Modeling for Industrial Heat Exchanger Type Steam Reformer

Transcription

1 Korean J. Chem. Eng., 18(1), (2001) Modelng for Industral Heat Exchanger Type Steam Reformer Yong Ho Yu and Mhal H. Sosna* R&D Center, Samsung Engneerng Co., LTD, 9- Seongbok-R, Suj-Eup, Yongn, Kyungg-Do , Korea *NVF Technogas-GIAP, LTD, Vorontsovo pole st. 10, Moscow 10064, Russa (Receved 2 September 2000 accepted 8 November 2000) Abstract In a heat exchanger type steam methane reformer, the temperature profles and mole fractons along the axal dstance from the top of the reformer can be predcted by usng the channel model, consderng radaton heat transfer. The cross-secton of the reformer tube was dvded nto several channels as concentrc crcles and then heat transfer and mass transfer at the nterfaces between adjacent channels were consdered. Because the steam reformer s operated at hgh temperature, the radaton and convecton were combned nto one heat transfer coeffcent to smplfy the transfer analyss. Ths model predcts the ndustral plant data very well; therefore, t may be used wth confdence to desgn the ndustral heat exchanger type reformer. Key words : Heat Exchanger Type Reformer, Methane Steam Reformng, Convectve Radaton Heat Transfer, Channel Model INTRODUCTION The steam reformer s wdely used for the producton of ammona, methanol, hydrogen, acetc acd and fuel cells [Park et al., 1998]. The conventonal furnace type reformer shows hgher energy consumpton because of the fuel combuston, though t produces export steam. In the conventonal furnace type reformer, the reacton heat for steam reformng s suppled by radaton heat transfer from flame formed by the combuston of fuels such as hydrocarbons or plant off-gas. Therefore, currently, two types of reformers such as the Pd-Ru membrane reactor [Nam et al., 2000; Km et al., 1999] as well as the heat exchanger type reformer [Schneder et al., 1992; Nrula, 1990] have been studed and appled, because they ncrease methane converson and reduce energy consumpton, respectvely. In the heat exchanger type reformer, the reacton heat of steam reformng s suppled by heat transfer of the hot gas generated from the secondary reformer. Therefore, the hydrocarbon consumpton can be reduced by 20-0%. In the heat exchanger type reformer, there are no serous hot spots due to ndrect heat transfer from the heatng gas temperature, unless carbon depost takes place on the catalyst. Moreover, the heat exchanger type reformer has the advantage of an ecologcally clean process because of no fuel combuston equpment except for the fred heater for HDS (Hydro-Desulphurzaton System). Snce 1988 the heat exchanger type reformer has been developed and commercalzed by lcensors such as GIAP, ICI and M. W. Kellogg [Schneder et al., 1992; Nrula, 1990]. Xu and Froment [1989b] and Elnashae et al. [1992] predcted profles of desgn parameters such as temperature, reacton rates, effectveness factor and partal pressure of each components for the ndustral steam reformer, usng the one dmensonal heterogeneous model by Froment and Bshoff [1979]. Whereas Sosna et al. To whom correspondence should be addressed. E-mal: yyh1@samsung.co.kr 127 [1989] developed the channel model for the conventonal type steam reformer usng plot plant data and well predcted the temperature and methane mole fracton along axal dstance from the top. However, the modelng of heat exchanger type reformer, let alone an ndustral scale desgn, has not yet been publshed. In the present study, temperature profles, mole fractons and reacton rates are predcted by usng the channel model Sosna et al. [1989], whch consders the counter-current heat transfer between hot gas and tube wall ncludng radaton heat transfer. In