1000 °C. The optimum reactor pressure is 4 bar as demonstrated in the second through fifth columns in Table 1. Reactor pressures over the range of 2 to 5 bar are evaluated. Results show a minimum total annual cost (TAC) of $56,960,000 per year at a reactor pressure of 4 bar. This should be compared with a TAC of $62,320,000 per year for the no-recycle process. It is clear that the use of CO2 provides a more economical process for this case (99% conversion and 1000 °C reactor temperature). In the recycle process, compression costs for the carbon dioxide increase as reactor pressure is increased, but compression cost for the synthesis gas decrease. However, higher reactor pressure requires higher CO2 recycle flow rates, which increase the loads on the absorber and stripper. More lean solvent is required, which increases stripper reboiler duty and increases the capital cost of the stripper (larger diameter). The absorber capital cost actually decreases slightly because its diameter is smaller at higher operating pressures. This quantitative comparison of the no-recycle and the recycle processes shows that the recycle process has better economics than the more simple process without recycle. This occurs because compression costs are very high with the low reactor pressures required for no-recycle operation. It should be emphasized that the conversion has been fixed at 99% in the cases considered up to now. However, does a 99% conversion design represent the economic optimum? In the following section, we explore the effect of changing the methane conversion and its impact on design parameters and economic results. 14429



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optimum design if the temperature limitation is lower? Lower reactor temperatures require lower reactor pressures to achieve the same conversion, which increases synthesis gas compression costs. We would expect that the optimum conversion would be shifted to a lower value to keep compression cost at reasonable levels. However, lower conversions result in more excess methane having to be fed to achieve the same hydrogen production rate. We explore this question in this section. Table 5 gives detailed results. Three cases are shown with conversions of 90, 93, and 95%. The corresponding reactor pressures are 3.412, 2.260, and 1.507 bar. Synthesis gas compression costs (capital and work) increase as conversion increases while carbon dioxide feed compression costs decrease slightly. The cost of purchasing additional methane decreases as conversion increases. Therefore, the basic design trade-off is between compression cost and raw material cost. The design that gives the minimum TAC has a 93% conversion when reactor temperature is limited to 900 °C. As expected, the TAC of this design ($60,510,000 per year) is higher than the TAC of the 1000 °C design ($54,790,000 per year, shown in Table 3). Both compression costs and fresh feed costs are higher in the lower conversion, lower reactor pressure process when reactor temperature is reduced. Figure 7 gives the flowsheet of the 900 °C design. Comparing these results with Figure 6 (the 1000 °C design) shows that a two-stage synthesis gas compression system is used in the 1000 °C process but a three-stage system is needed in the 900 °C process. This is due to the lower reactor pressure in the 900 °C process (2.26 versus 4.096 bar).

We assume that the cost of the carbon dioxide feed is negligible because of the desire to reduce CO2 emissions. In fact, there may be a credit for using CO2 if carbon tax legislation becomes a reality. The base-case conditions are those of the 99% conversion norecycle process described earlier in which the synthesis gas product contains 1973.8 kmol/h of H2. The two fresh feeds are 1000 kmol/h each in the base case. As an example, suppose we select a conversion of 95%. Reactor pressure can be raised to 4.096 bar with a 1000 °C reactor temperature, which reduces synthesis gas compression (but increases CO2 compression costs). However, the fresh feeds must be increased to 1056 kmol/h to produce the desired 1973.8 kmol/h of hydrogen in the synthesis gas. So the cost of supplying an additional 56 mol/h of methane is added to the operating cost. This amounts to almost $2,000,000 per year of additional expense to purchase methane. We consider two cases for the process with no recycle. The first is using a maximum temperature of 1000 °C. The second is when the maximum is reduced to 900 °C. 5.1. Reactor Temperature 1000 °C. Table 3 gives detailed results for a range of conversions. The minimum total annual cost case is for a 95% conversion. The fresh feeds are 1056 kmol/h, so an additional 56 kmol/h of methane must be included as an operating expense. This amounts to $1,913,000 per year. As conversion is decreased, synthesis gas compression cost decreases because reactor pressure increases. At 99% conversion, four compressors are required with three intercoolers. But no compression of the CO2 feed is required because it is supplied at 1 bar and the reactor pressure is 0.763 bar. No additional methane feed is required. At 97% conversion (reactor pressure, 2.372 bar), three synthesis gas compressors and one small CO2 compressor are required. Additional methane feed (27 kmol/h) is required. At 95% conversion (reactor pressure, 4.096 bar), two synthesis gas compressors and one small CO2 compressor are required. Additional methane feed (56 kmol/h) is required. At 93% conversion (reactor pressure, 5.948 bar), two synthesis gas compressors and two CO2 compressors are required. Additional methane feed (86 kmol/h) is required. Finally, at 90% conversion (reactor pressure, 8.994 bar), two synthesis gas compressors and two CO2 compressors are required. Additional methane feed (135 kmol/h) is required. The 95% conversion case gives the minimum total annual cost. Table 4 compares the syngas compositions at 95 and 99% conversions. Figure 6 gives the flowsheet for this optimum design. Conversion is 95%, reactor pressure is 4.096 bar, and reactor temperature is 1000 °C. The additional methane fresh feed required is 56 kmol/h. 5.2. Reactor Temperature 900 °C. All of the previous cases have used a reactor temperature of 1000 °C. What happens to the

