Solvent Recovery by Steamless Temperature Swing Carbon

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Solvent Recovery by Steamless Temperature Swing Carbon Adsorption Processes Pradeep K. Sharma and Phillip C. Wankat* School of Chemical Engineering, Purdue UniVersity, Forney Hall of Chemical Engineering, 480 Stadium Mall DriVe, West Lafayette, Indiana 47907-1283, United States

Adsorption with activated carbon followed by regeneration with steam is the most common technique for solvent recovery from gas streams. The resulting steam-solvent gaseous mixture when condensed can result in either two separate layers (immiscible) or a single liquid phase (miscible or miscible with an azeotrope). For miscible systems that form azeotropes, the downstream distillation system can easily be more expensive than the adsorbers. An alternative is to use hot nitrogen instead of steam for desorption. By adding a short oxygen/water removal step to the adsorption cycle, the need for distillation is avoided. The model system studied for the azeotropic case was recovery of 0.5 mol % isopropanol (IPA) from air with hot nitrogen for regeneration. Since the nitrogen cost dominates, nitrogen should be recovered and recycled as much as possible. Two- and four-bed adsorption schemes were studied, and the four-bed system, which recovers almost all of the nitrogen, was best. A preliminary total cost comparison showed that solvent recovery of IPA with steam regeneration has total costs that are ∼14% higher than for IPA recovery with hot nitrogen regeneration if the nitrogen is recovered and recycled. Introduction Organic solvents like isopropyl alcohol, acetone, ethanol, ethyl acetate, and toluene are extensively used in operations like painting, coating, and gluing.1,2 These volatile organic compounds (VOCs) often lead to environmental issues such as ozone depletion and smog.3 Solvents are the source of about 40% of the VOCs entering the atmosphere.1 Most VOCs are flammable in air, and hence their vapors can pose serious fire hazards. Consequently, they are either processed in oxygenfree environments, or excess air is pumped in to confine their vapor concentrations to well below the lower explosion limit (LEL).4 Further, some VOCs are carcinogenic, and others harm the kidneys and liver.5 Because of these issues, the Occupational Safety and Health Administration (OSHA) and the Environmental Protection Agency (EPA) regulate the emission of VOCs into the atmosphere. Fortunately, the recovery and reuse of solvents can help reduce fresh solvent requirements and hence completely or partially compensate for VOC abatement costs.1,6 Adsorption-based processes are commonly used to recover volatile solvents from gas streams.1,3,7,8 Pressure Swing Adsorption9 (PSA) and Temperature Swing Adsorption10 (TSA) are common processes for the separation/ purification of gas streams. PSA uses variation in column pressure, while TSA uses variation in column temperature to regenerate the adsorbent. Hybrid processes coupling PSA and TSA together for process improvement are also common.11–13 Solvent vapor capture and recovery from gas streams like air can be achieved by either PSA or TSA or both.1,6,8,14,15 The standard adsorption-based solvent recovery process uses activated carbon (adsorbent) to adsorb solvent vapor, and steam is generally used to desorb solvent (Figure 1). This is a special type of TSA process. If the feed is a solvent-air mixture, the solvent vapor concentration is below its LEL. During the regeneration step, steam immediately pushes out the air, removing oxygen from the column. Another advantage of using steam is its high energy density that allows rapid heating of the adsorber. The solvent-steam gas mixture coming out of the * To whom correspondence should be addressed. Tel.: 765-494-0814. Fax: 765-494-0805. E-mail: [email protected].

column is cooled to condense the steam and solvent. If the solvent is immiscible with water, then the condensate separates into two liquid layers. A simple decantation separates solvent from water. The solvent, highly pure, can now be reused. However, if the solvent and water are miscible, an additional separation, usually distillation, is required to recover the solvent. Further, if the solvent forms an azeotrope with water, extractive or azeotropic distillation will probably be required, which will increase the cost further. Hence, regeneration with steam is an excellent option when solvent is immiscible with water but may not be otherwise. Alternatively, hot, inert, nonadsorbing gas can be used instead of steam (e.g., N2 or CO2).16–19 Hot clean air cannot be used with flammable solvents because it contains O2. Increases in solvent concentration during thermal regeneration and the presence of O2 pose an explosion hazard. Use of an inert gas allows preferential solvent condensation and its recovery as pure liquid, while the uncondensed residual gas can be recycled back for readsorption. The two main operating costs are thermal energy (includes both heating and refrigeration) and the makeup requirements of fresh inert gas. The current work deals with the minimization of the combined operating cost of solvent recovery systems that are regenerated with a hot inert gas. A number of processes with hot N2 desorption were developed. The simplest process is shown in Figure 2. The feed is assumed to contain 5000 ppmv IPA () 1/4 LEL) which is a safe concentration.1 After the feed step, O2 and H2O are removed from the column using a short counter current purge with cold N2. However, hot N2 could be used as well. Removal of O2 before thermal regeneration is a measure to avoid explosion hazards as IPA concentration increases during thermal regeneration. Prior removal of H2O ensures high purity of the condensate in IPA during the thermal regeneration step. Hot nitrogen is fed counter-current to the feed direction to desorb IPA. This step is followed by cold nitrogen in the same direction which cools the bed and pushes the hot gas further toward the feed-end. The O2- and H2O-free IPA-N2 stream leaving the adsorbent column enters a condenser where most of the IPA condenses out (operating temperature of the condenser is T ) 4 °C), while the residual saturated gaseous

10.1021/ie1008019  2010 American Chemical Society Published on Web 10/26/2010

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Figure 1. Standard carbon adsorption process (modified from US Patent 4,440,549).