order to predct the temperature profles at the reformer tube surfaces, both the convecton heat transfer and the radaton heat transfer are consdered. The predcted results from ths model are compared wth the data from an ndustral scale plant. THE CHANNEL MODEL WITH COUNTERCURRENT HEAT TRANSFER 1. Reacton Knetcs In the steam reformng reacton, the possble knetc rate expressons consdered are as follows: [Xu and Froment, 1989; Km and Lee, 1991]. CH 4 +H 2 O=CO+H 2 H=206.1 kj/mol (1) CO+H 2 O=CO 2 +H 2 H= kj/mol (2) CH 4 +2H 2 O=CO 2 +4H 2 H=165.0 kj/mol () Km and Lee [1991] summarzed the knetc equatons of prevous researchers and developed the rate equaton usng ndustral catalyst when hgh hydrogen partal pressure nterferes wth the methane reacton. Xu and Froment [1989a] developed the rate equaton of all the above reactons, but ther equaton s too complex and the effectveness factor for reacton (2) has a mnus value due to the swtchng of ts drecton of the reacton through the ndustral reactor modelng. Though Elnashae et al. [1992] smulated the ndustral reformer usng the equatons by Xu and Froment [1989a],

2 128 Y. H. Yu and M. H. Sosna the effectveness factor shows stff varaton along the axal dstance from the top. Therefore, they showed the monotonc decrease of the effectveness factor for reacton of CO 2 only after elmnaton of CO-shft reacton (2), but the effectveness factor was stll negatve at the end of reformer tube. However, because the above reactons are known to be reversble, many researchers [Khomenko et al., 1971; Allen et al., 1975; Ko et al., 1995; Twgg, 1989] suggest that the reacton mechansm can be descrbed wth two smple equatons. (1) and (2). Khomenko et al. [1971] developed the rate equaton for methane converson usng 0.0 mm nckel fol n order to obvate any pore dffuson lmtaton. They concluded that the water shft reacton approaches equlbrum. Ther equaton can be consdered to apply to any brand of nckel catalyst, compensated wth an effectveness factor. Therefore, n ths study, the reacton rate equaton by Khomenko et al. [1971] was adopted for modelng the ndustral scale methane steam reformer. The rate equaton s as follows: k R 1 = 1 P CH4 P H 2O[ 1 P CO P H2 ( K p1 P CH4 P H 2O) ] P + l 2 H 2O 2P H 2 + l P H2 Where, k 1 = T exp[ 720/RT], cm (STP)/(m 2 h atm) l 2 = T exp[19520/rt], atm 1 l 2 = T 6.5 exp[46700/rt], atm 2 Accordng to the brochure of HTAS regardng the catalyst R-67-7H, the N surface area and bulk densty of catalyst range from,500-5,000 m 2 /kg and 900-1,000 kg/m. In order to convert the unt of k 1 nto kmol/m h atm, t s necessary to multply the factor of 0.18 by average values of N surface area and bulk densty n the above equaton. The rate equaton for CO-shft reacton s as follows as Sosna et al. [1989]. R 2 =k 2 P CO [1 P CO2 P H2 /(K P2 P CO P H2O )] (5) January, 2001 (4) Fg. 1. Schematc dagram of heat exchanger type reformer. 1. Feed gas 5. Reformer tubes 2. Product gas 6. Cover tubes. Hot gas nlet 7. Pgtals 4. Hot gas outlet 2. Channel Model for Tubular Reformer The schematc dagram for the heat exchanger type reformer s shown n Fg. 1. The feed (natural gas and steam mxture), whch s ntroduced to the top of reformer, flows through the tubular reformer. The steam reformng reacton takes place through the catalyst bed nsde the tube whle the hot gas passes through the annular space between reformer tube and cover tube. The hot gas from the secondary reformer s ntroduced at the bottom of the reformer shell. The product gases from reformer tubes are collected n the header located at the top of the reformer and get out from the reformer. Because the reacton heat s suppled n the radal drecton from the tube wall, temperature gradents may be