High Purity Hydrogen with Sorption-Enhanced Steam Methane

Article pubs.acs.org/IECR

High Purity Hydrogen with Sorption-Enhanced Steam Methane Reforming in a Gas−Solid Trickle Bed Reactor Ana Obradović* and Janez Levec National Institute of Chemistry, 1000 Ljubljana, Slovenia ABSTRACT: A sorption enhanced-steam methane reforming (SE-SMR) process was investigated in a countercurrent gas−solid trickle flow reactor packed with regularly stacked catalyst. The stacked catalyst was made of corrugated Pt/Ni/Al2O3 plates in the form of static mixers. Experiments were carried out in a temperature range of 550−600 °C, total pressure of 4.0 bar, and water to methane ratio of 4.0 in the reactor feed. Within the investigated conditions, experimental results offered a solid proof of concept for new continuous SE-SMR operation for production of hydrogen where sorbent can be regenerated separately from the catalyst. Experimental data were reasonably well described by a plug flow model for both the gas and solid phases. The mathematical model was further used to optimize the continuous SE-SMR operation: it could be demonstrated that the level of the hydrogen purity in the reactor exit stream could be controlled by a multistage reactor consisting of alternating catalytic active and stainless-steel inert static mixers.

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INTRODUCTION Hydrogen is considered the most promising and important fuel for the future because it offers a nonpolluting, inexhaustible, efficient, and potentially cost-effective energy source.1 Besides its application as a fuel, hydrogen is widely used in industrial production plants for a variety of hydro-treatment processes. It is mostly delivered in the form of compressed gas or, in case of large consumption, it is generated on-site. The most common process for hydrogen production is steam methane reforming (SMR) that was first implemented industrially in 1930.2 SMR can be described by three global equilibrium limited reactionstwo reforming reactions and water gas shift (WGS) reaction reforming reaction 1:

with in situ CO2 capture by a sorbent, has the potential to reduce the cost of hydrogen production. The concept is based on Le Chatelier’s principle, which states that both the conversion of reactants and the net reaction rate in an equilibrium controlled reaction can be increased by selectively removing some of the reaction products from the reaction mixture. In the case of SMR, CO2 can be continuously and effectively removed by a solid sorbent such as CaO CaO + CO2 ↔ CaCO3

(4)

In that way, the system of reforming, WGS, and sorption reaction lumps into an overall irreversible reaction. One of the first comprehensive investigations on the SE-SMR process was accomplished by Air Products and Chemical, Inc. (1997).3 They developed many process advantages, such as elimination of costly pressure swing adsorption (PSA). The combined process can be carried out either in a circulated fluidized bed reactor or in a reactor with multiple packed tubes.4,5 In both cases sorbent is admixed with catalyst. The circulated bed system operates continuously due to bed transport between the reformer and sorbent regenerator, whereas the second arrangement is discontinuous in operation regarding a single tube. A number of studies have been devoted to CO2 solid sorbents.4 Ca-based oxides from a naturally occurring precursor such as limestone and dolomite have the

° = 205.8 kJ mol−1 Δr H298

CH4 + H 2O ↔ CO + 3H 2

(1)

water gas shift reaction: CO + H 2O ↔ CO2 + H 2

° = − 41.2 kJ mol−1 Δr H298 (2)

reforming reaction 2: CH4 + 2H 2O ↔ CO2 + 4H 2

° = 164.6 kJ mol−1 Δr H298 (3)

The reforming reactions are highly endothermic and are favored by high temperatures, whereas WGS is mildly exothermic and proceeds with high CO conversion at lower temperatures. The thermodynamic nature of this reaction system dictates therefore a multiple step process to achieve high purity hydrogen from the SMR process. Sorption-enhanced steam methane reforming (SE-SMR), which combines SMR © XXXX American Chemical Society

° = −178.1 kJ mol−1 Δr H298

Special Issue: Tapio Salmi Festschrift Received: Revised: Accepted: Published: A

May 1, 2017 July 4, 2017 July 5, 2017 July 5, 2017 DOI: 10.1021/acs.iecr.7b01832 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Article

Industrial & Engineering Chemistry Research

Figure 1. Schematic diagram of SE-SMR experimental setup: (1) gas cylinders, (2) pressure reducers, (3) mass flow controllers, (4) water reservoir, (5) pump, (6) water preheating section, (7) gas preheating section, (8) vessel for collecting solids, (9) zone preventing gas-phase recirculation, (10) reaction zone with structured packing, (11) solid preheating and distribution zone, (12) bin, (13) coolers, (14) filter, (15) separator, (16) back pressure regulator.