Figure 2. Simple two-bed TSA process.

mixture is recycled back for readsorption. The gas stream flowing into the condenser is heat exchanged with the residual saturated gas stream flowing out of the condenser. Such thermal exchange is applied to all the schemes in this work. In Figure 2, all the fresh pure nitrogen used in regeneration appears in the recycle to the adsorber and is ultimately discarded to the atmosphere. Hence, a large amount of fresh nitrogen is used (and lost to ambience) and is a significant cost. One way to reduce pure nitrogen intake is the partial use of the uncondensed gas stream from the condenser to regenerate the adsorber. This gas can be heated and used to offset the total requirement of hot fresh nitrogen (Figure 3). A commercially available process also recycles N2 in a similar fashion.20 However, this process still requires fresh nitrogen for cooling

the bed. The impure gas from condenser cannot be used for cooling (at low temperatures) because of the presence of significant amounts of IPA in it which can adsorb. Nitrogen use can be further reduced by incorporating a second TSA unit. The IPA-N2 stream emerging during the regeneration of the first TSA is fed to a second TSA as shown in Figure 4. The second TSA unit removes IPA from the IPA-N2 mixture. Hence, the clean N2 stream emerging from the second TSA can be used by the first TSA as well as the second TSA for both thermal regeneration and cooling purposes. In this two-stage process, a small amount of N2 containing less than 100 ppm IPA is vented to the atmosphere in the “O2 and H2O removal” step. Hence, a small makeup of N2 is still required. This is the only fresh N2 required in this adsorption system. Also, the top

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Figure 3. Two-bed TSA with partial N2 reuse.

Figure 4. Four-bed, two-stage TSA.

stream from the second TSA can be heat exchanged with the top stream from the first TSA to reduce energy requirements. Theory and Modeling The governing equations are mass and heat balances, massand heat-transfer rates, and adsorption isotherms in a packed bed as provided in Table 1. The model system studied in this work is the adsorption of isopropyl alcohol (IPA) on activated carbon from a feed containing IPA vapor, water vapor, and air.21 IPA is a common solvent and forms an azeotrope with water. It is flammable in air, and hence its concentration in air is kept below the LEL (LELIPA ) 20 000 ppmv).

According to eq 1f (Table 1), a nonadsorbing species diffuses very fast, and hence very large mass transfer coefficient values were set for air and water vapor (H2O adsorption has been suppressed as discussed later). For IPA, the mass transfer coefficient was estimated using eqs 1e through 1h. Table 2 summarizes the appropriate initial and boundary conditions for the mass and energy balance for fixed bed adsorptive processes studied in this work. Table 3 lists important parameter values for simulating IPA adsorption-desorption on activated carbon. The adsorption beds were assumed to be initially filled with clean air. Although this is purely a simulation study, the adsorption isotherms and mass transfer coefficients were taken from experimental literature. For distillation, the NRTL method

Ind. Eng. Chem. Res., Vol. 49, No. 22, 2010 Table 1. Mass and Energy Balance, Mass and Heat Transfer Rates, and Adsorption Isotherm

Table 2. Initial and Boundary Conditions for Mass and Energy Balance

Mass and Energy Balances22 εe

Initial Conditions27

∂ci ∂cjpore,i ∂(Vfci) ∂qji + Kd,i(1 - εe)εp + Fs(1 - εe)(1 - εp) + εe ∂t ∂t ∂t ∂z 2 ∂ ci εeEz 2 ) 0 (1a) ∂z

FfCp,fεe

js ∂T j* ∂T ∂T + FfCp,fεp(1 - εe) + FsCp,s(1 - εp)(1 - εe) + ∂t ∂t ∂t 2 ∂(νT) ∂T - Ez,TFfCp,fεe 2 ) 0 FfCp,fεe ∂z ∂z

yIPA(z,t)0) ) 0;

yH2O(z,t)0) ) 0;

yNz(z,t)0) ) 0.79;

Boundary Conditions27 Adsorption (Feed Step)

(1b)

∂ci ) 0; ∂z(z)L,t)

ci(z)0,t) ) cifeed ;

T(z)0,t)

ci(z)L,t) ) (1c)

cipurge ;

∂ci ) 0; ∂z(z)0,t)

T(z)L,t)

∂qjIPA ∂t

Pore Diffusion Controlled Mass Transfer Coefficient21 60Deff,i kMTC,solid,i ) d2p

(1d)

(1e)

Effective Diffusivity, Deff, is a Function of Surface (Ds) and Knudsen Diffusion (Dk)23 εp ∂ci Deff,i ) Ds,i + Dk,i (1f) Fp ∂qi



T Mi

m2 /s

Ds,i ) Ds0,i exp(-Ei /RT)

m2 /s

Heat-Transfer Coefficient, hHTC, from the Colburn j-Factor hHTC ) jCp,fνfFfPr-2/3

(1g)

(1h)

(1i)

where Pr ) µCp,f/kez depends on the thermal conductivity kez24,25 with26 j ) 0.983Re-0.41 for Re > 190 and j ) 1.66Re-0.51 for Re < 190 (1j) Cp,f ) yN2Cp,N2 + yO2Cp,O2 + yIPACp,IPA + ywaterCp,water, Cp values from Perry’s24 Adsorption Isotherm Since water vapor is kept below 50% RH, only IPA adsorbs21 m(yIPAP) q* (1k) IPA ) [b + (yIPAP)n]1/n

(

b ) b0 exp -

nδHads,IPA RT

)

(2b)

(1l)

has been suggested for the IPA/water mixture. Table 4 consists of the NRTL model parameters. The experimentally unavailable parameters were estimated from relevant correlations. Aspen Adsim was used to simultaneously solve differential mass and energy balances, mass- and energy-transfer equations,