developed n the radal drecton due to the consumpton of heat as the reacton s occurrng. Whereas, the steam methane reacton rate depends on the temperature and a molar concentraton gradent n the radal drecton may develop. It s necessary to analyze the heat transfer and molar dffuson n the radal drecton. Therefore, the present model was developed under the concept that the cross-secton of the reformer tube s dvded nto several concentrc crcles and the ncrease n radus of concentrc has to be equal. The ncrement of radus, Dr, s equvalent to the volumetrc dameter of catalyst partcle to have a physcal meanng of channel sze, because the reactons are performed nsde of the catalyst partcle. Therefore, f we dvde the cross-secton nto n 1 concentrc crcles, n channels exst nsde the tube. If we assume that good mxng occurs nsde one channel and heat and mass transfer occur at the nterface between channels, the mass and heat balances for the j component n -th channel can be descrbed by followng equatons, respectvely: m + N j Z = ξ jk η k R k S + 2πr + 1 De 1 + j 1 r( C j C j ) k = 1 for =1,..., n 2πr De j r C 1 ( j C j ) m C p m T Z = Q k ( 1 ε)r k S + 2πr + 1 h + 1 r( T + 1 T ) k = 1 2πr h r( T T 1 ) where, n s defned as an nteger value of R/d v +0.5 Boundary condtons are, o At Z=0, N j =N j and T =T n for =1...n (8) At r=0, De j1 =0 and h 1 =0 (9) At r=d /2, T n+1 =T W, h n+1 =h W (10) The man parameters for estmatng the mole flow and temperature along the axal dstance from the top are mass transfer coeffcents besdes the reacton rate. The mass transfer coeffcent, De j for the j component between channels can be defned by the followng equaton of Sosna et al. [1989]: De j =0.28 D j Pr Re (11) Molecular dffusvty of j component to bulk phase, D j, can be (6) (7)

3 Modelng for Industral Heat Exchanger Type Steam Reformer 129 obtaned by the rule of Blanc s law calculaton wth the molecular dffusvty of one component j to the other component phases, whch was predcted by Sherwood et al. [1975]. On the other hand, the heat transfer coeffcent h n the catalyst bed can be defned by followng equaton: h =h o +v d v C p C /Bo (12) where h o mples the thermal conductvty and radaton heat transfer coeffcent. h o =k g +h r (1) Here, the radaton heat transfer coeffcent can be calculated by the followng equaton as defned by Kulkarn and Doraswamy [1980]: h r = T ε d v +(1 ε)/[1/(10 k g )+(1/ T ε d v )] (14) Moreover, the second term of Eq. (12) means the convectve contrbuton to the conductvty n the radal drecton caused by flow n the catalyst bed as defned by Westerterp [1986]. Accordng to hs study, the Bodensten Number, Bo can be predcted by the followng equaton as per the prevous researcher. Bo=8 [2 (1 2(d v /D)) 2 ] (15) In order to calculate the heat transfer coeffcent between nsde tube wall and catalyst bed, the followng correlaton equaton was adopted. Nu w =0.09 Re w 0.8 Pr 1/ (16) Where, Re w =2 r ρ g V/µ From the above Eq. (16), the heat transfer coeffcent between tube wall and n-th channel can be calculated. In order to supply the reacton heat by the hot gas from secondary reformer, a countercurrent double ppe heat exchanger was adopted. Therefore, the hot gas from secondary reformer passes upward through the annular space between cover tube and reformer tube and then s cooled down. However, for the sake of smplcty, the heatng gas would be consdered to be heated from the outlet gas temperature up to the nlet temperature of the hot gas from the secondary reformer. For the hot gas from the secondary reformer, the heat transfer can be consdered as follows. C p, hg m hg T hg / Z=q 1 +q 2 (17) Where, q 1 means the heat transfer from hot gas to reacton tube. Where q 2 means the heat