were made of corrugated pure nickel plates, which were catalytically enhanced by the successively deposited layers of Pt (30 nm), Al2O3 (1.5 μm), and again Pt (30 nm), followed by the annealing treatment.6,7 The hydraulic diameter and porosity of the element were 4.3 mm and 0.83, respectively. The total geometric surface area of the plate catalytic section was 2889 cm2, and the average geometric surface area of a single structured plate catalyst element was 321 cm2. After every two catalytic elements, a ring with the fixed o.d. of 43 mm, and i.d. conically reduced from 43 to 32 mm in a distance of 10 mm, was inserted to redistribute the flow of solid particles from the tube wall back to the reactor core. The inert SS element at the bottom was used to stabilize the gas-phase flow at the entrance of the catalytic section. Below the plane where the reactants were introduced into the reactor, another SS element was installed to prevent gas recirculation in the reactor entrance region, see zone 9 in Figure 1. Three SS elements along with two spacings were mounted above the catalyst bed, see zone 11 in Figure 1, to preheat and distribute the sorbent (CaO) particles at the entrance of the reaction zone. Thermocouples were installed along the bed, with the tips positioned 1 mm from the reactor wall in order not to disturb the solid-phase flow. Distilled water was fed by a Beckman 114 M Solvent Delivery Module pump into 1/4 in. long SS tubing with inert packing in which it was vaporized. The flow of methane (99.999%), hydrogen (99.999%), and nitrogen (99.999%) was monitored by mass flow controllers (Brooks 5850TR) and mixed with water vapor prior to entering an electrical preheater made of coiled 1/4 in. SS tubing. The gaseous mixture was heated to the operation temperature and was fed into the reactor below the first packing SS element. A special feature of

advantage of being widely available and inexpensive. They also possess high CO2 capacity and react rapidly over a wide range of temperatures and pressures. However, their primary disadvantage is associated with the high temperature required for sorbent regeneration and deactivation with multicycle operation.4 Recently, Obradović et al.6,7 introduced a novel Pt/Ni/Al2O3 structured plate-type catalyst for SMR. A catalyst in the form of static mixer is an excellent example of a regularly stacked-type packing recommended for gas−solid trickle flow operation due to reduced pressure drop, low static hold-up of solid particles, and rapid radial solid distribution throughout the column.8 Static mixer elements with corrugated plates have been recently efficiently employed for a solid sorption process (CaO−CO2) by Obradović and Levec9 in a countercurrent gas−solid trickle flow reactor, where they served as a solid flow distributor and gas-phase mixer. This work describes proof of concept of SESMR carried out in a continuously operated gas−solid trickle flow reactor with the same structured packing elements as described in Obradović and Levec,9 but now catalytically active for SMR. This operation is advantageous over the fluidized bed concept, since the used solid sorbent can be regenerated separately from the catalyst.

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EXPERIMENTAL SECTION The experimental apparatus for the SE-SMR experiments is depicted in Figure 1. The reactor of 43 mm i.d. was made of a 316Ti tube and consisted of nine (9) catalytically active structured elements (42 mm in height) and an identical one at the bottom but made of inert SS material, see Figure 2a. The structured elements B

DOI: 10.1021/acs.iecr.7b01832 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Figure 2. (a) Schematic representation of the reaction zone with structured packing; (b) single structured plate-catalyst element, tilted perspective; (c) SEM picture of plate-catalyst element.

the apparatus was feeding nitrogen via a mass flow controller through the bin and solid sorbent feeder in the amount of 10% of the total volumetric flow, see Figure 1. This arrangement eliminated convective transport of gaseous components from the reaction zone into the solid preheating zone and bin, and improved solid sorbent feeding. The gaseous product reaction mixture passed through the filter and condenser/separator to remove dust and water prior to gas stream analysis. The reactor exit stream was continuously checked for CO2 and CO concentration by an IR instrument (Binos 1001, Rosemount), and the composition analysis was done by a gas chromatograph (Agilent 7890 A). The sorbent used in the experiments was three times calcined CaO of the fraction 500−710 μm. The calcination of the sorbent was performed separately from the trickle bed reactor in nitrogen recirculating atmosphere for 10 min at the temperature of 840 °C (heating rate of 10 °C min−1, Nabertherm furnace) and the sorbent was cooled to the

ambient temperature before it was used for SE-SMR experiments. The same CaO material was used in the previous study,9 which explains why for this set of experiments it was calcined for the third time. The solid sorbent feeder was designed as a gear wheel of 70 mm in diameter and 20 mm in width, which was mounted offcentrally in a housing. Twenty half-cylindrical grooves of radius 4 mm were equally spanned around the wheel circumference. The wheel was driven by an asynchronous motor (Siemens). A magnetic coupling (Dexter) with torque 3 N m was interposed between the wheel and the motor. The motor speed was controlled by an inverter (FRENIC-Mini, Fuji Electric). A catalyst stabilization procedure was carried out at 500 °C and 1.0 bar. It was performed in a gaseous mixture of constant composition (FH2/FCH4 = 0.15; FH2O/FCH4 = 4.0; FN2/FCH4 = 0.85), which entered the bed at the volumetric flow rate of 3.0 LSTP min−1, whereas 0.3 LSTP min−1 of pure nitrogen was fed through the empty bin. After the catalyst reached a stable C