6. OPTIMUM DESIGN CONVERSION WITH RECYCLE In Section 4 we fixed the conversion at 99% and demonstrated that a process with recycle was better than one without. In Section 5 we considered the no-recycle process and found the optimum conversion. Now we apply the same approach to the recycle process. A feed ratio of CO2/CH4 = 1.05 is assumed, which will require a CO2 recovery system. The fresh feeds are adjusted to keep the production of hydrogen constant at 1973.8 kmol/h. A range of conversions are explored. Reactor temperature is 1000 °C. Results are given in Table 6. The case with the minimum total annual cost is with 95% conversion. Notice that the 99% conversion case with feed ratio of 1.05 has a TAC ($60,240,000 per year) that is lower than the corresponding no-recycle case shown in Table 3 ($62,320,000 per year). So the recycle process is better than the no-recycle process with this 99% conversion. This is what we found in Section 4. However, if we compare the optimum conversion cases, the no-recycle process at its optimum 95% conversion ($54,790,000 per year) is better than the recycle process at its optimum 95% conversion ($56,830,000 per year). The TAC decreases in the recycle process as conversion is decreased but not as rapidly as in the no-recycle process. A feed ratio of CO2/CH4 = 1.05 has been used in the calculations discussed above. During the review of this paper, a perceptive question was asked about the effect of designing for different feed ratios. To answer this question, runs were made with a feed ratio of CO2/CH4 = 1.1. As shown in Figure 2, a higher ratio permits the use of a higher reactor pressure for a fixed conversion. For a 95% conversion and 1000 °C, reactor pressure increases from 5.928 to 7.428 bar, which reduces work in the syngas compressors. However, it increases work in the CO2

Table 4. Synthesis Gas; No Recycle; 1000 °C conversion (%) flow rate H2 CO CO2 CH4 H2O

(kmol/h) (mol %)

95

99

4118 47.9 49.5 0.50 1.30 0.80

3980 49.6 49.9 951 ppm 0.30 0.20 14430



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Figure 6. Fixed H2 production; 1000 °C; 95% conversion; 4.096 bar.

Table 5. Conversion Cases (No Recycle; 900 °C)

3. Heat duty in the stripper increases by 0.744 MW because of the higher CO2 recycle and the higher lean solvent flow rate. This represents an energy cost of $182,500 per year. Therefore, operating the recycle process with higher feed ratios is unattractive. Thus, the most economical design does not have CO2 recycle and operates with a 4.096 bar reactor pressure when reactor temperature is 1000 °C. In the following sections, we study the dynamics of this process and develop an effective control structure.

conversion (% CH4) 90 reactor pressure fresh CH4 Fresh CO2 capital

93

95

(bar) (kmol/h) (kmol/h) (106 $) FEHE reactor furnace KCO2 K1 K2 K3 cooler HX1 HX2

3.412 1131 1131

2.260 1084 1084

1.507 1054 1054

1.432 2.975 3.920 7.464 7.619 7.530 0.3702 0.3292 0.3523

1.358 2.944 2.656 8.665 8.862 8.764 0.34057 0.3405 0.3455

1.309 2.925 1.667 10.09 9.874 9.969 0.3319 0.3505 0.3584

total (106 $/yr) reactor furnace compression

31.98

34.29

36.84

39.60 6.320

39.10 7.119

38.79 8.014

total

45.92

46.22

44.80

excess CH4

(106 $/yr)

