Improving Product Recovery in Fractional Crystallization Processes

Ind. Eng. Chem. Res. 1999, 38, 823-832

823

Improving Product Recovery in Fractional Crystallization Processes: Retrofit of an Adipic Acid Plant Marcos A. B. Cesar and Ka M. Ng* Department of Chemical Engineering, University of Massachusetts, Amherst, Massachusetts 01003-3110

A systematic method is presented for improving the recovery of the desirable product in an existing fractional crystallization process. Solid-liquid-phase behaviors and techniques relevant to product recovery are discussed. Retrofit targets are identified through an analysis of the phase diagram of the mixture under consideration. Conceptual design techniques consisting of heat and mass balances, heuristics, and short-cut unit models are used to determine the necessary changes in flow sheet structure and equipment design in order to meet the retrofit objective. The method is illustrated with an adipic acid plant. Introduction Because of the intense competitive pressure, there is a continual need to improve the existing plants in the chemical processing industries for a wide variety of retrofit objectives such as1,2 (a) reducing operating costs (increasing raw materials efficiency and optimizing the use of energy), (b) increasing throughput (debottlenecking), (c) improving product quality, (d) improving process operability (flexibility and controllability), (e) reducing the environmental impact, (f) improving process safety, (g) changing the process route or introducing a new technology, (h) making a new product or processing a new feedstock. Irrespective of the goal of a retrofit project, three general tasks usually need to be accomplished: (1) identification of retrofit incentives and targets, (2) generation and analysis of new process alternatives, including new flowsheets and/or changes in operating conditions, and (3) determination of a retrofit strategy in view of the existing flowsheet structure and equipment. The techniques used for generating retrofit alternatives are similar to those employed in process synthesis. For example, dominant design variables and economic tradeoffs must be identified. However, a retrofit design requires special considerations for the constraints in the existing plant and a proper match between the old and new equipment. For this reason, the optimal retrofit is usually different from the optimal grassroots design. Retrofitting solids plants is particularly challenging. Despite their economic importance, the literature on process synthesis and retrofit methods for solids plants is still rather limited.3-5 The purpose of this research is to develop a systematic procedure for improving product recovery in a solids plant with a fractional crystallization subsystem. Fractional crystallization is a separation method for recovering pure solutes from a multicomponent solution. When a suitable sequence of operations such as heating, cooling, dilution, and evaporation is employed, it is possible to completely separate a mixture into its individual components.6-9 In many processes, however, the technique is applied to isolate a single product, since * To whom correspondence should be addressed. E-mail: [email protected].

Figure 1. Typical flowsheet structure of a fractional crystallization process.

the other components are impurities of little commercial value. In these cases, a liquid purge stream is often required to prevent the impurities from building up within the process. Figure 1 shows a typical process flowsheet using fractional crystallization to separate the desired product (A) from two other impurities (B and C) in a solvent (S). Before the crystallization step, there are usually one or more stages of concentration where some or all of the unconverted reactants, reaction byproducts, or reaction solvents are removed. The crystallizer feed (stream 6) consists of a multicomponent solution from which crystals of A are recovered. Part of the effluent mother liquor is directly recycled to the reactor via stream 11. The remainder is purged from the process (stream 13), after the solvent is recovered and recycled to the reaction system via stream 14. Unlike separation schemes for complete separations, a certain amount of dissolved product is lost along with the impurities in the purge stream. In this paper, we determine, through an analysis of multicomponent phase diagrams, the process configurations and operating conditions that minimize the product loss. We focus on ternary mixtures that do not form hydrates or compounds. The design issues in the retrofit

10.1021/ie9803671 CCC: $18.00 © 1999 American Chemical Society Published on Web 02/04/1999

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Ind. Eng. Chem. Res., Vol. 38, No. 3, 1999

Figure 2. Isothermal phase diagram for a three-solute system.