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Industrial & Engineering Chemistry Research operation, the flow of reactants was switched to pure nitrogen until the gas-phase in the reactor exit stream was free of reactants and products. At this point, the bin was opened and filled with sorbent (CaO). The pressure of 4.0 bar was built up with nitrogen and the temperature was raised to 550 °C. Shortly after, a gaseous reaction mixture, see Table 1, was introduced into the reactor and the flow of nitrogen through the bin was simultaneously established.

determined during the SMR kinetics determination experiments.7 • There is no diffusion resistance within the catalyst.7 • Only the carbonation reaction proceeds in the reactor parts with no catalytically active elements. • There is no diffusion resistance for the three times calcined sorbent used in this study. For the fresh (one time calcined) sorbent, the average pore diameter was 759.0 nm, see Table 2. With every calcination, the average pore diameter increases.12

Table 1. Experimental Conditions Used in the SE-SMR Study at Two Different Temperaturesa T

Qtotal

QN2,bin

ṁ s

gradT

expt

(°C)

(LSTP min−1)

(LSTP min−1)

(kg m−2s−1)

(°C m−1)

SMR, ref SE-SMR SE-SMR SE-SMR SMR, ref SMR, ref SE-SMR SE-SMR SE-SMR SMR, ref SE-SMR SE-SMR SMR, ref

550 550 550 550 550 600 600 600 600 600 600 600 600

3.0 3.0 5.0 7.0 3.0 3.0 3.0 5.0 5.0 3.0 3.0 7.0 3.0

0.3 0.3 0.5 0.7 0.3 0.3 0.3 0.5 0.5 0.3 0.3 0.7 0.3

0 0.30 0.30 0.30 0 0 0.30 0.30 0.20 0 0.20 0.20 0

Table 2. Parameters Used in the SE-SMR Simulation study kinetic and structural parameters of fresh (1× calcined) sorbent used in the studya parameter

608.7 608.7 629.1

units

value

m4 mol−1 s−1 3.6 × 10−9 ks0 −1 Ea J mol 18000 S0 m2 m−3 6.3 × 106 L0 m m−3 2.29 × 1013 ε0 0.4847 ψ0 3.75 ρp kg m−3 1555 dpore nm 759.0 u ̅p m s−1 0.0905 w 0.9371 kinetic and adsorption constants of the Pt/Ni/Al2O3 plate-type catalyst used in the studyb

645.6 663.1 728.2 679.2 679.2

a

P = 4.0. bar, (FH2O/FCH4)0 = 4.00, (FH2/FCH4)0 = 0.15, (FN2/FCH4)0 = 0.10.

constant

unit

k1 = 1.00 × 1015 exp [−250641 J mol−1/(RT)] k3 = 2.40 × 1015 exp[−267282 J mol−1/(RT)] KCH4 = 2.8 × 10−9 exp[43930 J mol−1/(RT)] KH2O = 1.23 × 106 exp[−99878 J mol−1/(RT)] KCOc = 8.23 × 10−10 exp[70650 J mol−1/(RT)] KH2c = 6.12 × 10−14 exp [82900 J mol−1/(RT)] reaction equilibrium constants

Steady-state operation of the SMR process was established by continuously monitoring the reactor exit stream for CO2. The composition of the exit stream was then analyzed by a GC (Agilent 7890A). After a few minutes of steady SMR operation, the flow of solid sorbent particles was started: it consequently reduced the level of CO2 in the reactor exit stream. The composition of the steady-state SE-SMR process, which took approximately 10 min to establish after the introduction of CaO into reactor, was also resolved by GC analysis. Reference SMR tests were done in order to determine the catalyst activity/ stability with time on stream. Model Formulation. A model of the SE-SMR process, which takes place in the gas−solid trickle flow reactor explained above, is needed for the simulation and optimization purpose. In setting up a pseudohomogeneous steady state model, the following assumptions were made • Plug flow for both phases. It is known that static mixers induce plug flow even for low gas-phase Reynolds number.10 Because of the used redistributors and coarse solid sorbent particles with low mass flux, axial dispersion in the solid-phase is also minimized.9,11 • Reactor operates with negligible pressure gradients in the axial direction.9 • Reactor operates isothermally in the parts where the external heating is applied, which was also confirmed experimentally. Because the base of the structured packing is made of metal (Ni) and is in contact with the reactor tube walls, radial temperature gradients are negligible. • No mass transfer limitations to/from the catalyst surface at the operating conditions used (Table 1), as

mol Pa1/2 m−2 s−1 mol Pa1/2 m−2 s−1 Pa−1 Pa−1 Pa−1

K1 = 1.198 × 1023 exp [−26830 K/T] K2 = 1.767 × 10−2 exp [4400 K/T] K3 = 2.117 × 1021 exp [−22430 K/T] reactor and catalyst geometrical parameters

Pa2 Pa2

parameter

unit

value

Drd acatd εBd catalyst element heightd redistributor height

m m2 m−3

0.043 523.86 0.83 0.042 0.01

m m

a

Obradović and Levec.9 bObradović et al.7 cKCO and KH2 are adopted from the work of Xu and Froment.13 dProvided by Sulzer Chemtech Ltd.