∂νf ) 0; ∂z(z)0,t) ∂T ) Tpurge ; )0 ∂z(z)0,t)

νf(z)L,t) ) νpurge ;

(2c)

Table 3. Parameter Values for IPA Adsorption on Activated Carbon21,28 adsorbent

rpore τ

∂νf ) 0; ∂z(z)L,t) ∂T ) Tfeed ; )0 ∂z(z)L,t)

νf(z)0,t) ) νfeed ;

js ∂T j* ∂T + FfCp,fεp(1 - εe) ) ∂t ∂t

Dk,i ) 97

(2a)

Purge Step

Linear Driving Force (LDF) Approximations22 ∂qji ) kMTC,solid(q* j i) i - q ∂t

-hMTCRp(Ts - T) + (1 - εp)(1 - εe)Fs∆Hads,IPA

yO2(z,t)0) ) 0.21; T(z,t)0) ) Tinitial

Mass and Heat Transfer

FsCp,s(1 - εp)(1 - εe)

11605

bulk density, Fb solid heat capacity, Cp,s bed porosity, εe particle porosity, εp mean pore radius, rpore BET surface area solid density, Fs particle diameter, dp

activated carbon15

ap T Ds E

400 kg/m3 813 J/kg K 0.411 0.707 15.2 Å 1400 m2/g 2270 kg/m3 3.5 mm (lab scale), 1.4 mm (industrial scales) 60025.8 J/mol kez ) 0.0026 (300 K), 0.0034 (400 K), 0.0041 (500 K) 977 m-1 4 4 × 10-7 m2/s 27,011.6 J/mol

Isotherm Coefficients m b0 n lower explosion limit (LEL)1 upper explosion limit (UEL) water-IPA azeotrope29 OSHA permissible exposure limit30

7.457 mol/kg 1.2383 × 107 Pan 0.598 2 vol % 12 vol % 80.3-80.4 °C, 87.4-87.7 wt % 0.04 vol %

heat of adsorption, ∆H thermal conductivity, kez

and equilibrium relations. Aspen Adsim is a commercial simulator for adsorption processes and has been benchmarked previously against the standard Lapidus and Amundson solution under limiting conditions.33 Simulations of the adsorption processes presented in the current work include the following assumptions: (1) Oxygen and nitrogen do not adsorb.21 (2) Below 50% RH, water adsorption and capillary condensation are negligible.3,6,10,28,34 (3) No reactions other than adsorption take place. (4) No radial dispersion. (5) Axial dispersion caused by the second derivative terms in eqs 1a and 1b is negligiblesaxial spreading is caused by the finite rate of mass transfer. (Essentially the same results will be obtained whether axial dispersion is included or not. However, because including axial dispersion in ADSIM is easy, it was included.) (6) Momentum balance follows the Karman-Kozeny equation.

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Table 4. Distillation For the IPA-water system, the NRTL model has been suggested in the literature.31 The NRTL model connects the activity coefficient of a component to the mole fractions of the components present in the concerned liquid phase.

∑xτ G ln γ ) ∑x G j ji

ji

j

i

k

+

∑ j

ki

k

xjGij

∑x G k

k

[

τij

kj

∑x τ G ∑x G

Gij ) exp(-Rijτij) τij ) Rij + Rij ) cij,

bij + eij ln T T

τii ) 0,

Gii ) 1

m mj

mj

m

k

k

kj

]

(4a)

(4b) (4c) (4d)

In Aspen Plus, once the NRTL method is selected, Aspen Plus uses its own model parameters as a default setting. However, the following rigorous parameters are available in the literature. IPA ) i, water ) j aij ) 0, aji ) 0, bij (K) ) 185.4, bji (K) ) 777.3, cij ) 0.5, eij ) 0, eji ) 0. The derivation and mathematical formulation for the NRTL method can be found elsewhere.32

(7) Ideal gas behavior. (8) Linear lumped parameter driving force for mass and heat transfer. (9) Constant heat of adsorption. All the above assumptions can be relaxed if appropriate data are available. Aspen Adsim solves the governing partial differential equations using the method of lines.35 Space dimension was discretized using the first-order upwind differencing scheme (UDS1). Integration was carried out using the Implicit Euler method with a variable time step of 0.1-2 s. Time- and space-dependent variables like velocity, density, etc. were computed at each grid point and time step. Adsorption isotherm and mass transfer coefficients, obtained from experimental literature, were custom built into ADSIM. The parameters are in Table 3. For the nonadsorptive part like distillation columns, heat exchangers, etc., Aspen Plus was used using the parameters provided in Table 4. Counter-current flow shell and tube heat exchangers were used throughout this study. Results and Discussion Recovery was calculated using the following expression IPA Recovery (%) ) moles of IPA condensed in t hours × 100 moles of IPA present in feed air fed in t hours All the cyclic processes were run until the cyclic steady state (CSS) was achieved. The cyclic steady state is achieved when all the bed profiles (temperature, concentration, etc.) at the start of a cycle are similar to that at the end of that cycle. A relative error tolerance of 10-6 was used. As was mentioned earlier, IPA vapor is flammable. Its concentration in air is generally restricted to at or below 1/4 of its LEL value (LEL: 20 000 ppm), although one source reported 25-35% LEL as a safe range of operation.1 In this study, the more conservative limit was used, and the feed is 5000 ppm IPA in air (25% LEL). Air is assumed to be at 35 °C with 85% relative humidity. Water vapor has been reported to compete in adsorption onto activated carbon above 50% relative humidity. Rudisill et al. reported pure water isotherms onto activated carbon. Water adsorption at 40% RH was less than 10% (based