4.467

2.860

1.852

TAC

(106 $/yr)

61.04

60.51

60.93

energy

7. DYNAMIC MODELING To explore the dynamics of the dry methane reforming process, a realistic furnace model must be used. Fuel and air are fed into the furnace through burners, and the heat of combustion is transferred into the process stream, which is flowing inside tubes in the radiant and convective sections of the furnace. This heat transfer is an extremely complex mixture of radiant and convective effects. It can be approximately modeled by using a tubular reactor in which a value of the flux of heat (GJ/m2) is used to size the dimensions of the process tubes (total heattransfer area). Then, with a fixed area, the heat generated by burning fuel with air in the combustion zone is transferred into the process stream. 7.1. Reformer Reactor Model. The RPlug reactor model is used in the Aspen simulations with the “Reactor Type” specified to be Reactor with specif ied external heat flux prof ile. Jones and Pujado13 recommend heat-flux values for fired furnaces in the of range 6800 cal/(s m2) (9000 Btu/(h ft2). The heat-transfer rate in the reactor is 61.55 Gcal/h, so a heat-transfer area of 2500 m2 sets the size of the furnace tubes. Selecting a tube diameter of 0.1 m and a tube length of 10 m, the number of tubes is 800. The tubes are filled with catalyst (voidage, 0.6; density, 2000 kg/m3). Heat transfer to the solid catalyst is specified. Reactor pressure drop is 0.1 bar. The number of intervals in the Report option in Aspen Plus is set at 50 (giving a 50-lump model in Aspen Dynamics). Figure 8 shows the temperature and composition profiles in the reactor using a fixed heat flux of 6800 cal/(s m2) and a heat duty 61.55 Gcal/h. Note the very large drop in temperature at the inlet portion of the tubes because of the endothermic reaction.

compressors. The net effect is a slight decrease in compressor work of 0.629 MW, which is worth $333,000 per year. However, the higher feed ratio has three other effects that negatively impact costs. 1. The flow rate of fresh methane must be increased by 15 kmol/h to maintain the same production rate of hydrogen in the syngas. This represents an increase in raw material cost of $512,000 per year. 2. The increase in CO2 recycle (65.1 instead of 44.5 kmol/h) increases flow rates through the reactor, which increases furnace duty by 1.22 MW, representing an increase in fuel cost of $646,000 per year. 14431



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Figure 7. Fixed H2 production; 900 °C; 93% conversion; 2.26 bar.

Table 6. Conversion Cases with FR = 1.05; Recycle; 1000 °C

simulation with an RPlug model, reaction kinetics must be specified. There is little need for a complex rigorous kinetic model when studying the dynamics of a reactor operating near some design conditions. Therefore, for simplicity, we develop approximate power law, first-order kinetics for the reaction.

conversion (%) reactor pressure recycle lean reflux capital

energy

cost additional CH4

93

95

97

99

(bar)

7.972

5.927

3.961

1.912

(kmol/h) (kmol/h) (kmol/h) (106 $) absorber stripper FEHE preheater reactor furnace KCO21 KCO22 K1 K2 K3 total HX drums

55.53 1800 140.3

44.59 1500 143.4

34.42 1400 162.5

24.93 900 161.6

0.1962 0.08283 1.607 0.2084 2.993

0.2058 0.07455 1.557 0.1882 2.972

0.2300 0.0714 1.555 0.1851 2.950

0.3045 0.0556 1.470 0.1388 2.930

3.379 3.466 7.194 7.060 0 2.987 0.2193

2.836 2.960 8.096 8.502 0 2.863 0.2193

4.316 0 7.158 7.268 7.172 3.009 0.4397

2.912 0 10.91 9.587 9.366 2.463 0.4079

total (106 $/yr) reactor furnace compression stripper QR

27.33

28.50

32.31

38.00

39.10

39.56

39.21

38.88

4.822 0.8209

5.269 0.6850

6.126 0.6337

8.152 0.409

total

44.73

44.83

45.33

47.03

(106 $/yr)