• Particles are of spherical shape and uniform in size (dp = 605 μm). • Mass transfer rate from the gas to the sorbent particle surface is estimated by the Ranz−Marshall correlation. • The dynamic hold-up of solid-phase is constant along the reactor height. The continuity equation for gaseous species j consumed or formed by catalytic reactions within the differential element of reactor can be expressed as dFj dz D

−

Dr 2π acatrj = 0 4

(5) DOI: 10.1021/acs.iecr.7b01832 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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where L2 is the total length of pores per unit volume and ε2 porosity of the 3× calcined sorbent. L2, S2, ψ2, and ε2 can be determined from the fresh (1× calcined) sorbent parameters L0, S0, ψ0, and ε0. The latter ones can be determined from the mercury porosimetry data as

where acat is the geometrical surface area of the catalyst per unit volume of reactor. The following boundary conditions apply at the reactor (bottom) entrance Fj = Fj,0 ,

Fj = 0,

j = CH4 , H 2 , H 2O, N2

j = CO, CO2

z=0

(6)

z=0

(7)

L0 =

Net reaction rate of a component j is related to the rates per reaction i (eq 1−4) as rCH4 = −(r1 + r3)

(8)

rH2O = −(r1 + 2r3)

(9)

rH2 = 3r1 + 4r3 rCO2

S0 = 2

ε0 =

(10)

1 = r3 − r4 acat

Local rate expressions for catalytic reforming reactions 1 and 3 (per geometrical catalyst surface area, in units mol m−2 s−1) are given in the form proposed by Xu and Froment13

SN = S0

2

+ K H2OpH O /pH 2

(13)

4

βdyn = (14)

2

The constants in the above equations for the catalyst used in this study are provided in the work of Obradović et al.,7 and are summarized in Table 2. The rate of carbonation reaction (in units mol m−3 s−1) can be described by the random pore model.15 In the kinetically controlled regime, this model reduces to

+ 0.52N 1

1 (1 − 0.075)

+ 0.52N

+ 0.075 (21)

+ 0.075 (22)

ṁ s ρp u p̅

(23)

(24)

z=L

(25)

Knowing the weight fraction of CaO in the sorbent w, and the sorbent particle density ρp (Table 2), CCaO,0 can be calculated. The second term in eq 24 represents the transport of CO2 from the bulk gas-phase to the sorbent surface, and the unknown concentration of CO2 at the particle surface can be calculated from eq 26. In the steady-state, CO2 transport from the bulk gas-phase is equal to CO2 disappearance rate in the sorbent, therefore the following relation applies

( ) CCaO CCaO,0

1 − ε2

⎛ pCO ⎞ kga⎜ 2 − CCO2, i⎟ ⎝ RT ⎠ ksS2CCaO 1 − ψ2 ln = βdyn(CCO2, i − CCO2,eq)

4πL 2(1 − ε2) S2

(20)

1 1 (1 − 0.075)

CCaO = CCaO,0

(16)

2

ν0(r ) dr

with the boundary condition

where CCO2,i and CCO2,eq stand for the concentration of CO2 at the sorbent particle surface and the equilibrium concentration of CO2, respectively. βdyn is the dynamic hold-up of solid sorbent in the reactor column. S2 represents the surface area per unit volume of a three times calcined sorbent used in this study and ψ2 structural sorbent parameter, defined as ψ2 =

(19)

⎛ pCO ⎞ ṁ s dCCaO − kga⎜ 2 − CCO2, i⎟ = 0 ⎝ RT ⎠ ρp dz

2

r4 = βdyn(CCO2, i − CCO2,eq)

ν0(r ) dr r

where ṁ s is the solid-phase flow rate per unit of cross sectional area of the empty reactor column, up̅ is the average particle velocity in the reactor column, and ρp is the particle density (Table 2). The mass balance for the solid sorbent can be written as

The WGS reaction is assumed to be in equilibrium at the reforming conditions,7,14 and the partial pressure of CO can be calculated from the reaction equilibrium constant pH pCO 2 2 K2 = pCO pH O (15)

ksS2CCaO 1 − ψ2 ln

(18)

where N corresponds to the number of cycles a sorbent underwent (N = 0 for the fresh sorbent). The assumption is made that the total sorbent pore volume does not change with the calcinations exerted on the sorbent (εN = ε0). βdyn can be calculated from

where DEN = 1 + K COpCO + K H2pH + K CH4pCH

∞

dr

∞

LN = L 0

(12)

⎛ pH 4 pCO ⎞ k3 ⎜ 2 2 2⎟ p p /DEN2 r3 = − K3 ⎟⎠ pH 3.5 ⎜⎝ CH4 H2O 2

πr 2

∫0

∫0

ν0(r )

with ν0(r) denoting the pore radii distribution of the fresh sorbent. For a sorbent that underwent a multicycle carbonation process, the sorbent structural parameters described above can be calculated using the structural parameters of the fresh sorbent and the empirical relation proposed by Grasa et al.16