on water adsorption at 100% RH).36 Further, Huggahalli and Fair showed that at high humidity water competes with organic compounds. For example, acetone and water mixtures at low humidity conditions have essentially no water adsorption.34 Hence, feed air is heated to 50 °C where relative humidity becomes 39% before it is fed to the adsorber. Results for 5000 ppm Feed. Scheme 1: Simple TSA [Figure 2]. For a lab scale throughput of 21.3 SCF/h, two adsorbers with dimensions 12 cm (L) × 12.8 cm (D) were used. Addition of recycle made the IPA concentration increase to an average of 6832 ppm and a maximum value of ∼8000 ppm. This is an issue with this scheme since this concentration is above the recommended safe range (25-35% LEL) of operation (8000 ppm ) 40% LEL). The average IPA concentration in purified air was fixed at 100 ppm. The adsorption step was 3.8 h with a superficial feed velocity of 0.02 m/s at ∼1 bar. In other words, 3.8 h was the breakthrough time such that cycle average concentration of purified air released to atmosphere was 100 ppm IPA. However, because the bed was completely clean the adsorption step for the first cycle produced purer air (∼50 ppm). Incomplete regeneration and mixing of effluent during the H2O and O2 removal step yielded 100 ppmv IPA in purified air at CSS. For regeneration, hot nitrogen at 473 K flows at a superficial velocity of ∼0.01 m/s. It takes 3.8 h for the center of the thermal wave to reach to the other end. This gives an estimate of the amount of N2 required. This amount of N2 was arbitrarily fixed. After a short counter-current O2 and H2O removal step with cold N2 (3 min long, ∼0.2 m/s), hot counter current N2 was passed at 0.1 m/s followed by cold N2. Keeping the total amount of N2 purge fixed, the ratio of hot to cold N2 was varied. Hot N2 thermally desorbs IPA but leaves the column hot and unsuitable for the next adsorption cycle, while cold N2 prepares the bed for the next cycle but hardly desorbs IPA. The ratio of hot to cold N2 that produced 100 ppm product purity at a cyclic steady state with the least amount of hot nitrogen was 1.25. With two beds operating out of phase, 7.6 h was the total cycle time: adsorption + H2O/O2 removal + hot N2 regeneration + cold N2 cooling. Cyclic steady state was achieved in 5 to 10 cycles. The effluent from the H2O and O2 removal step could be sent for readsorption instead of releasing it with the purified air. However, the amount of this stream is very small, and it does not increase the exit air concentration above 100 ppm IPA. Figure 5(A) shows the cycle structure for scheme 1. Scheme 2: Simple TSA with Partial N2 Reuse [Figure 3]. For the same lab scale throughput of 21.3 SCF/h, two adsorbers with dimensions of 12 cm (L) × 10.8 cm (D) were used. The cycle time was again 7.6 h. Part of the hot N2 in the regeneration cycle was replaced by heated residue gas from the condenser (15 000 ppm IPA). All of the residue gas from the condenser cannot be used due to timing constraints. A complete cycle for this scheme is comprised of five steps: adsorption, O2, and H2O removal, hot pure N2, hot residual gas from condenser, and cold pure N2. Scheme 1 was used as the starting point for this scheme. As the recycle stream decreased, adsorber diameters were reduced so that an average purity of 100 ppm IPA and similar velocities were obtained. The peak IPA concentration that enters the adsorbers is ∼7000 ppm (35% LEL). This is due to decreased recycle flow. Average IPA concentration entering adsorbers is 5532 ppm. Figure 5(B) shows the cycle structure for scheme 2. Scheme 3: Four-Bed TSA-TSA [Figure 4]. This scheme is more complex than schemes 1 and 2. Unlike the previous two schemes, the fresh feed (5000 ppm IPA-air mixture) is introduced directly into the adsorbers (there is no mixing with

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Figure 5. (A) Time scheme: two-bed simple TSA. (B) Time scheme: simple TSA w/partial N2 reuse.

Figure 6. Time scheme: four-bed TSA-TSA.

a recycle prior to adsorption). Therefore, the TSA 1 adsorber is smaller for the same feed. However, the presence of the second TSA unit offsets this cost advantage. For simplicity, all beds are of the same size, 12 cm (L) × 9.52 cm (D). Adsorption was carried out at 323 K in TSA 1, while it was at 310 K in TSA 2. TSA 1 produced the required purity of 100 ppm IPA. TSA 2 is not required to produce 100 ppm IPA since the exiting

gas stream is not discarded to the atmosphere. It was observed from simulation results that a 105 ppm IPA exiting gas stream from TSA 2 could be used to maintain a 100 ppm IPA exiting gas stream from TSA 1. Figure 6 shows the cycle structure for scheme 3. Preliminary Comparison. All the above schemes are compared for the same feed conditions in Table 5. Scheme 1

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Table 5. Utility Comparison among Three Schemes for the Laboratory Scale Unita scheme 1

scheme 2

scheme 3

21.3 0.32 15.14 7.48 0.02 47 1278.38 96.83% 8000

21.3 0.23 11.51 7.24 0.015 22 580.14 97.52% 7000

21.3 0.36 22.3 7.27 0.019 1.01 27.2 97.79% 5000

feed rate adsorbent cooling water refrigeration steam (0 psig to 300 psig) nitrogen liquid IPA recovery maximum IPA concentration inside adsorber with O2 present (i.e., during adsorption)

SCF/h kg/(100 SCF feed h) kJ/(100 SCF feed h) kJ/(100 SCF feed h) kg/(100 SCF feed h) SCF/100 SCF feed SCF/kg liquid IPA ppm IPA

a Outlet air maximum average IPA concentration: 100 ppm. Inlet air conditions: 35 °C, 1.1 bar, 5000 ppm IPA. Standard Conditions: 0 °C, 1 atm: 1/4 LEL ) 5000ppm (mol basis).

