Comparative Performance Analysis of Industrial Scale Catalytic Steam Reformer with Membrane Steam Reformer

Anton Friedl, Jiří J. Klemeš, Stefan Radl, Petar S. Varbanov, Thomas Wallek (Eds.) Proceedings of the 28th European Symposium on Computer Aided Process Engineering June 10th to 13th, 2018, Graz, Austria. © 2018 Elsevier B.V. All rights reserved. https://doi.org/10.1016/B978-0-444-64235-6.50124-8

Comparative Performance Analysis of Industrial Scale Catalytic Steam Reformer with Membrane Steam Reformer Arun Senthil Sundaramoorthy.,a Arun Prem Anand Natarajan.,a Sundaramoorthy Sithanandamb,* a

Department of Chemical Engineering, Sri Venkateswara College of Engineering, Sriperumbudur, 602117, India

b

Department of Chemical Engineering, Pondicherry Engineering College, Puducherry, 605014, India

[email protected]

Abstract The objective of this work is to develop a framework to compare the performance of an industrial scale Catalytic Steam Reformer (CSR) with that of a Membrane Steam Reformer (MSR) and to decide on the choice of MSR as a viable alternative to CSR for production of hydrogen. For this purpose, a mathematical model is developed in this work to represent the industrial scale operation of MSR. Using the mathematical models, operations of CSR and MSR are optimized to achieve maximum methane conversion. The advantage of higher methane conversion achieved in MSR as compared to CSR has to be weighed against the additional cost of Pd membrane used in MSR. A mathematical equation is presented in this work to calculate and compare the unit costs of hydrogen as a function of optimal values of various operating parameters obtained for CSR and MSR. Keywords: Process intensification, Process design, Process optimization, Membrane steam reforming, Techno-economic analysis

1. Introduction With increasing demand for hydrogen as a clean fuel and with abundant availability of natural gas resources across the globe, Catalytic Steam Reforming of methane is gaining prominence as the most attractive process for production of hydrogen. In a conventional CSR, the maximum achievable conversion of methane is limited to equilibrium conversion. Whereas, in a MSR , the equilibrium is shifted in favour of higher methane conversion by letting the hydrogen produced in the reactor to permeate out through a Pd membrane. Although a number of studies have been reported (Silva et al, 2016) on laboratory scale MSR, there is no evidence of any industrial application of MSR technology. This is mainly due to enormous cost of preparation of Pd membranes (Criscuoli et al., 2001), which is expected to decrease in future with growing number of industrial applications of MSR. The present work addresses the need to develop a mathematical model to appropriately represent the industrial scale operation of MSR and evaluate the economic feasibility of replacing the existing CSR technology by MSR technology. For this purpose, the design

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of industrial scale side fired CSR reported in Rajesh et al. (2000) is taken and is suitably altered to represent the industrial scale MSR. A mathematical model for the MSR is developed by modifying the model equations for CSR reported in Rajesh et al. (2000) to account for transport of hydrogen through Pd membrane. A systematic comparative analysis of optimum performances of CSR and MSR is carried out for various surface areas of Pd membrane tubes used in MSR. An equation for calculating the unit production cost of hydrogen as an explicit function of optimal values of feed and operating parameters of CSR and MSR is presented in this work for the purpose of economic feasibility analysis.

2. Process Description of CSR and MSR Sweep Gas (Steam)

Reactor Feed

Reactor Feed

di

1 2

Z=0

Z=0

2

1

3 1: Pd Membrane 5 2: Catalyst Pellet 3: Reactor Tube Wall 6 4: Flame 5: Burner 6: H2 Diffusion through Memb.