(11)

⎛ pH 3 pCO ⎞ k1 ⎜ 2 ⎟ /DEN2 r1 = − p p K1 ⎟⎠ pH 2.5 ⎜⎝ CH4 H2O 2

∞

∫0

(17)

CCaO CCaO,0

(1 − ε2) (26)

E


Article

Industrial & Engineering Chemistry Research Mass transfer coefficient, kg, for the operating reactor conditions can be estimated from the Ranz−Marshall correlation, which for a single particle in an undisturbed gas flow reads17 k g dp

Sh =

Dm

1/3 = 2.0 + 0.6Re1/2 p Sc

(27)

Reynolds number in eq 27 is based on the average local particle and gas velocities proposed by Kiel et al.18 Rep =

ρg d p u p̅ − μg

ug εb

(28)

Figure 3. Experimental vs simulation yields of H2 and CO2 in SE-SMR steady state operation at 550 °C, P = 4.0. bar, (FH2O/FCH4)0 = 4.00, (FH2/FCH4)0 = 0.15, (FN2/FCH4)0 = 0.10.

Superficial gas velocity changes along the reactor axis due to the change in molar flow rate of the reaction mixture, and can be evaluated by

ug =

FtotRT P(Dr 2π /4)

At 600 °C, however, the reference SMR tests showed that the catalyst deactivation was up to 33% of the initial catalyst activity. The deactivation was accounted for by the empirical function proposed by Obradović et al.,7 and consequently acat was adjusted (the reaction zone height remained the same). Experimental vs calculated yields of H2 and CO2 at 600 °C are depicted in Figure 4. It can be observed that at both temperatures, the SE-SMR model described well the trends, but somewhat overpredicted the experimentally obtained yields. This can be due to the following reasons: • It was assumed that all the catalytic elements were identical, and also identical to the catalyst element which was used in the SMR kinetics determination.7 The same procedure, as described in Obradović et al.6 for a single structured element, was followed for catalytic activation of the surface (Figure 2c) of nine structured elements (Figure 2a,b) used in this study, with the difference that all the structured elements were pretreated simultaneously. This simultaneous pretreatment could have resulted in possibly different surface morphologies for the structured elements due to a different flow pattern in the furnace around each of the elements • The kinetic and adsorption parameters for SMR,7 kinetic parameters for carbonation reaction,9 and solid sorbent hydrodynamic parameters9 were determined in separate experimental studies, and therefore a better model estimation could not be expected. It should be mentioned that the reaction rate constants for SMR, k1 and k3, were taken from the lower limit of the confidence region, whereas all the other parameters were the optimal ones determined in the separate studies (Table 2)7,9 • Deviation from the plug flow behavior of the gas and solid-phase along the reactor column, and the possibility of an axial dispersion for both phases. Larger discrepancy between the model and experiment was expectedly found at higher level of catalyst deactivation (Figure 4b). The decrease of the catalyst activity at 600 °C can be explained by the sorbent dust blockage of the catalytic active sites. The question arises why the sorbent experienced obviously increased fragmentation at 600 °C in the SE-SMR process compared to that at 550 °C. For the reactor systems with sorbents, two attrition phenomena are observed: primary and secondary fragmentation. Primary fragmentation occurs after the injection of the particles in the bed, as a consequence

(29)

The gas−solid contact area for sorbent particles was determined as

a=

6βdyn dp

(30)

The molecular diffusion coefficient of CO2 in the multicomponent gas mixture, Dm, is calculated according to the relation proposed by Fairbanks and Wilke,19 in which the binary diffusion coefficients of CO2 with respect to each component of the mixture are calculated by the Chapman Enskog theory.20 The viscosity of the gas mixture can be predicted by the method of Wilke.20 After the catalyst stabilization procedure was performed, failure of one of the heaters at the bottom of the reactor column occurred, as shown in Figure 2. According to the information obtained at the thermocouple locations TC1, TC2, and TC3, a linear axial temperature gradient was assumed from the entrance of the catalyst layer to the TC3 location, at a distance of 0.112 m (Figure 2) dT = gradT dz

(31)

The estimated temperature gradients are summarized in Table 1 for different operating conditions. To obtain the outlet molar flow rates of the gaseous components in the reaction mixture, a system of ordinary differential equations described by eq 5 and eq 24 (and eq 31 for the bottom of the reactor) was solved. This was done using the subroutine ode15s for the system of ordinary stiff differential equations in Matlab R2014a software (Mathworks, Natick, Massachusetts, USA). The parameter values used during the simulations are summarized in Table 2. In the parts of the reactor with redistributors, no catalytic reactions were taken into account, but only the carbonation on the sorbent. The concentration of the sorbent at the inlet of the reactor was guessed, and this guess was checked at the outlet of the reactor, where the initial concentration of the sorbent is known. Model vs Experimental Measurements. Figure 3 shows the yields of H2 and CO2 at steady state SE-SMR operation at 550 °C determined experimentally and predicted by the model. There was no catalyst deactivation observed at 550 °C. F


Article


Figure 4. Experimental vs simulation yields of H2 and CO2 in SE-SMR steady state operation at 600 °C, P = 4.0 bar, (FH2O/FCH4)0 = 4.00, (FH2/ FCH4)0 = 0.15, (FN2/FCH4)0 = 0.10: (a) ṁ s = 0.30 kg m−2 s−1, (b) ṁ s = 0.20 kg m−2 s−1.