transfer from hot gas to cover tube. Therefore, q 1 and q 2 can be defned by the followng equatons, respectvely, consderng radaton heat transfer: q 1 =(h c, 1 +h r, 1 )πdo (T g T w, o ) (18) q 2 =(h c, 2 +h r, 2 )πdc (T g T c ) (19) It can be consdered that all of the heat transfer from the hot gas to the cover tube by Eq. (19) s transferred to the reformer tube by radaton. q 2 =(h c, 2 +h r, 2 )πdo (T g T c )=h r, πdc (T c T w, o ) (20) Therefore, the equaton can be expressed n terms of T w, o, T g and heat transfer coeffcents. The convectve heat transfer coeffcent h c, 1 between the hot gas and the reformer tube wall and the convectve heat transfer coeffcent h c, 2 can be calculated by followng equatons [Bloch, 1967]. Nu 1 =0.86 (Do/Dc) Re c 0.8 (21) Where, Nu 1 =(Dc Do) h c, 1 /2k g Nu 2 =[1 0.14(Do/Dc) 0.6 ] Re c 0.8 (22) Where, Nu 2 =(Dc Do) h c, 2 /2k g The radaton heat transfer coeffcents can be calculated wth the emssvty as follows, h r, 1 = (ε w +1) ε g T g [(1 (T w, o ).6 )/2(1 T w, o )] (22) h r, 2 = (ε c +1) ε g T g [(1 (T c ).6 )/2(1 T c )] (2) h r, = (1 ε g ) ε c ε g T c [(1 (T c ).6 )/(1 T c )] (24) Throughout the above equatons, the temperature profles of the reformer tube, cover tube, reacton gas and hot gas can be obtaned for tube length wth an ncrease of length by 0.01 m. RESULTS AND DISCUSSION 1. Model Verfcaton The model developed n ths study ncludes all possble phenomena takng place n the heat exchanger type reformer through a rgorous heat and mass transfer analyss as well as the knetc equaton by Khomenko et al. [1971]. In order to verfy ths model, data from an ndustral scale heat exchanger type reformer, whch was used n ammona producton n Chna, and predcted values by ths model are presented n Table 1. Frst, from the top of the catalyst bed, the mass and heat balance equatons for each channel are numercally solved wth an ncrease of length untl the heatng gas temperature, T g s the temperature of heatng gas from the secondary reformer, T g, n. The accuracy of the model s proved by the actual results as shown n Table 1. The predcted methane content n the product s very accurate, 25.97%, compared wth the actual content, 25.71%, wth a dfference of 0.28%, and for CO 2 by 0.8%, Ths result s better than the results by Elnashae et al. [1992]. Though the steam reformng reacton tends to reach the equlbrum state, the actual product composton has a gap from ts chemcal equlbrum state at the outlet temperature. Whle the equlbrum constant can be calculated from the actual product composton, the equlbrum temperature can be deduced from the equlbrum constant because the equlbrum constant s a functon of temperature. The approach temperature s defned as the dfference between actual temperature and equlbrum temperature. Therefore, the approach temperature represents the catalyst actvty. Throughout the operaton results, the approach temperature to the equlbrum for methane converson s 7.6 o C. 2. Temperature Profles The temperature profles of the reacton gas, tube temperature, hot gas and cover tube are shown n Fg. 2. Even though the temperature of the cover tube s close to that of hot gas, the radaton heat transfer from cover tube to reformer tube for total heat flux s Korean J. Chem. Eng.(Vol. 18, No. 1)