3 4

4

1: Catalyst Pellet 2: Reactor Tube Z Wall dZ 3: Flame 4: Burner

Z dZ

Z=L Z=L

Reactor Exit

do

Sweep Gas + Hydrogen Reactor Exit

1(b)

1(a)

Figure 1: (a) Single reactor tube in a CSR (b) Single reactor tube in a MSR Three chemical reactions listed below take place in the Catalytic Steam Reformer: CH 4 + H 2O =CO + 3H 2

;

CO + H 2O = CO2 + H 2

;

CH 4 + 2 H 2O =CO2 + 4 H 2

Kinetic rate equations reported by Xu and Froment (1989) are used in this work. Industrial scale side fired CSR (Fig.1a) reported in Rajesh et al. (2000) is taken in this study. CSR contains nt = 176 numbers of catalyst filled reactor tubes (outer diameter do = 102 mm, inner diameter di = 79.5 mm, length L = 11.95 m) placed in a refractory lined furnace. The catalyst is Ni on Al2O3 pellet of diameter Dp = 17.4 mm. The furnace has 112 burners. The feed gas is a mixture of CH4, H2O, CO2, H2 and N2. FCH4 is the molal feed rate of methane. S/C, H/C, D/C and N/C are respectively the molar ratios H2O/CH4, H2/CH4, CO2/CH4 and N2/CH4 in the feed gas.

Comparative Performance Analysis of Indstrial Scale CSR with MSR

701

The design of CSR (Fig.1a) is altered into a MSR as shown in Fig.1b. In MSR, each reactor has two concentric tubes. Inner empty tube (radius ro ) is made of Pd membrane of thickness δ = 7.5µ m through which a sweep gas (steam) is passed. Annular section is the reactor packed with catalyst pellets through which feed gas is sent. H2 permeates through Pd membrane wall into the inner tube at a molar flux of J H 2 ( Sieverts law) J H2 =

(

Qo

)

Qo  E   E exp  − P  PH0.5 exp  − P − PH0.5 = 2 2p δ δ  RT   RT

(

   PyH 2 

)

0.5

(

− Pp y H 2 p

)

0.5 

 

(1)

Pp is the permeate pressure and yH 2 p is mole fraction of hydrogen in permeate. Values

of parameters Qo and E p are taken from Johannessen et al. (2005)

3. Mathematical Modelling of CSR and MSR In this work, one dimensional mathematical model for side fired industrial scale CSR reported in Rajesh et al. (2000) is adapted with some modifications. Further, a model is developed to represent the industrial scale side fired MSR by suitably modifying the model equations of CSR to account for flux of hydrogen through the Pd membrane. The model equations for CSR and MSR are combined using a switching parameter α which takes a value α = 1 for MSR and α = 0 for CSR. Model Equations for Catalytic Reactor Section of CSR and MSR: Over all mass balance equation is  4π ro L   dG = −α   J H 2  dz  Ac  

(2)

Component molal balance equation is written in terms of xi =

yi M

as

 x  4π ro L    (2 xi − 1)  2π ro L   dxi L ρbηi Ri = + α (1 − β )  i   J H 2  + αβ    J H 2  dz G  G  Ac    G   Ac 

(3)

i = CH 4 , H 2O, CO, CO2 , H 2

Over all heat balance equation is dT L = dz GC pg

   π DiU  α  A   c

  4U  + (1 − α )    di

 III   U (2π ro )   η j r j −∆H j − α  m   (Twi − T ) + ρb  (T − T p )   Ac     j=I

∑ (

)

(4)

Pressure drop in reactor channel is calculated using Kozney – Karman equation 1.75 LG 2 (1 − ε b ) dP = − dz ϕ s D pε b3 ρ g

(5)

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Ac : cross sectional area of the reactor channel, yi : mole fraction, M : average molecular weight, Ri : rate of formation of chemical species, G : mass velocity of gas , β : switching parameter ( β = 1 for i = H 2 and β = 0 for i ≠ H 2 ), ηi : catalyst effectiveness factor (equal values of 0.03 assigned for all three reactions as reported in E. Johannessen et al. 2005), C pg : specific heat capacity of gas, T : gas temperature, Twi : inner tube wall temperature, U : catalyst bed side heat transfer coefficient, Tp :

permeate gas temperature, U m : overall heat transfer coefficient across the membrane. Model Equations for Permeate Tube Section of MSR: Hydrogen mole balance, Heat balance and Pressure drop equations are as follows,

dyH 2 p dz

(

 2π r L 1 − y o H2 p =  Gp 

dT p )  J=

 