Figure 5. (a) Multistage and (b) one stage gas−solid trickle bed SE-SMR concept for high purity H2 outlet gas stream. P = 1.2 bar, T = 550 °C, FCH4,0 = 2.1 × 10−4 mol s−1, ṁ s = 0.30 kg m−2 s−1, (FH2O/FCH4)0 = 6.00, (FH2/FCH4)0 = 0.10.

of thermal stresses due to rapid heating, while secondary fragmentation is related to the mechanical stresses due to collisions with other particles or with the internals.21,22 The

reason, therefore, probably lies in the higher degree of thermal stress that sorbent particles underwent on their way from the bin, which was kept at the room temperature, through the G



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preheating section at 600 °C to the reaction zone. Some recent research on the mechanical stability of sorbent during the calcium looping process showed that increased attrition resistance could be achieved using calcium aluminate as a binder,23 applying thermal pretreatment or chemical doping,24 or incorporating CaO on an attrition resistant carrier.25 Multistage Gas−Solid Trickle Bed Concept for High Purity Hydrogen. Additional simulation, with the same parameter values as used in the experimental validation part discussed above, was done in order to introduce a multistage gas−solid trickle bed SE-SMR concept that could provide an outlet gas stream with the desirable purity of hydrogen for low temperature SE-SMR operation. The operating conditions chosen for the simulation were the following ones: P = 1.2. bar, T = 550 °C, FCH4,0 = 2.1 × 10−4 mol s−1, ṁ s = 0.30 kg m−2 s−1 (FH2O/FCH4)0 = 6.00, (FH2/FCH4)0 = 0.10, L = 8 m (see Table 2 for parameter values). The high inlet steam to methane ratio and low pressures are favorable for SE-SMR operation.26 However, the following multistage concept can be applied at less favorable operating SE-SMR conditions, as well. The idea of the multistage concept is to have successive catalytic and inert packing elements in the reactor column, as shown in Figure 5a. In that way, CO2 and CO percentages are lower at the reactor outlet compared to the case where there is no inert packing (Figure 5b). Namely, in the parts of the SE-SMR reactor with the catalyst, the WGS reaction is assumed to be in equilibrium instantaneously. When CO2 is taken from the reactor system with the sorbent in the reactor parts with the inert packing, a new WGS equilibrium state is reestablished at the entrance of the next stage of the SE-SMR reactor with the catalyst. To establish this new equilibrium WGS state, more CO is converted into the product side in the WGS. Since CO is produced in the reforming reaction (eq 1), and CO2 also in the second reforming reaction (eq 3), these two equilibrium limited reactions (but not in equilibrium at SE-SMR reaction conditions) undergo enhancement in net rate, as well. The calculated molar percentages on a dry basis at the reactor outlet for the multistage SE-SMR concept (Figure 5a) were the following ones: 97.61% H2, 2.00% CH4, 0.29% CO2, and 0.10% CO, whereas for the one-stage SE-SMR concept the composition at the reactor outlet was 86.25% H2, 2.89% CH4, 9.19% CO2, and 1.67% CO (Figure 5b). In both the cases the reactor zone was 8 m long (Figure 5), but with the multistage concept, only 5 m of the reactor zone was catalytically active (Figure 5a). The calculated sorbent conversion per reactor pass in the multistage SE-SMR concept was 3.2%, which indicates that the sorbent could be reused many times prior to its regeneration. It should be noted that higher pressures and lower steam to methane ratios usually applied in the industry would require a considerably higher reactor column to obtain high purity hydrogen in the outlet stream. A somewhat similar multistage concept was introduced by Lee et al. for SE-WGS27 and SE-SMR28 processes, but in a fixed-bed reactor, where the reactor column was divided into two sections with different catalyst to sorbent ratios. However, with the concept introduced in this study, a high purity hydrogen stream is produced in a continuous way regarding the sorbent.