Table 8. Scaling Equations22

Table 6. Utility and N2 Cost Chart Energy Cost37

∆Pold

steam

∆Pnew

pressure

price ($/1000 kg)

450 psig 150 psig 50 psig

14.5 10.5 6.6

RN )

[

)

( )( Qold Qnew

] ( )(

(L/LMTZ)new Qold ) (L/LMTZ)old Qnew

cooling water temperature

price ($/m3)

89.6 °F

0.02 Refrigeration

temperature

price ($/GJ)

40 °F 10 °F

4 5.5 N2 Costa

on-site production range (SCF/h)

price range ($/100 SCF)

25000

0.25-0.3 0.15-0.25 0.05-0.1

a

Courtesy: Praxair, Inc., NY, 2009.

uses the maximum amount of N2. If schemes 1 and 2 are compared, scheme 2 is a clear winner in all areas except simplicity. The biggest advantages are safety and in fresh N2 requirement. Scheme 2 uses less than 50% of the fresh N2 used by scheme 1. The two-bed commercial process20 is similar to scheme 2 and hence better than scheme 1. On the other hand, scheme 3 is safer since IPA concentration never goes above 1/4 LEL (during adsorption), and scheme 3 cuts the fresh N2 demand by almost 98% compared to scheme 1. However, if scheme 3 is compared to the previous two schemes, it has higher utility requirements. Hence, there is a trade-off between N2 makeup and utility requirements. Safety consideration always point to scheme 3. To judge whether the fresh N2 reduction in scheme 3 offsets its higher utility needs, a preliminary operating cost estimate was done. Utility costs and N2 on-site production costs are provided in Table 5. Results from Table 5 and cost data from Table 6 were used to estimate operating costs for each scheme (Table 7). The

Across-section,new Across-section,old

tF,new tF,old

)

)( )( ) Lold Lnew

Across-section,new Across-section,old

dp,new dp,old

2

(8a)

)( )( ) Lnew Lold

dp,old dp,new

(L/V)new (L/V)old

(8b) (8c)

highest and lowest costs of N2 production were used to calculate upper and lower bounds on N2 cost. From Table 7, it is clear that even the lower bound on N2 cost is orders of magnitude greater than utility costs. This implies that N2 cost dominates, and reduction in fresh N2 demand more than offsets the increase in utility requirements. Thus, the safest scheme (i.e., #3) also has the lower operating costs. Later sections of this work will show that scheme 3 is in fact better in overall cost as well. The high N2 demand of scheme 1 and scheme 2 single handedly outran all of their advantages except for the advantage of simplicity in small-scale applications. Scale Up. Industrially, solvent recovery units process feed anywhere from 60 000 SCF/h to as high as 30 000 000 SCF/h. To scale up the adsorption processes, the scaling equations in Table 8 were used.22 This method assumes that a good design (the old design) exists. Then the dimensionless heat and mass balance equations are simplified for pore diffusion control and negligible axial dispersion. The resulting equations allow accurate conditions for the new design. The variable RN is a surrogate for purity, and if RN ) 1 the two designs will have equal product purities. For high-throughput solvent recovery systems, use of large horizontal cylindrical columns with a thin horizontal adsorbent layer in the middle is common (see Figure 7).38 The smaller thickness and larger cross-sectional area of the adsorbent bed helps to prevent particle attrition and yields lower pressure drop. In addition, horizontal beds are less expensive than vertical beds.37 Industrially, a column diameter of 12 ft (3.66 m) and a bed thickness of 1 m with downward feed flow has been reported.38 These dimensions lead to less than 1% variation in cross-sectional area across the adsorbent layer.

Table 7. Operating Cost Comparison among Three Schemes in $/100 SCF Feed/h steam + refrigeration nitrogen (low: $0.05/100 SCF) nitrogen (high: $0.3/100 SCF) total

2

scheme 1

scheme 2

scheme 3

0.00016 0.005 0.03 0.00516-0.03016

0.00012 0.0023 0.0138 0.00242-0.01392

0.00016 0.00011 0.00065 0.00027-0.00081

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Figure 7. Horizontal column adsorber. Table 9. Scaled up Design for Scheme 3 feed rate adsorbent particle diameter column diameter column length bed thickness no. of columns N2 requirement distillation (enrichment) feed rate no. of theoretical trays D L

134 644 SCFM 1.4 mm (spherical) 12 ft (3.66 m) 54 ft (16.46 m) 1m 4 (installed horizontally and all of same dimensions) 1363 SCFM 50.11 kmol/h (99.88 mol % IPA) 2 1.5 ft 9.6 ft

Scheme 3 was chosen for scale-up because it is safer and has lower operating costs than schemes 1 and 2. The scaled up design is shown in Table 9. Feed flow rate and particle diameter were adjusted to ensure that the maximum pressure drop is less than 0.1 bar and the cleaned air contains less than 100 ppm IPA (averaged over one adsorption step). The amount of feed is 134 644 SCFM. For this feed rate, the IPA content in the liquid product coming out of the 4 °C condenser is 50 kmol/h. Such an IPA rate was purposely selected such that the “with N2” process can be compared to the “with steam” process, as discussed later. Further, the product liquid stream from the condenser is ∼99.88 mol % IPA, where the impurity is mostly dissolved N2. A small enrichment unit (a small distillation) was placed to enrich IPA to 99.9999 mol % IPA (parameters in Table 9). The selection of such a high purity is again for the comparative purposes as discussed later. The amount of IPA appearing in the IPA-rich product is 50 kmol/h. A complete (capital as well as operating) preliminary annualized cost analysis is done to estimate the cost of producing 100 ppm IPA air from a 5000 ppm IPA feed processed at the rate of 134 644 SCFM. For capital costs, equations presented by Douglas36 were used with the Marshall and Swift (M&S) cost index for the year 2009 as provided in Chemical Engineering Magazine.39,40 Norit Americas, Inc. was contacted for the cost of activated carbon. For heat exchangers, Coulson and Richardson (ref 42) and Arifin and Chien (ref 45) were referred for estimates on heat transfer coefficients. Table 10 lists various parameters required to estimate the capital costs. The amount of adsorbent and N2 required suggests that their costs should be taken as $ 8.77/kg and $0.05/100 SCF, respectively.