H2

;

dz

 (2π r L)U o m   G pC pp 

  (T − T p )  

;

dPp dz

= −

2 4 fL  v p  d p  2g   

(6), (7), (8) G p : molal flow rate of permeate gas per unit area, G pp : specific heat capacity of the

permeate gas, f : friction factor, v p : gas velocity, d p : permeate tube diameter

4. Simulation and Optimization of CSR and MSR operations The main objective of this work is to quantify the improvement in performance of MSR over CSR when the surface area ( Am ) of Pd membrane is increased. For this purpose, identical tube lengths (L) and catalyst bed volumes are taken for both MSR and CSR and Am is varied by changing the inner membrane tube radius ro . To accommodate equal catalyst bed volumes in CSR and MSR, the inner diameter Di of the outer tube of MSR is increased as Di = γ di . By changing the value of parameter γ > 1 , the values of ro and Am are changed respectively = as ro

di = Am π di γ 2 − 1Lnt . γ 2 − 1 and 2

Matlab®R2015a was used to simulate CSR and MSR models. Simultaneous stiff ordinary differential equations (Eq. 2-8) were solved using function subroutine ode15s to obtain the profiles of mole fractions, temperature and pressure for different feed and operating conditions. Factors that influence CH 4 conversion ( M c ) in CSR are feed temperature Ti , feed pressure Pi , furnace gas temperature Tg , S/C and H/C. In addition to all these factors listed for CSR, conversion of CH 4 in MSR is also influenced by permeate feed temperature T pi , permeate feed pressure Ppi and sweep ratio Sr which is the ratio of feed rate of sweep gas (steam) to the feed rate of methane.

Comparative Performance Analysis of Indstrial Scale CSR with MSR

703

In this work, CSR and MSR operations were optimized to achieve maximum CH 4 conversion ( M c ) for fixed values of FCH4 = 1000 kmol/h, D/C = 0.091 and N/C = 0.02, subject to operational constraints and bounds (Rajesh et al. 2000) listed here, Two ≤ 1200 K , F ≤ 5000 kmol / h , 725 ≤ Ti ≤ 900 K , 2400 ≤ Pi ≤ 3000 kPa , 2.0 ≤ S / C ≤ 6.0 , 0.01 ≤ H / C ≤ 0.5 , 1375 ≤ Tg ≤ 1650 K , 1.0 ≤ Sr ≤ 3.0 , 600 ≤ T pi ≤ 800 K , 200 ≤ Ppi ≤ 300 kPa . MSR is optimized for different values of γ varying between 1.15

and 2. The constrained optimization problem is solved using fmincon optimization routine of Matlab® R2015a.

5. Results and Discussion Optimal values of feed and operating variables and the corresponding values of methane conversion ( M c ), H 2 production rate ( PH 2 ) and CO2 production rate ( PCO2 ) are

reported in Table 1. In CSR, an optimal methane conversion of only 49% was achieved, whereas 99.9% conversion was achieved in MSR for γ = 2 .Although methane conversion increases steadily with increase in γ , the increase is not very significant for γ > 1.75 suggesting that γ > 1.75 is not economical. Further, optimal operation of MSR requires the feed gas to be at a higher pressure Pi and furnace gas to be at a higher temperature Tg compared to CSR in order maintain a higher hydrogen concentration gradient across the Pd membrane. Table 1: Comparison of Key Optimum Performance indices of CSR and MSR Operating Parameters

CSR

MSR γ = 1.15 γ = 1.30 γ = 1.40 γ = 1.65 γ = 1.75 γ = 2.00

Tg ( K )

1,575.5

1,644.5

1,650.0

16,50.0

1,643.2

1,630.4

1,606.7

Ti ( K )

Pi (kPa ) S /C Sr Tpi ( K )