CONCLUSIONS It was experimentally demonstrated that the sorption enhanced-steam methane reforming (SE-SMR) process could be efficiently carried out in a continuous mode of operation using countercurrent gas−solid trickle flow reactor. The reactor was packed with special static mixers made of corrugated Pt/ Ni/Al2O3 plates, catalytically active for SMR reactions. These static mixers acted as distributors for solid sorbent at the same time. Simultaneous sorption of carbon dioxide increases hydrogen concentration in the exit stream, but due to the thermodynamic constrains in the reaction mixture it reaches an equilibrium value. Simulations showed that this could be overcome using an alternating catalytic and inert elements packing. Depending on the length of catalytic and inert alternating packing it is possible to control the level of hydrogen purity in the reactor exit stream. Also, with the alternating packing the total height of the catalytic part needed to achieve a certain level of CH4 conversion is lower compared to that of the SE-SMR case with catalytic elements only. This kind of multistage trickle bed concept can be applied to SEWGS reaction systems, as well as to the other types of equilibrium limited systems in which a sorbent could be applied. However, for the industrialization of the process, the mechanical stability of sorbent should be further addressed using binders, applying thermal pretreatment or chemical doping, incorporating sorbent on an attrition resistant carrier, or some other technique.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. ORCID

Ana Obradović: 0000-0003-4492-1576 Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors would like to thank the Slovenian Research Agency (ARRS) (Program P2-0152) for financing this work, Sulzer Chemtech Ltd, Switzerland, for their cooperation and provision of the structured elements, and Matjaž Debeljak, the director of former Solkanska Industrija Apna, Nova Gorica, Slovenia, for the donation of CaO material for this research. We are also thankful to our colleague Dr. Gorazd Berčič for his useful technical suggestions.

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H

NOMENCLATURE a = specific interfacial area between gas and sorbent particles (m2 m−3) acat = geometrical catalyst surface area per unit volume of reactor (m2 m−3) CCaO = sorbent concentration per sorbent particle (mol m−3) Cj = concentration of species j in gas-phase (mol m−3) CCO2, i = concentration of CO2 at sorbent particle surface (mol m−3) CCO2, eq = equilibrium concentration of CO2 (mol m−3) Dm = molecular diffusion coefficient of CO2 in multicomponent gas mixture (m s−1) dp = sorbent particle diameter (m) Dr = reactor diameter (m) Ea = activation energy of carbonation reaction (J mol−1) Fj = molar flow rate of component j (mol s−1) DOI: 10.1021/acs.iecr.7b01832 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Article

Industrial & Engineering Chemistry Research Ftot = total molar flow rate (mol s−1) ΔrH298 ° = standard reaction enthalpy (kJ mol−1) kg = gas−solid mass transfer coefficient (m s−1) ks = rate constant of carbonation surface reaction (m4 mol−1 s−1) ks0 = pre-exponential factor of rate constant ks (m4 mol−1 s−1) k1, k3 = rate coefficient of reforming reactions 1 and 3 (mol Pa1/2 m−2 s−1) K1, K3 = equilibrium constants of reforming reactions 1 and 3 (Pa2) K2 = equilibrium constant of WGS reaction (/) KCH4, KCO, KH2 = adsorption equilibrium constant for CH4, CO, and H2 (Pa−1) KH2O = dissociative adsorption constant of H2O L = total length of structured catalyst bed (m) L0 = total length of pores per unit volume (m m−3) LN = total length of pores per unit volume of (N+1) times calcined sorbent (m m−3) ṁ s = mass flux of solid sorbent (kg m−2 s−1) P = total pressure (bar, Pa) pj = partial pressure of j species in gas phase (bar, Pa) pCO2,i = partial pressure of CO2 at sorbent particle surface (bar, Pa) r1, r3 = rates of reaction 1 and 3 per geometrical surface area of the catalyst (mol m−2 s−1) r4 = rate of reaction 4 (mol m−3 s−1) rj = net reaction rate of component j per geometrical surface area of the catalyst (mol m−2 s−1) Q = volumetric flow rate (m3 s−1, L min−1) S0 = surface area per unit volume of fresh (1 time calcined sorbent) (m2 m−3) SN = surface area per unit volume of (N + 1) times calcined sorbent (m2 m−3) Scat = geometrical surface area of catalyst (m2) ug = superficial gas-phase velocity (m s−1) u̅p = mean particle velocity (m s−1) VB = volume of empty reactor column (m3) w = weight fraction of CaO in the calcined sorbent (/) yj = molar fraction of species j in gas phase