Table 10. Carbon Cost and Capital and Operating Cost Assumptions Activated Carbona 1 paper bag ) 20 kg 1 pallet ) 800 kg less than pallet (1-39 paper bags) ) $21.12 per kg 1-4 pallets ) $10.56 per kg 5-9 pallets ) $9.99 per kg 10-14 pallets ) $9.28 per kg 15 pallets plus ) $8.77 per kg Assumptions material of constructionsstainless steel both for columns and heat exchangers payback period: 3 years hours of operation/year: 8000 h installed cost horizontal column ) 57% of vertical column (based on equations in Seider37 applied to the size of columns used here.) M&S Index: 1468.6 (annual index for 2009) Heat Transfer Coefficients Distillation hHTC,reboiler ) 1420 W/(m2 K) with a differential T ) 25 °C41 [hot fluid ) steam, cold fluid ) aqueous solutions], kettle type reboiler hHTC,condenser ) 850 W/(m2 K), ∆Tcooling water ) 16.67 °C [hot fluid ) organic vapors, cold fluid ) water], floating head type condenser Adsorption fixed tube sheet heat exchangers 42 2 hHTC,condenser ) 315 W/(m K) [hot fluid ) organics (with some noncondensables), cold fluid ) water] hHTC,heater/cooler ) 285 W/(m2 K) [hot fluid ) steam, cold fluid ) gas], [hot fluid ) gas, cold fluid ) water] ) 50 W/(m2 K) [hot fluid ) gas, cold fluid ) gas] a

Courtesy: NoritAmericas, Inc., Dec. 2009.

With all the required data and equations, annualized costs can be readily estimated for scheme 3 for a feed rate of 134 644 SCFM of air with 5000 ppm IPA (Table 11). It should be noted that the nitrogen make upstream, though considerably smaller than in schemes 1 and 2, still contributes 12% to the total cost. Also, from Table 5, we know that scheme 1 needs ∼45 times and scheme 2 needs 21 times more fresh N2 than scheme 3. This implies that annualized cost for fresh N2 usage will be ∼$15 000 000/year for scheme 1 and ∼$7 000 000/year for scheme 3. Both of these costs are higher than the total annualized

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Table 11. 2009 Cost for the Four-Bed Steamless Solvent Recovery Processa costs

($/year) Adsorption

columns (purchase and installation) adsorbent heat exchangers steam, refrigeration, and cooling N2 make-up

820 000 280 000 710 000 520 000 330 000

Distillation equipment Column + trays + reboiler + condenser utility steam + cooling water total annualized cost (TAC) a

30 000 14 000 2 700 000

IPA purity in the final IPA product ) 99.9999 mol %.

Table 12. Design Specifications for Steam Regenerated Adsorber (Two Beds) feed rate adsorbent particle diameter column diameter column length bed thickness no. of columns steam

134 644 SCFM 1.4 mm (spherical) 12 ft (3.66 m) 71 ft (14.33 m) 1m 2 (installed horizontally) 12 000 kg/h

cost for scheme 2. Hence, it is clear that scheme 3 is economically better than the other two schemes for large-scale operation. As mentioned in the initial sections of this paper, a common adsorption based solvent recovery process practiced industrially uses steam. Therefore, an economic comparison of scheme 3 with a “steam regenerated adsorption based solvent recovery unit” was done. Steam Regenerated Solvent Recovery. In steam regenerated adsorbers, steam condensation is a major concern. The presence of steam after bed regeneration affects the adsorption capacity.15,43 Frequently, the coadsorption data for solvent and water are not available or are of poor quality. The following simplifying

assumptions for the design process provide satisfactory approximation to experimental results:44 (1) solvent’s adsorption is unaffected by water adsorption; (2) the temperature throughout the adsorber is that of saturated steam during regeneration; (3) and negligible variation in gas velocity across bed length Using these assumptions and the corresponding simplified equations,44 steam requirements can be estimated. For 1 kg of solvent recovered, around 4 kg of 1 atm steam is needed (T ) 373.15 K).15,43,44 The aim is to process the same feed (134 644 SCFM, 5000 ppm IPA) and compare the steam regenerated process to an earlier developed steamless process. A steam-IPA mixture coming out of adsorbers during regeneration, when condensed, contains 7.085 mol % IPA, and the rest is water. For simplicity in comparison, particle diameter, column diameter, and bed thickness have been kept unchanged (see Table 12). Once steam removes IPA from the adsorber and the mixture is condensed, it needs to be dewatered to recover IPA. Arifin and Chien45 optimized a two-column azeotropic distillation scheme using cyclohexane as the entrainer. The final IPA-rich product stream was 99.9999 mol % IPA. However, their feed was 50-50 mol % IPA-water at 100 kmol/h. Hence, a preconcentrator (distillation column) is used to enrich IPA from 7.085 to 50 mol % IPA. The three-column distillation scheme is shown in Figure 8 where C-1 is the preconcentrator. Columns C-2 and C-3 are taken from Arifin and Chien’s work. It should be noted that the four-bed steamless process uses an enrichment unit to yield 99.9999 mol % IPA product. Complete cost estimation was done for IPA recovery (using steam) wherein adsorption is followed by steam regeneration, condensation, preconcentration, and azeotropic distillation. The equipment cost provided in Arifin and Chien’s work was updated to the year 2009 using the M&S annual index. Table 13 provides an itemized cost for this process. Arifin and Chien45 also presented economic comparison between three-column

Figure 8. Distillation scheme downstream to steam regenerated solvent recovery [partially adapted from Arifin and Chien45]. “real” for columns refers to the trayed section of the column, and numbering starts from the topmost tray.