4 10 Y. H. Yu and M. H. Sosna Table 1. Comparson between plant data and model predctons Reformer tubes: Effectve tube length=12. m Number=168 Inner/outer dameter=0.102/0.114 m Cover tubes: Inner dameter=0.15 m Process gas: Natural gas wth followng composton (vol% dry bass): CH %, C 2 H %, C H %, C 4 H %, C 5 H %, N 2.51% Feed condtons: Feed temperature, T n =49 o C Feed pressure, P n =2.5 Mpa Natural gas flow rate=10,877 Nm /h S/C rato=2.78 Hot gas data: Inlet temperature=955 o C Outlet temperature=515 o C Composton (vol%, wet bass) CH %, CO 7.75%, CO %, H 2 5.7%, N %, H 2 O 5.12% Flow rate=8,226 Nm /h Pressure=2.2 Outlet condton Plant data Model predctons Temperature, o C Pressure, MPa Product composton (dry bass), % CH CO H CO N Fg. 2. Temperature profles along the axal dstance from the top of reformer tube. 15% at the bottom of the tube and 5.5% at the top of the tube of heat flux accordng to the calculaton results. Therefore, radaton heat transfer should not be gnored for modelng of the heat exchanger type reformer because of hgh temperature operaton. The temperature dfference at the bottom of the reformer s hgher than Fg.. Profles of mole fractons along the axal dstance from the top of reformer. that at the top of the reformer, whch means that the much more methane steam reacton (1) takes place at the bottom of the reformer of the hgh temperature regon. The temperature profle s a typcal one for a countercurrent exchanger system wth reacton. Furthermore, the predcted temperature of the product gas s accurate compared wth the actual data; therefore, ths model can be consdered as reasonable. From ths temperature profle both of the cover tube and reformer tube, the thckness of the reformer tube can be determned for the purpose of cost reducton.. Mole Fracton Profles The predcted mole fractons of CH 4, H 2, CO 2, CO along the axal dstance are shown n Fg.. The methane fracton decreases gradually by 6 meters and ts decrease becomes greater over 8 meters. Ths trend s dfferent from the result by Xu and Froment [1989b]. Because they calculate the heat transfer wth the constant flame temperature and overall heat transfer coeffcent, the converson rate of methane along the axal dstance s dfferent. As can be seen n ths fgure, the CO content s very low up to 8 meters, because the shft reacton goes to equlbrum easly as reported by Khomenko et al. [1971]. However, the CO 2 fracton would not ncrease further over 11 meters, because the CO-shft reacton s exothermc, whle the hydrogen content ncreases further because the methane steam reacton s endothermc. 4. Effectveness Factor Many researchers [Namaguch and Kkuch, 1988; Twgg, 1989; Xu and Froment, 1989b; Elnashae et al., 1992] have reported the effectveness factor for ndustral nckel catalysts. Namaguch and Kkuch [1988] reported that the effectveness factor has a value of Twgg [1989] reported that the effectveness factor may, dependng on condton, only be as hgh as 0. at the nlet regon and as low as 0.01 at the ext. However, Xu and Froment [1989b] and Sosna et al. [1989] reported that the effectveness factor for methane converson has a value of The effectveness factor for CO 2 reacton has dscontnuty and shows mnus values due to the reverse drecton of ths reacton. However, Elnashae et January, 2001

5 Modelng for Industral Heat Exchanger Type Steam Reformer 11 al. [1992] reported that the effectveness factor has values rangng from to 0.08 along the axal dstance after they elmnated the CO-shft reacton. The effectveness factor for methane reacton n ths study can be obtaned as 0.04 by 8 meters from the top of the reformer tube and 0.12 over 8 meters, usng the wolf algorthm wth ntal guess n order to match all calculated values wth ndustral data. Ths value enables one to predct the nlet and outlet data of ndustral plant. It can be consdered the reason why the effectveness factor ncrease at the bottom of reformer s nduced from the dfferent heatng system compared wth conventonal fred heater type reformer. Moreover, the effectveness factor for methane steam reformng s close to ts defnton because the knetc equaton of Khomenko et al. [1971], whch s appled to