(6) Obradović, A.; Likozar, B.; Levec, J. Catalytic Surface Development of Novel Nickel Plate Catalyst with Combined Thermally Annealed Platinum and Alumina Coatings for Steam Methane Reforming. Int. J. Hydrogen Energy 2013, 38 (3), 1419. (7) Obradović, A.; Likozar, B.; Levec, J. Steam Methane Reforming over Ni-Based Pellet-Type and Pt/Ni/Al2O3 Structured Plate-Type Catalyst: Intrinsic Kinetics Study. Ind. Eng. Chem. Res. 2013, 52 (38), 13597. (8) Verver, A. B.; Van Swaaij, W. P. The Hydrodynamic Behaviour of Gas−Solid Trickle Flow over a Regularly Stacked Packing. Powder Technol. 1986, 45 (2), 119. (9) Obradović, A.; Levec, J. Countercurrent Gas−Solid Trickle Flow Reactor with Structured Packing: Hydrodynamics and CaO−CO2 Reaction. AIChE J. 2015, 61 (5), 1601. (10) Thakur, R. K.; Vial, C.; Nigam, K. D.; Nauman, E. B.; Djelveh, G. Static Mixers in the Process IndustriesA Review. Chem. Eng. Res. Des. 2003, 81 (7), 787. (11) Nikačević, N. M.; Duduković, A. P. Solids Residence Time Distribution in Gas-Flowing Solids-Fixed Bed Contactors. Ind. Eng. Chem. Res. 2005, 44 (16), 6509. (12) Alvarez, D.; Abanades, J. C. Determination of the Critical Product Layer Thickness in the Reaction of CaO with CO2. Ind. Eng. Chem. Res. 2005, 44 (15), 5608. (13) Xu, J.; Froment, G. F. Methane Steam Reforming, Methanation and Water-Gas Shift: I. Intrinsic Kinetics. AIChE J. 1989, 35 (1), 88. (14) Jakobsen, J. G.; Jørgensen, T. L.; Chorkendorff, I.; Sehested, J. Steam and CO2 Reforming of Methane over a Ru/ZrO2 Catalyst. Appl. Catal., A 2010, 377 (1), 158. (15) Bhatia, S. K.; Perlmutter, D. D. A Random Pore Model for Fluid−Solid Reactions: I. Isothermal, Kinetic Control. AIChE J. 1980, 26 (3), 379. (16) Grasa, G. S.; Abanades, J. C. CO2 Capture Capacity of CaO in Long Series of Carbonation/Calcination Cycles. Ind. Eng. Chem. Res. 2006, 45 (26), 8846. (17) Ranz, W. E.; Marshall, W. R. Evaporation from Drops. Chem. Eng. Prog. 1952, 48 (3), 141446. (18) Kiel, J. H.; Prins, W.; Van Swaaij, W. P. Modelling of NonCatalytic Reactions in a Gas-Solid Trickle Flow Reactor: Dry, Regenerative Flue Gas Desulphurisation Using a Silica-Supported Copper Oxide Sorbent. Chem. Eng. Sci. 1992, 47, 4271. (19) Fairbanks, D. F.; Wilke, C. R. Diffusion Coefficients in Multicomponent Gas Mixtures. Ind. Eng. Chem. 1950, 42 (3), 471. (20) Reid, R. C.; Prausnitz, J. M.; Poling, B. E. The Properties of Gases and Liquids, 4th ed.; McGraw-Hill, Inc.: New York, 1987. (21) Scala, F.; Montagnaro, F.; Salatino, P. Attrition of Limestone by Impact Loading in Fluidized Beds. Energy Fuels 2007, 21 (5), 2566. (22) Scala, F.; Salatino, P.; Boerefijn, R.; Ghadiri, M. Attrition of Sorbents during Fluidized Bed Calcination and Sulphation. Powder Technol. 2000, 107 (1), 153. (23) Manović, V.; Anthony, E. J. Long-Term Behavior of CaO-Based Pellets Supported by Calcium Aluminate Cements in a Long Series of CO2 Capture Cycles. Ind. Eng. Chem. Res. 2009, 48 (19), 8906. (24) Blamey, J.; Anthony, E. J.; Wang, J.; Fennell, P. S. The Calcium Looping Cycle for Large-Scale CO2 Capture. Prog. Energy Combust. Sci. 2010, 36 (2), 260. (25) Hu, Y.; Liu, W.; Sun, J.; Li, M.; Yang, X.; Zhang, Y.; Xu, M. Incorporation of CaO into Novel Nd2O3 Inert Solid Support for High Temperature CO2 Capture. Chem. Eng. J. 2015, 273, 333. (26) Halabi, M. H.; De Croon, M. H.; Van der Schaaf, J.; Cobden, P. D.; Schouten, J. C. Reactor Modeling of Sorption-Enhanced Autothermal Reforming of Methane. Part II: Effect of Operational Parameters. Chem. Eng. J. 2011, 168 (2), 883. (27) Jang, H. M.; Kang, W. R.; Lee, K. B. Sorption-Enhanced Water Gas Shift Reaction using Multi-Section Column for High-Purity Hydrogen Production. Int. J. Hydrogen Energy 2013, 38 (14), 6065. (28) Lee, C. H.; Mun, S.; Lee, K. B. Application of Multisection Packing Concept to Sorption-Enhanced Steam Methane Reforming Reaction for High-Purity Hydrogen Production. J. Power Sources 2015, 281, 158.

Greek Letters

βdyn = Vsorb/VB εB = Vvoid/VB ε0 ρp ρg μg ψ

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dynamic holdup of solid sorbent (-) porosity of structural packing (-) porosity of fresh sorbent (-) sorbent particle density (kg m−3) gas density (kg m−3) viscosity of gas-phase (kg m−1 s−1) structural parameter of solid sorbent (-)

REFERENCES

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High Purity Hydrogen with Sorption-Enhanced Steam Methane

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