Ind. Eng. Chem. Res., Vol. 49, No. 22, 2010 Table 13. Steam Regenerated Adsorption + Azeotropic Distillation Costs costs

Table 14. Cost of Retrofitting an Existing Adsorber + Distillation Process

($/year) Adsorbers

columns (purchase and installation) adsorbent heat exchanger steam and cooling TAC adsorption

510 000 190 000 220 000 640 000 1 560 000

Distillation C-1 column equipment cost C-1 column steam cost C-1 column cooling water cost C-2 column equipment cost C-2 column steam cost C-2 column cooling water cost C-3 column equipment cost C-3 column steam cost C-3 column cooling water cost makeup entrainer cost TAC distillation (three column) TAC distillation (two column) TAC adsorption + distillation (three column) TAC adsorption + distillation (two column)

180 000 100 000 10 000 560 000 210 000 30 000 290 000 250 000 20 000 300 1 650 000 1 535 000 3 210 000 3 090 000

distillation (preconcentrator column + recovery column + azeotropic column) and two-column distillation (preconcentrator/ recovery column + azeotropic column). The two-column configuration was shown as 5-7% cheaper than the threecolumn system (Table 13). We assumed that the two-column system is 7% cheaper than the three-column system. If the total annualized cost in Table 13 (for two-column distillation configuration) is compared with that in Table 11, it is clear that the process with steam regeneration has ∼14% (with twocolumn azeotropic distillation) higher total costs than the process with hot N2 regeneration. If the solvent is immiscible with water, then the distillation scheme is not required and a relatively inexpensive decantation unit replaces distillation. So, for a simple and quick comparison for the immiscible case, total annualized costs for “adsorption alone” are compared for steam regeneration (Table 13) and hot N2 regeneration (Table 11). It is clear that use of steam is a better choice for immiscible systems as it is ∼42% cheaper than hot N2 regeneration. For the miscible azeotrope case, if we break up the total costs (or TACs) for steam and hot N2 cases, into capital and operating cost, steam regeneration is expensive both in capital and in operating cost individually. Therefore, if utilities or equipment costs change in the near future, the use of N2 still remains promising. Retrofitting. If an adsorber-distillation combination (steam regenerated adsorption) is already in place, an important question is: Will it be profitable to remove distillation and introduce a second adsorption train as in scheme 3? We assume that the distillation unit (distillation columns along with their corresponding reboilers and condensers) can be salvaged at 15% of its original cost. New incurring equipment costs will be two adsorption columns, adsorbent and heat exchangers. The salvage value of the distillation equipment is a gain and can be deducted from the new equipment cost. However, operating cost for “adsorption plus distillation” is higher compared to the fourbed “steamless adsorption” process. Thus switching to steamless adsorption leads to a reduction in operating cost. Now, the new net equipment cost can be divided by reduction in operating cost (or savings in operating cost) to estimate the number of years required to recover the new net equipment cost. This comes out to be almost 9 years as illustrated in Table 14. Such

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costs Second Train Adsorption 2 columns (D × L ) 12 ft × 54 ft each) distillation (D × L ) 1.5 ft × 9.6 ft) Adsorbent extra heat exchangers distillation salvage value (15%) net equipment cost

-1 230 000 $a -90 000 $a -420 000 $a -1 470 000 $a +431 000 $a -2 780 000 $a

Operating Costs adsorption + distillation four-bed adsorption (hot N2 regeneration) operating cost savings Years to recover net extra equipment cost

-1 170 000 $/year -864 000 $/year +306 000 $/year 9 years

a These are one-time equipment costs and not annualized cost. Care must be taken when comparing these values with values in previous tables.

a long period of cost recovery is hardly appealing. Hence, retrofit does not appear to be a good option. Discussion The hot N2 regeneration process has been simulated for one set of conditions in the present work. However, process optimization can bring further improvement. Some of the important parameters are temperature (adsorption, regeneration, condenser (4 °C currently); type of activated carbon used as adsorbent; and gas velocity (adsorption, regeneration). The material of construction in this work is taken as stainless steel, which is expensive. Costs are expected to plummet if this requirement is relaxed and a relatively cheaper material like carbon steel is used. In addition, costs can probably be reduced by using plate-and-frame heat exchangers instead of shell-andtube. As mentioned earlier, water vapor does not compete appreciably in adsorption below 50% RH. Hence, recycle streams in schemes 1 and 2 dilute the fresh feed and bring down the humidity level which allows fresh feed to have higher moisture. However, the recycle also has a maximum IPA concentration of ∼15000 ppm (saturated at T ) 4 °C) resulting in IPA concentrations greater than 1/4 LEL inside the adsorber in the presence of O2. Hence, safety measures will always point toward scheme 3. We have calculations for two casessmiscible azeotrope and completely immiscible. Another possible case is miscible but no azeotrope where only one simple distillation column is required. For a quick observation from the available data, TAC for C-1 (Table 13) is $ 290 000/year and when added to “TAC adsorption” of Table 13 yields $1 850 000/year. This cost is for producing 50 mol % IPA liquid product (C-1 distillate composition: 50 mol % IPA) using steam regeneration and is less than the steamless system (TAC in Table 11) which however produces very high purity IPA (99.9999 mol % IPA) with the help of a small enrichment unit. For a miscible system with no azeotrope, column C-1 would be designed to produce the desired solvent purity. Depending on the relative volatility of the solvent, the steam regenerated system may or may not be more economical than the steamless system. However, we can imagine four cases: (i) strong adsorption and low relative volatility; (ii) weak adsorption and high relative volatility; (iii) strong adsorption and high relative volatility; (iv) weak adsorption and low relative volatility. Case (i) is expected to favor the steamless adsorption process 3 over steam regeneration since the adsorbers will be relatively