3.577

2.499

1.493

0.5398

CH4 + CO2 ⇔ 2H 2 + 2CO

First we assume that the methane conversion in the very large reactor (800 tubes) is 93 mol %, which is 2 mol % lower than chemical equilibrium. Then we guess a value for the preexponential factor in the forward reaction k0F, which will be adjusted to give 93% conversion. A typical activation energy (50 000 kJ/kmol) for the forward reaction is specified, giving the forward reaction rate 9 F. 9F = kFPCH4PCO2 = k 0Fe−50 000/ RT PCH4PCO2

Pressures are in Pa, and temperature is in K. The reaction is vapor phase, and the reaction rate has units of kmol/(s m3). The kinetics of the reverse reaction are found by first determining the temperature dependence of the chemical equilibrium constant KEQ. A REquil reactor model is run at two temperatures (900 and 1000 °C), and the ratio of the concentrations of the products to the reactants at chemical equilibrium gives KEQ. KEQ(T ) =

TAC

(10 $/yr)

57.42

56.83

57.60

2

yCH yCO 4

2

ln KEQ = 31.447 − 2.958 × 104 /T

For a specified value of k0F, the specific reaction rate of the reverse reaction kR can be calculated. kR =

6

yCO yH

60.24

kF KEQ

The value of k0F that gives 93% methane conversion is found to be 5 × 10−9.kmol s−1 m−3 Pa−2. The final expressions used for the forward and reverse reaction rates are given in the following equations.

7.2. Reaction Kinetics. The RGibbs reactor model predicts a thermodynamically limited maximum methane conversion of 95% at 1000 °C and 4.096 bar. For the purposes of dynamic 14432



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Figure 8. (a) Temperature profile in reactor with fixed heat flux and (b) composition profiles in reactor with fixed heat flux.

9RWGS = 10−11 e−50 000/ RT PCO2PH2

9F = 5 × 10−9 e−50 000/ RT PCH4PCO2 9R = 1.094 × 10−22 e−195 900/ RT PCOPH2

This gives a water concentration in the synthesis gas of 0.2 mol %. 7.3. Furnace Model. To model the combustion of fuel (methane) with air in the fire box of the furnace, a simple RGibbs model is used. The required flow rate of fuel to provide the heat to the process stream is 715.3 kmol/h. The flow rate of air required to give 5% excess oxygen above the stoichiometric amount needed for complete combustion of methane to water and carbon dioxide is a factor of 10 larger than the fuel flow rate.

Note that the activation energy of the reverse reaction is much larger than the activation energy of the forward reaction because the reaction is highly endothermic. Because the RGibbs model predicted a small amount of water formed, a simple reverse water−gas shift reaction is assumed with the following kinetics. 14433



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Figure 9. Flowsheet equations for transferring combustion heat into reactor.

Figure 10. Basic control structure.

length, the number of tubes is 12 000. The tube volume is 9.4 m3. The shell volume is assumed to be equal to the tube volume. The Aspen Plus file is exported into Aspen Dynamics as a pressure-driven simulation.

CH4 + 2O2 ⇔ 2H 2O + CO2

The fire box temperature is assumed to be fixed at 1200 °C, and the pressure is 1.1 bar. The 61.55 GJ/h of heat removed from the fire box matches the heat input required in the process stream. The heat generated and transferred depends on the fuel flow rate. In the control structure developed below, the reactor exit temperature is controlled by manipulating the flow rate of the fuel, with the flow rate of the air ratioed to the flow rate of the fuel. The Aspen Dynamics feature of Flowsheet equations is used to transfer the furnace (“COMBUST”) heat into the reactor. Figure 9 gives these calculations. Remember that the heat removed in fire box is negative and the heat added in the reactor is positive according to the Aspen convention. 7.4. Equipment Sizing. There are three units in the process with significant volume that provide the dynamic capacitance of the unit (reactor, fire box and FEHE). Remember that all streams are gas phase, so the dynamics are quite fast. The size of the reactor has already been determined. The 800 tubes have an internal volume of 62 m3. The size of the fire box (volume outside the tubes) is assumed to be equal to the tube internal volume. The FEHE is very large (3740 m2) and has significant volume. Assuming heat exchanger tubes 0.01 m in diameter and 10 m in