of a fractional crystallization process are discussed. The retrofit procedure is illustrated with an examplesthe production of adipic acid. Phase Behavior and Techniques for Improving Product Recovery Figure 2 shows the isothermal solubility surfaces of three solutes (A, B, and C) dissolved in the solvent S. The tetrahedral phase diagram can be in mole-fraction or mass-fraction coordinates. At a given crystallization temperature, A will precipitate out if the feed composition (point F) lies within the prism associated with the solubility surface of A. Point M represents the composition of the mother liquor after crystallization. By compressing the phase diagram onto the base as viewed from the apex, we obtain a two-dimensional representation called the Ja¨necke projection. The double saturation lines, which represent solutions saturated with two solutes, divide the triangle into three regions, one for each solute. These lines intersect at point T, the triple saturation point, at which all three solutes have reached their solubility limit. At fixed selectivities of A, B, and C in the reaction system, crystallization of pure A always occurs along the line AD, along which the ratio of B to C is constant. To improve the product recovery of A, defined as the production rate of crystals of A relative to the rate of generation of A by reaction, the mother liquor composition should be as far away from point F as possible. In doing so, we decrease the concentration of A relative to that of B and C in the mother liquor, thereby reducing the loss of A in the liquid purge. However, if the mother liquor composition is moved beyond point D, which is a double saturation point, cocrystallization of another solute (in this case B) will take place. Therefore, operation at point D corresponds to the maximum recovery of pure A at the given crystallization temperature. There are three basic techniques for improving the recovery of A. The most intuitive is to vary the crystallization temperature. Figure 3a shows the solubility surfaces at two different temperatures, Tc1 (higher) and Tc2. If the solubility of A decreases significantly with temperature, it is possible to increase the product recovery by reducing the temperature and moving the mother liquor composition from point M1 to point M2.

Figure 3. (a) Effect of crystallization temperature on product recovery. (b) Effect of solvent evaporation on product recovery. (c) Effect of mother liquor recycle (purge ratio) on product recovery.

The second technique for improving product recovery is by evaporating solvent. If solvent is removed from a solution with composition F in Figure 3b, the composition moves along the line F-F′ to a higher A concentra-

Ind. Eng. Chem. Res., Vol. 38, No. 3, 1999 825

tion. On the Ja¨necke projection, all compositions along this line are represented by the same point. Because the mother liquor composition moves in the direction of the double-saturation point D, the product recovery is increased. Therefore, by adjustment of the degree of solvent evaporation, it is possible to reach the double saturation point at any crystallization temperature. The third technique is based on recycling of mother liquor from the crystallization system to the reaction system. For an operating plant, the amount of impurities removed in the liquid purge must balance the amount generated by reaction. Therefore, by recycling more of the mother liquor to the reaction system, we increase the concentration of impurities in all streams of the process. We define the purge ratio (Rp) as the amount of mother liquor purged from the process (stream 12) relative to the amount after the solid/liquid separation step (stream 10). Figure 3c illustrates the effect of the purge ratio on the crystallizer operation. When the purge ratio is decreased and consequently the amount of mother liquor recycle is increased, the composition at the crystallizer feed will follow the line F-F1. Operation at point F1 leads to an effluent mother liquor saturated with solutes A and B, thereby giving the maximum recovery of A. Thus, by adjusting the purge ratio, we can reach the double saturation point at any given crystallization temperature, without the need for evaporating solvent from the system. Phase Behavior Types in Maximizing Product Recovery There are two types of phase behaviors of relevance to maximizing product recovery. Figure 4a depicts the phase diagram of a system with type I behavior. Point F represents the composition of the crystallizer feed for the situation where no mother liquor is recycled to the reactor (purge ratio ) 100%). If we operate the crystallizer at temperature Tc1 without solvent evaporation, the composition of the resulting mother liquor is represented by point M1. To improve the recovery of A, we need to move the operation closer to the double saturation point D1. One option is to remove solvent from the crystallizer to the limit where the composition F′1 is achieved. Another alternative is to move the crystallizer feed composition from point F to point F1, by recycling of mother liquor to the reaction system. The two techniques can also be applied simultaneously to obtain solution compositions between points F1 and F′1 in the diagram. If we decide to operate at a different crystallization temperature, maximum product recovery will be obtained at a double saturation point with a different composition. Thus, a type I phase diagram is defined to be one in which operation at a lower temperature Tc2 leads to a new double saturation point D2 which has a lower concentration of A. Therefore, the maximum recovery of A corresponds to the lowest, practical crystallization temperature. The phase diagram of a system with type II behavior is illustrated in Figure 4b. Again, point F represents the crystallizer feed composition if no mother liquor is recycled to the reactor. By operating the crystallizer at temperature Tc1 without solvent evaporation, we obtain a mother liquor with composition M1. As the solubility of A decreases at lower temperatures, the product recovery would be somewhat improved by temperature reduction. However, in a type II system, the concentration of A at the double saturation point decreases at

Figure 4. (a) Improving product recovery in a system with type I phase behavior. (b) Improving product recovery in a system with type II phase behavior.