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small but the distillation column will be relatively large and expensive to operate. However, the large heat exchanger requirement for gas systems can play a vital role and may break the economics. For case (ii) scheme 3 may not be economical compared to steam regeneration with a distillation column because the adsorber will be fairly large but the distillation column will be fairly inexpensive to operate. For cases (iii) and (iv), all one can say without a detailed economic analysis is case (iii) will be less expensive than (iv). Detailed economic analysis of both adsorption and distillation is needed for each solvent to choose between steam and steamless processes. Conclusions The conventional adsorption-based solvent vapor recovery process using steam for regeneration is an excellent process for water immiscible solvents due to its simplicity and economics. For water-miscible solvents, energy intensive distillation can pose an economic challenge. The situation is usually worse for miscible azeotropes. Using an inert gas like hot nitrogen for regeneration can avoid the distillation step. For small scale units, with more dilute feeds so that safety is not compromised, scheme 1 may be practiced because of its simplicity. At large industrial scales, N2 needs to be recycled as in scheme 3. From the preliminary cost estimates, scheme 3 is more economical than the steam regenerated adsorber-distillation combination for recovery of IPA from air. Acknowledgment This project was partially funded by the National Science Foundation (NSF) [CTS 0754906]. The financial support from Eastman Chemical Company during summer 2009 is highly appreciated. Authors are very grateful to Dr. Brian Kromer (Praxair, Inc., NY) for help on N2 costing, to Norit-Americas, Inc., for help on carbon costing, and finally to AspenTech’s technical support team for their help with ADSIM. Nomenclature A ) adsorbent bed area perpendicular to fluid flow (m2) ap ) external surface area per volume (1/m) b ) Toth adsorption isotherm coefficient (Pan), eq 1k b0 ) Toth adsorption isotherm constant (Pan), eq 1l ci ) concentration of solute i in fluid (mol/m3) jci,pore ) average concentration of solute i in pore (mol/m3) ci0 ) saturated bed solute concentration in fluid (mol/m3) cf ) concentration (or molar density) of feed (mol/m3) Cp,f ) fluid-phase heat capacity (J/(kg K)) Cp,s ) solid-phase heat capacity (J/(kg K)) Deff,i ) effective diffusivity (m2/s) Dk,i ) knudsen diffusivity (m2/s) dp ) particle diameter (m) Ds,i ) surface diffusivity (m2/s) Ds0,i ) surface diffusivity constant (m2/s) Ei ) activation energy (kJ/kmol) Ez ) axial dispersion coefficient (m2/s) Ez,T ) thermal axial dispersion coefficient (m2/s) hHTC ) heat-transfer coefficient (W/(m2 K)) ∆Hads ) heat of adsorption (kJ/kmol) j ) j-factor for calculating the heat-transfer coefficient Kd or Kd,i ) fraction of the interparticle volume that species i can penetrate kez ) thermal conductivity of the gas phase (W/(m K)) kMTC,solid ) linear lumped parameter mass-transfer coefficient (1/s)

L ) bed length (m) LMTZ ) length of mass transfer zone (m) m ) Toth adsorption isotherm coefficient (mol/kg adsorbent), eq 1k Mi ) Molar mass (kg/kmol) for Knudsen diffusivity estimation, eq 1g N1,N2, N3 ) no. of trays in columns C-1, C-2, and C-3, respectively NF1, NF2, NF3 ) feed tray number in columns C-1, C-2, and C-3, respectively Pr ) Prandtl number; Pr ) µCp,f/kez P ) pressure (Pa), eq 1k ∆P ) pressure change (Pa) Q ) molar flow rate (kmol/h), eq 8a q or qi ) amount of solute i adsorbed (mol/kg adsorbent) qj ) average amount of solute adsorbed (mol/kg adsorbent) q*i ) equilibrium amount adsorbed of species i (mol/kg adsorbent), eq 1k qji ) average amount of solute i adsorbed (mol/kg adsorbent) qjIPA ) average amount of solute IPA adsorbed (mol/kg adsorbent) Re ) Reynolds number rpore ) pore radius (m) in Knudsen diffusivity estimation, eq 1g R ) universal gas constant (kJ/K kmol), eq 1l t ) time (s) tF ) feed step time (s) T ) temperature (K), eq 1l, 1g Tj * ) average equilibrium temperature (K) Tj s ) average solid-phase temperature (K) V or Vf ) bulk fluid superficial velocity (m/s) Vpurge ) superficial velocity of purge stream (m/s) yi ) mole fraction of species i in the gas phase z ) axial distance (m) Greek Symbols εe ) external porosity (m3 void/m3 bed) εp ) internal porosity (m3 pore/m3 particle) Fb ) bed density (kg/m3) Fs ) solid-phase density (kg/m3) Ff ) fluid phase density (kg/m3) τ ) tortuosity

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ReceiVed for reView April 2, 2010 ReVised manuscript receiVed October 2, 2010 Accepted October 4, 2010 IE1008019