8. BASIC CONTROL STRUCTURE Figure 10 shows the basic control structure for this gas system. The control loops are listed below. 1. Methane feed is flow controlled by manipulating the control valve in the feed line. 2. Carbon dioxide feed is ratioed to the methane feed using a multiplier whose one input is the methane flow rate and the other is a fixed ratio of 10. The output of the multiplier is the remote set point of the carbon dioxide flow controller, which is on cascade. The flow controller output sets the power to the carbon dioxide feed compressor. 3. Reactor exit temperature is controlled by manipulating fuel flow rate. A fuel-to-feed ratio provides feedforward action for disturbances in feed flow rate (ratio at design conditions is 0.6674). The temperature controller output signal adjusts the ratio. A 0.5 min lag in the feed flow signal provides dynamic compensation. The output signal from the multiplier is the remote set point of the fuel flow controller. 4. Air flow rate is ratioed to fuel flow rate using a multiplier. 14434



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Figure 11. Effect of lag (0.5 min).

Figure 12. Thermal dynamics of isolated reactor.

A 3 min deadtime is used in the reactor exit temperature loop. Relay-feedback testing and Tyreus−Luyben tuning give KC = 0.13 and τI = 48 min with a temperature transmitter range of 800−1200 °C and an output range of 0−1.2 ratio. Pressure controllers have KC = 1 and τI = 5 min. Figure 11 shows why a small lag is needed in the feedforward loop. An instantaneous increase in fuel when feed is increased

5. Exit temperatures from the two water-cooled heat exchangers are controlled by manipulating heat removal. 6. Reactor pressure is controlled by manipulating power to the first syngas compressor. 7. Discharge pressure of the first compressor is controlled by manipulating power to the second syngas compressor.

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Figure 13. 20% feed flow rate disturbances.

Figure 14. Furnace firing control structure.

causes temperature to spike up. Although this is a gas-filled system, the large amount of catalyst (50 000 kg) in the reactor tubes provides thermal capacitance. Figure 12 illustrates the dynamics of the reactor in isolation for a 10% increase in the set point of an inlet stream flow controller. The exit flow rate responds in about 0.05 h, but the exit temperature takes 0.3 h to come to steady state.

Figure 13 gives responses to large 20% disturbances in the set point of the methane flow controller. Solid lines are for an increase, and dashed lines are for a decrease. Effective regulatory level control is achieved. Peak transient deviations in reactor exit temperature are about 10 °C before returning to the set point in about 3 h. The bottom right graph in Figure 13 shows that the methane content in syngas yCH4 varies somewhat with throughput as 14436



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Figure 15. Action of the control structure. Fuel lags air for increases and leads air for decreases.

Figure 16. Comparison of basic and firing controls.

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Figure 17. Ramp 20% feed disturbances with firing controls.

It is important to note that there is a deterioration of performance of the reactor temperature control loop when the firing controls are used. The lower right graph shows larger transient deviations of the reactor temperature (drops to 940 °C at 0.25 h). The presence of the lag changes the controller tuning. Remember that these responses are for a very large 20% step disturbance in feed flow rate, which is an extreme test of the effectiveness of the control structure. Smaller disturbances or ramp disturbances are less severe and produce smaller deviations. Figure 17 gives responses when the feed flow rate is ramped either up or down 20% over a 0.5 h period with the firing control structure in use. The reactor temperature drops to only 973 °C for the ramp up (instead of the 940 °C with the step disturbance).

expected because of changes in reactor residence time. Reactor pressure deviations are small and quickly eliminated.

9. FURNACE FIRING CONTROL STRUCTURE In the basic control structure discussed in the previous section, the flow rate of the air is ratioed to the flow rate of the fuel. In many fired furnace applications it is important to never operate the fire box with a shortage of air, i.e., there should always be excess oxygen to guarantee complete combustion of the fuel. An excess of fuel can cause black stack gas, which is frowned on by the EPA. In extreme cases when a large excess of fuel accumulates in the fire box, the rapid combustion when the air flow rate eventually increases can cause an unsafe overpressuring of the furnace. This safety and operational problem can be eliminated by using a control structure in which the air is always in excess. When the reactor temperature controller desires an increase in fuel, the air should be increased first. When the reactor temperature controller desires a decrease in fuel, the fuel should be decreased before the air is decreased. Figure 14 shows a furnace firing control structure14 that achieves these objectives. The control elements use high and low selectors in combination with a lag. A 1 min first-order lag is used in the simulations reported below. Figure 15 illustrates the action of the control structure. For an increase in firing (left graph), the air comes up first (“leads”) before the fuel comes up. For a decrease in firing (right graph), the air goes down after (“lags”) the fuel goes down. Figure 16 compares the responses using the basic control structure and using the firing control structure. The disturbance is a 20% increase in throughput. The lower left graph shows that the oxygen composition of the stack gas undergoes a significant drop when the basic control structure is used. The firing control structure maintains an excess of oxygen.