higher temperatures. If we operate the crystallizer at a higher temperature Tc2, we can improve recovery by moving the mother liquor composition close to point D2. This is obtained by evaporation of solvent from the crystallizer or recycling of mother liquor to the reactor, to the limit where points F′2 or F2 are achieved, respectively. Evaporative crystallizers are preferred in this case because they can operate at temperatures higher than the feed temperature. The maximum recovery of A for a type II system corresponds to the highest, practical crystallization temperature such as Tc3. Process Design Using Joint Solubility Equations For designing a fractional crystallization process, we need a mathematical description of the solubility behavior of the system. A convenient method is to use the joint solubility equations, which express the solubility of each solute (S, kg of solute/kg of solvent) as a function of the concentration (c, kg of solute/kg of solvent) of the other solutes. Assuming linear behavior, the solubility equations of a three-solute system are8

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SA′(T,cB,cC) ) R1 + R2cB + R3cC

(1)

SB′(T,cA,cC) ) R4 + R5cA + R6cC

(2)

SC′(T,cA,cB) ) R7 + R8cA + R9cB

(3)

The slanted prime signifies that a single solute is saturated and the R’s are parameters. The parameters R1, R4, and R7 are the pure component solubilities of A, B, and C, respectively. The rest of the parameters can be determined by measurement of the joint solubilities of pairs of solutes at the three saturation points. We develop below an expression for the overall product recovery as a function of the key operating variables, beginning with the effluent mother liquor from the crystallizer in Figure 1, stream 10. Because the mother liquor must be saturated with A, the flow rate (F, kg/h) of A in stream 10 is

FA(10) ) R1FS(10) + R2FB(10) + R3FC(10)

(4)

We assume that the impurities B and C are completely purged from the process via stream 13. The amounts of B and C in stream 10 can be expressed in terms of their respective production rates (P, kg/h) in the reaction system:

FB(10) ) PB/Rp

(5)

FC(10) ) PC/Rp

(6)

where Rp is the fraction of B or C that is purged from stream 10 (also see eq 26). Let the amounts of reactant R and solvent S consumed in the reaction be CR and CS, respectively. For a conversion xR, the amount of R entering the reactor is CR/xR. For a fixed solvent-toreactant ratio at the reactor inlet (RS), we have

CR FS(4) ) FS(6) ) RS - CS xR

(

CR - CS xR

)

(8)

FA(9) PA - RpFA(10) ) PA PA

(9)

Combining eqs 4-6, 8, and 9, we obtain

[

Y ) 1 - R1

(

Rp(1 - Rv) CR R S - CS PA xR

FB(10) FS(10)

)] ( ) () - R2

PB PA PC R3 (10) PA

As discussed previously, the maximum recovery of A by crystallization corresponds to an operation close to the solubility limit of another solute. For a three-solute system, the Ja¨necke projection can be used to determine which double saturation line will be achieved at given

R4 + R1R5 + cC(R3R5 + R6) 1 - R2R5

)

(11)

The double slanted prime signifies that two solutes are saturated. Combining eq 11 with eqs 5, 6, and 8 gives

RP(1 - Rv) )

PB(1 - R2R5) - PC(R3R5 + R6) CR RS - CS (R4 + R1R5) xR

(

)

(12)

Equation 12 can be used to determine the values of purge ratio and solvent evaporation that produce a mother liquor saturated with A and B. Substitution of eq 12 into eq 10 yields

[Y]max A,B ) 1 -

(

) (

)

PB R1 + R2R4 PC R3R4 - R1R6 PA R4 + R1R5 PA R4 + R1R5

(13)

Similarly, for a mother liquor saturated with both A and C, the following expressions can be obtained:

SC′′(cB) )

FC(10) FS(10)

RP(1 - Rv) )

[Y]max A,C ) 1 -

The overall product recovery Y is the flow rate of A in stream 9 divided by the total amount of A generated by reaction:

Y)

SB′′(cC) )

(7)

Then, assuming that a fraction of solvent Rv is evaporated in the crystallizer and complete solid/liquid separation, we can determine the amount of solvent in the mother liquor stream:

FS(10) ) (1 - Rv) RS

product selectivities in the reactor. After identifying the impurity that will cocrystallize, we can develop expressions for the maximum product recovery based on the joint solubility equations. If, for example, the mother liquor is saturated with both A and B (cA ) SA′′ and cB ) SB′′), we have

)

R7 + R1R8 + cB(R9 + R2R8) 1 - R3R8

PC(1 - R3R8) - PB(R9 + R2R8) CR RS - CS (R7 + R1R8) xR

(

(

)