ths model, s based on no mass transfer effect. CONCLUSIONS The model developed n ths study was verfed wth the data from an ndustral scale reformer that s operatng n the ammona ndustry. Ths model s accurate and reasonable for predctng the profles of each temperature and mole fractons along the axal dstance. Furthermore, radaton heat transfer s mportant for predctng accurate temperature profles as well as the heat transfer rate. Ths model can be used wth hgh confdence to desgn an ndustral heat exchanger type reformer. ACKNOWLEDGMENT The authors would lke to express ther apprecaton to Dr. Evenchck for computer programmng ths model. NOMENCLATURE Bo : Bodensten Number [C C p V d v /ε k g ] C : molar concentraton [kmol/m ] C p : heat capacty of gas mxture [kcal/ o C kg] D : molecular dffusvty to bulk phase [m 2 /h] De : effectve dffusvty n radal drecton [m 2 /h] Dc : nsde dameter of cover tube [m] D : nsde dameter of reformer tube [m] Do : outsde dameter of reformer tube [m] d v : volumetrc dameter of catalyst partcle [m] h : heat transfer coeffcent between channels [kcal/m 2 o C h] h r : radaton heat transfer coeffcent [kcal/m 2 o C h] h c : convecton heat transfer coeffcent [kcal/m 2 o C h] k : ntrnsc rate constant of reacton [kmol/m h atm] k g : thermal conductvty of gas [kcal/m o C h] K p : equlbrum constant of reacton N : mole flow rate [kmol/h] Nu : Nusselt Number Q : reacton heat [kcal/m ] Pr : Prandtl Number q : heat transfer rate [kcal/m h] r : channel radus [m] R : reacton rate [kmol/m h] Re : Reynolds Number n channel, 2 r ρ g V/µ Re c : Reynolds Number for double ppe, (Dc Do) ρ g V/µ S : cross-sectonal area of channel [m 2 ] T : temperature [K] Z : axal dstance from the top [m] Greek Letters ζ jk : stochometrc coeffcent of component, j n the k-th reacton ε : catalyst bed porosty or emssvty η : effectveness factor Subscrpts and Superscrpts c : cover tube g :gas : -th channel number j : j-th component k : reacton number w : tube wall 1 : from hot gas to reformer tube 2 : from hot gas to cover tube : from cover tube to reformer tube REFERENCES Allen, D. W., Gerhard, E. R. and Lkns Jr., M. R., Knetcs of the Methane-Steam Reacton, Ind. Eng. Chem. Process Des. Dev., 14, 256 (1975). Bloch, A. G., Thermal Radaton-Heat Transfer, Energy, Lenngrad (1967). Elnashae, S. S. E. H., Adrs, A. M., Solman, M. A. and Al-Ubad, A. S., Dgtal Smulaton of Industral Steam Reformers for Natural Gas usng Heterogeneous Models, Can. J. Chem. Eng., 70, 786 (1992). Froment, G. F. and Bschoff, K. B., Chemcal Reactor Analyss and Desgn, John Wley, New York (1979). Khomenko, A. A., Apelbaum, L. O., Shub, F. S., Snagovsk, S. and Temkn, M. I., Knetcs of the Reacton of Methane wth Water Vapor and the Reverse Reacton of Hydrogenaton of Carbon Monoxde on the Surface of Nckel, Knet. Katal., 12, 42 (1971). Km, D. H. and Lee, T. J., Knetcs of Methane Steam Reformng, HWAHAK KONGHAK, 29, 96 (1991). Km, J. H., Cho, B. S. and Y, J. H., Smulaton on the Methane Steam Reformng n Pd-Membrane Reactor, HWAHAK KONGHAK, 7, 210 (1999). Ko, K. D., Lee, J. K., Park, D. and Shn, S. H., Knetcs of Steam Reformng over A N/Alumna Catalyst, Korean J. Chem. Eng., 12, 478 (1995). Kulkarn, B. D. and Dorswamy, L. K., Estmaton of Effectve Transport Propertes n Packed Bed Reactors, Catal. Rev.: Sc. Eng., 22, 41 (1980). Nam, S. W., Yoon, S. P., Ha, H. Y., Hong, S. A. and Maganyuk, A. P., Methane Steam Reformng n a Pd-Ru Membrane Reactor, Korean J. Chem. Eng., 17, 288 (2000). Namaguch, T. and Kkuch, K., Intrnsc Knetcs and Desgn Smulaton n a Complex Reacton Network: Steam-Methane Reformng, Chem. Eng. Sc., 4, 2295 (1988). Korean J. Chem. Eng.(Vol. 18, No. 1)

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