900.0 2,400.0 3.879 -

878.8 2,505.6 3.879 2.998 800.0

891.8 3,000.0 3.879 3.000 800.0

900.0 3,000.0 3.878 2.595 800.0

900.0 2,519.4 3.877 3.000 800.0

900.0 2,988.5 3.863 2.971 799.6

900.0 2,539.7 3.647 3.000 793.1

Ppi (kPa )

-

200.0

200.0

200.0

200.0

200.0

200.0

PH 2 (kmol / h)

1,817.9 346.7 49.0

3,039.3 712.6 77.6

3,497.1 836.2 88.7

3,668.0 876.7 93.0

3,891.8 925.7 98.9

3,945.2 951.4 99.8

3,938.0 940.3 99.9

PCO2 (kmol / h)

M c (%)

The economics of choice of MSR over CSR has to take into account influence of various parameters listed in Table 1 on the fixed and operating costs of steam reformer.

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In this work, an explicit equation (Eq. 9) is derived to calculate the unit cost of production of hydrogen CPH as a function of these parameters (Table 1) CPH =

(9)

CRM + CFR + CSP + COT + COC H P 2

æ M ö  CH = CM ç C ÷ F CRM CH 4 + CH 2O ( S / C + a Sr ) F 4 è 100 ø

(

)

M ö æ  CO + C  CSP = CMS ç1 - c ÷ F CH 4 + CCS - CCO2 P HS P H 2 2 100 ø è

; ;

 CH + a C æ p d g 2 - 1 Ln ö CFR = CR F Pd ç i t÷ 4 è ø

d2 ù  d1 + a éc S F  = COC c1 FP ê 2 r CH 4 PPi ú i ë û

( )

(

)

b3 ù  b1 + a T b2 + a é a S F  COT = a1 FT ê 3 r CH 4 TPi ú 2 f i ë û

( )

( )

(

)

Here, CM , CH 2O CCO2 : unit costs of methane, steam, carbon dioxide, CR : annualized fixed cost of catalytic reactor tubes per unit kmole of methane processed, CPd : annualized fixed cost of palladium membrane tubes per unit area, CMS , CCS , CHS : annualized fixed and operating costs of separators for methane, carbon dioxide hydrogen per unit kmole of separation, ai , bi , ci , di : coefficients of empirical cost correlations for COT cost for thermal heating and COC cost for feed gas compression. By comparing the unit production costs of hydrogen CPH calculated for CSR and MSR for different values of γ , one can make a decision on the choice of MSR over CSR and the value of γ at which the cost is minimum.

6. Conclusions In this paper, a mathematical model has been developed to represent the industrial scale operation of a side fired Membrane Steam Reformer (MSR). The mathematical models of CSR and MSR are used to compare their optimal performances and to evaluate the suitability of MSR over CSR. The economics of choice of MSR in place of CSR is to be evaluated by taking in to account a number of cost factors. A frame work is developed in this paper to calculate the unit cost of production of hydrogen by representing various cost factors as functions of optimal feed and operating parameters.

References A. Criscuoli, A. Basile, E. Drioli, O. Loiacono, 2001, An economic feasibility study for water gas shift membrane reactor, Journal of Membrane Science181,21-27 E. Johannessen, Kristin Jordal,2005,Study of a H2 seperating membrane reactor for methane steam reforming at conditions relevent for power processes with CO2 capture, Energy Conversion and Management,46,1059-1071 J.K. Rajesh, S. K. Gupta, G.P. Rangaiah, A. K.Ray, 2000, Multiobjective optimization of steam reformer performance using genetic algorithm, Ind. Eng. Chem. Res., 39,3,706-717 J. D. Silva, Cesar Augusto Moraes de Abreu, 2016, Modelling and simulation in conventional fixed-bed and fixed-bed membrane reactors for the steam reforming of methane, International Journal of Hydrogen Energy,41,27,11660-11674 J. Xu, G.F. Froment, 1989, Methane steam reforming, methanation and water-gas shift: I. Intrinsic kinetics, AIChE Journal, 35,1,88-96