) (

(14)

(15)

)

PB R2R7 + R1R9 PC R1 - R3R7 PA R7 + R1R8 PA R7 + R1R8

(16)

Next, we demonstrate the use of this generic method to investigate retrofit alternatives for improving the overall product recovery of an existing adipic acid plant. Production of Adipic Acid Process Description. Adipic acid is an important dicarboxylic acid produced on a large scale worldwide. The predominant commercial route to adipic acid is based on the air oxidation of cyclohexane to a mixture of cyclohexanol and cyclohexanone (KA oil), followed by the oxidation of the KA mixture with nitric acid. The second step of this route is of particular interest because it includes recovery and purification of adipic acid by fractional crystallization. A typical industrial adipic acid plant flowsheet is illustrated in Figure 5. Fresh feed (stream 2) and recycle streams of nitric acid (streams 12, 15, and 18) containing the copper and vanadium catalysts are combined with the KA mixture and fed into the oxidation reactor. Oxidation is carried out at 60-80 °C and nearly ambient pressure, with total conversion of KA and 92-95% selectivity to adipic acid. The oxidation gases (NOx and CO2) are removed by air (stream 5) in a bleaching column and treated with water to recover some nitric acid. The liquid reaction effluent (stream 6) is then concentrated in a distillation column to remove the

Ind. Eng. Chem. Res., Vol. 38, No. 3, 1999 827 Table 1. Values of Some Operating Variables in the Adipic Acid Plant production rate of the dry adipic acid concentration of succinic acid, the final product residual moisture of the final product reaction conditions temperature concentration of cyclohexanol in fresh KA concentration of HNO3 in fresh nitric acid KA conversion net consumption of HNO3 mass ratio HNO3/KA at reaction feed reaction selectivities (basis ) 1 mol of KA converted) adipic acid glutaric acid succinic acid concentration of HNO3 at reaction feed (organic-free basis) excess of oxygen in the bleacher concentration of HNO3 in solution after absorption crystallizer feed temperature crystallization temperature mother liquor composition adipic acid glutaric acid succinic acid HNO3 water concentration of water in the liquid purge cake moisture after filtration (saturated cake) wash ratio (kg of water/kg of dry cake) residual cake moisture after dewatering

25 000 kg/h 50 ppm 0.1% 70 °C 60 mol % 60% 100% 0.9 kg/kg of adipic acid 7 0.92 mol/mol of KA 0.06 mol/mol of KA 0.02 mol/mol of KA 50% 100% 40% 80 °C 50 °C 6.42% 27.42% 8.17% 30.15% 27.84% 60% 46.8% 3 15%

To control the concentration of impurities in the process, part of the mother liquor is diverted to a purge treatment section via stream 16. In this section, nitric acid is removed by evaporation (stream 18) and the copper and vanadium catalyst is recovered by ionexchange treatment. The liquid purge (stream 19) consisting of an aqueous solution of dibasic acids including adipic acid is discarded as waste. An existing adipic acid plant with a production rate of 600 tons/day (or 25 000 kg/h) is considered. A list of operating variables used in the overall balance calculations is shown in Table 1. Some of the values were estimated, while others were obtained from standard references, such as the Kirk-Othmer Encyclopedia of Chemical Technology and the Encyclopedia of Chemical Processing and Design. Identification of Retrofit Incentives. The composition of the effluent mother liquor from crystallization is a key variable. Let us consider the solubility behavior of adipic, glutaric, and succinic acids in an aqueous solution. The effect of HNO3 on the solubility of the organic acids is not considered. The phase equilibria were calculated using the solubility equation:10

xeq i ) Figure 5. Simplified flowsheet structure of the adipic acid process.

excess water and the small amounts of monobasic acids (MBA) generated in the reaction. The concentration column also receives the recovered nitric acid solution (stream 9) and a wash liquor stream from the crystallization step (stream 22). Adipic acid is separated from other dibasic acids (DBA), such as glutaric and succinic acids, by vacuum crystallization at 50-60 °C. The solid product is recovered by filtration, washing, and drying. Additional purification steps for obtaining polymergrade adipic acid including solvent switching from a nitric acid solution to water, charcoal treatment, and recrystallization from water are omitted.