10. CONCLUSION The dry methane reforming process presents some interesting design trade-offs among the numerous variables. Reactor pressure affects reactor performance and compression costs in inverse ways, which presents a major trade-off. Reactor temperature is also important with operation at a maximum limit being the most economical. Using higher CO2/CH4 ratios by recycling of CO2 does not appear to be economically attractive. The more simple process without recycle shows smaller total annual cost at its corresponding optimum methane conversion. An effective control structure is developed and tested for large disturbances. Realistic modeling of the fired-furnace reactor as implemented in Aspen software is demonstrated.

■

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel.: 610-758-4256. Fax: 610-7585057. 14438



Article

Notes

The authors declare no competing financial interest.

■

REFERENCES

(1) Baliban, R. C.; Elia, J. A.; Floudas, C. A. Biomass and natural gas to liquid transportation fuels: Process synthesis, global optimization and topology analysis. Ind. Eng. Chem. Res. 2013, 52, 3387−3406. (2) Reitmeyer, R. E.; Atwood, K.; Bennerr, H. A.; Baugh, H. M. Production of synthesis gas by reacting light hydrocarbons with steam and carbon dioxide. Ind. Eng. Chem. 1948, 40 (4), 620−626. (3) Wang, S.; Lu, G. Q.; Millar, G. J. Carbon dioxide reforming of methane to produce synthesis gas over metal-supported cataysts: State of the art. Energy Fuels 1996, 10, 896−904. (4) Olsbye, U.; Wurzel, T.; Mleczko, L. Kinetic and reaction engineering studies of dry reforming of methane over Ni/La/Al2O3 catalyst. Ind. Eng. Chem. Res. 1997, 36, 5180−5188. (5) Quiroga, M. M. B.; Luna, A. E. C. Kinetic analysis of rate data for dry reforming of methane. Ind. Eng. Chem. Res. 2007, 46, 5265−5270. (6) Lim, Y.; Lee, C.; Jeong, Y. S.; Song, I. H.; Lee, C. J. Optimal design and decision for combined steam reforming process with dry methane reforming to reuse CO2 as a raw material. Ind. Eng. Chem. Res. 2012, 51, 3387−3406. (7) Yang, S.; Yang, Q.; Man, Y.; Xiang, D.; Quia, Y. Conceptual design and analysis of a natural gas assisted coal-to-olefins process for CO2 reuse. Ind. Eng. Chem. Res. 2013, 52, 14406−14414. (8) Noureldin, M. M. B.; Elbashir, N. O.; El-Halwagi, M. M. Optimization and selection of reforming approaches for syngas generation from natural/shale gas. Ind. Eng. Chem. Res. 2014, 53, 1841−1855. (9) Julian, M. Process Economics lecture. Lehigh University, Bethlehem, Pa, 2013. (10) Turton, R., Bailie, R. C., Whiting, W. B., Shaelwitz, J. A. Analysis, Synthesis and Design of Chemical Processes, 2nd ed; Prentice Hall: New York, 2003. (11) Walas, S. M. Chemical Process Equipment, Selection and Design; Butterworth-Heinemann: Boston, MA, 1990. (12) Douglas, J. M. Conceptual Design of Chemical Processes; McGrawHill: New York, 1988. (13) Jones, D. S. J., Pujado, P. R. Handbook of Petroleum Processing; Springer: Dordrecht, The Netherlands, 2006; p 1047. (14) Luyben, W. L. Chemical Reactor Design and Control; Wiley: Hoboken, NJ, 2007; p 412.

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Design and Control of the Dry Methane Reforming ... - ACS Publications

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