[ (

∆Hmi 1 1 1 exp γi R Tmi T

)]

(17)

Here, T is temperature, ∆Hmi, the heat of fusion, and Tmi, the melting temperature. The liquid-phase activity coefficient, γi, was predicted using the UNIFAC groupcontribution method. Figure 6 shows the Ja¨necke projection of the threedimensional solubility surfaces. The ternary diagram represents mass compositions of different solutions on a dry basis. At a given equilibrium temperature, the double saturation lines divide the triangle in three single saturation regions, one for each component. Because glutaric acid is much more soluble than the other components, its saturation region is considerably smaller. This means that high concentrations of glutaric

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Figure 7. Effect of temperature on the concentration of adipic acid at the double saturation point. Figure 6. Ja¨necke projection of the solubility surfaces for an aqueous solution of adipic, glutaric, and succinic acids predicted by the UNIFAC method.

acid are required for its crystallization to occur. The dashed line represents a locus of compositions with a constant ratio of glutaric to succinic acids. This ratio is determined by the relative selectivities of these components in the reaction system. Thus, crystallization of adipic acid always occurs along this material balance line provided that the reaction selectivities remain unchanged. If desirable, the solubility parameters in eqs 1-3 can be backed out from the data for Figure 6. The necessary equations and calculational procedure are provided elsewhere.8 If all of the mother liquor were purged from the process (Rp ) 100%), the composition of the crystallizer feed (on a dry basis) would be represented by point F1. By crystallizing adipic acid from this solution at 50 °C, the resulting mother liquor would correspond to point M1. At this process condition, the loss of product in the purge would be excessive because of the high relative concentration of adipic acid in the mother liquor. In the existing process, we assume that most of the mother liquor from the crystallizer is recycled to the oxidation reactor. When the purge ratio is reduced from 100% to 1.5%, the crystallizer feed composition is moved from point F1 to point F2. Consequently, crystallization at 50 °C results in a mother liquor with composition M2, which is close to the solubility limit of succinic acid. Further reduction of the purge ratio is not desirable, because it would lead to cocrystallization of succinic acid. This means that the existing process is already assumed to be operating at the maximum possible recovery of adipic acid, for the given crystallization temperature. Despite this assumption, the potential for improved product recovery still exists. By examining the effect of temperature on the solubility data, we observe that the concentration of adipic acid in the double saturation points decreases at lower temperatures or, in other words, the system behaves as a type I phase diagram. This effect is better quantified in Figure 7 which shows the ratio of the amount of adipic acid to the sum of glutaric and succinic acids RA at the double saturation point as a function of crystallization temperature. Clearly, the temperature affects RA significantly, especially in the range between 60 and 30 °C. Because the production of glutaric and succinic acids is assumed to

be constant, changes in the ratio RA have a direct impact on the amount of adipic acid that is lost in the liquid purge. Thus, there is an incentive for improving the product recovery of the existing process by reducing the crystallization temperature. Because the process is already operating close to a double saturation point, our retrofit strategy is to increase the purge ratio, while simultaneously reducing the crystallization temperature, to keep the concentration of succinic acid always close to its solubility limit for each temperature value. Let us examine the necessary equations for determining Figures 6 and 7 as well as the flow rates in other streams in some detail. Balances around the Reaction System. We define the selectivities SA, SGL, and SSU as the fractions of KA converted that correspond to the production of adipic, glutaric, and succinic acids, respectively. For oxidation reactions, it is reasonable to assume that these selectivities remain constant despite the change in operating conditions of the retrofitted plant. Thus, for an adipic acid production rate of PA (kmol/h), the amounts of organic compounds involved in the oxidation reaction are (note that P has units of kg/h in the general analysis)

FKA ) PA/SA

(18)

PGL ) FKASGL

(19)

PSU ) FKASSU

(20)

PCO2 ) PGL + 2PSU

(21)

To simplify our calculations for the recovery of the nitric acid consumed, we assume that the reacted nitric acid is reduced to a mixture of NO and N2O. Also all NO is reconverted to nitric acid by bleaching and absorption with water according to the following reaction:

2NO + 3/2O2 + H2O f 2HNO3

(22)

The N2O formed during oxidation cannot be recovered and corresponds to a net loss of nitric acid in the process. The consumption of fresh nitric acid FHNO3 (mol of HNO3/mol of adipic acid) is experimentally determined, and we calculate the amount of N2O as follows:

PN2O ) PAFHNO3/2

(23)


For a given fraction of cyclohexanol in the KA mixture, xCHL, the net production of water can be estimated by an atomic balance of hydrogen in the reaction system:

NPH2O ) (2xCHL + 10)FKA + (FHNO3 - 10)PA - 8PGL - 6PSU 2 (24) Similarly, we can use an atomic balance of oxygen to calculate the amount of NO produced by oxidation:

PNO ) [4(PA + PGL + PSU) + 2PCO2 + PN2O + NPH2O - FKA - 3FHNO3PA]/1.5 (25) Once the value of PNO is known, the amount of nitric acid recovered and the corresponding water consumption can be estimated using the stoichiometry of reaction (22). Balances around the Crystallization System. An important operating variable in this process is the purge ratio of mother liquor from crystallization, Rp, which is the ratio of the flow rate in stream 16 to that in 14 of Figure 5:

Rp )

F(16) F(14)

(26)

Different values for the purge ratio result in changes of flow rate and composition for most of the streams. Crystallization is assumed to occur by adiabatic vacuum cooling. A fraction Rv of the amount of solvent in the crystallizer feed, FS(11), is evaporated and recycled to the reactor after condensation:

Rv )

F(12) FS(11)

(27)

The vapor fraction Rv is fixed by the crystallizer heat balance, according to the following relationship:

λvFS(11)Rv ) F(11)Cp(Tf - Tc) + λcFA(26) (28) Here, λv and λc are the heats of vaporization and crystallization, respectively. The left-hand side represents the heat removed by solvent evaporation; the two terms on the right-hand side represent the sensible heat released by the feed and the heat released by crystallization, respectively. Because we recycle mother liquor to the reactor, the total flow rate of the crystallizer feed stream F(11) depends on the value of the purge ratio Rp. An overall material balance around the crystallization and filtration steps gives

F(11) ) RvFS(11) + F(16)/Rp + F(20)

(29)

Similarly, we can write the balance above for the solvent S (nitric acid + water) and solve for Rp:

Rp )

FS(16) (1 - Rv)FS(11) - FS(20)

(30)

When eqs 28-30 are combined, the following expression for the vapor fraction Rv can be obtained:

[

{

Rv ) λcFA(26) + Cp(Tf - Tc) F(20) +

}] [ { [ ]}]

F(16) [F (11) - FS(20)] / FS(11) λv + FS(16) S F(16) -1 Cp(Tf - Tc) FS(16)

(31)

Overall Balances. The production rate of dry adipic acid is fixed at FA(26) ) 25 000 kg/h. Because the mother liquor contains some dissolved product, part of the adipic acid generated by reaction is lost in the liquid purge. We define the variable RA as the ratio of the amount of adipic acid to the amount of other dibasic acids in the mother liquor:

RA )

FA(16)

(32)

FGL(16) + FSU(16)

The generation of glutaric and succinic acids by reaction must be equal to the removal of these impurities in the liquid purge. Thus, the loss of adipic acid in the purge is

FA(16) ) RA(PGLMWGL + PSUMWSU)

(33)

Here, MW is molecular weight. For a given production rate of dry adipic acid, FA(26), we can substitute eqs 18-20 into eq 33 to get the total amount of product generated by reaction:

PA )

FS(26) (34) RA MWA (S MWGL + SSUMWSU) SA GL

[

]

Once the value of PA is known, the material balances for the reaction system can be calculated as described earlier. We have summarized the key heat and mass balance equations above. Cesar11 provides a more detailed discussion of the calculational procedure and reports all of the flows in this adipic acid plant. In addition, various solids processing models which are essential for examining the operating conditions are also presented in the thesis. Simulation Results Again, the key operating variables in our retrofit analysis are the crystallizer temperature (Tc) and the purge ratio (Rp). The impact of different process conditions on the operation of the existing plant is examined below. Table 2 presents a list of equipment parameters and physical properties used in the retrofit study. Figure 8 demonstrates the effect of the simultaneous variations of Tc and Rp on the overall product recovery. If the crystallization temperature is reduced from 50 to 20 °C, the purge ratio must be increased from 1.5 to 12%, to avoid the cocrystallization of succinic acid. Because the concentration of adipic acid in the purge decreases at lower temperatures, the overall product recovery is improved from 98.6% to 99.2%. The existing crystallizer is assumed to operate by vacuum adiabatic cooling. Thus, the flash temperature is regulated by pressure control in the crystallizer. Figure 9 shows the required reduction in pressure for

830


Figure 8. Effect of retrofit on product recovery. Table 2. Values of Some Equipment Parameters and Physical Properties for the Adipic Acid Plant crystallizer volume filter area filter parameters cycle time vacuum level submergence medium resistance air temperature for drying typical porosity of filter cake heat of crystallization (adipic acid in water) crystallization kinetics for adipic acid12 densities adipic acid glutaric acid succinic acid nitric acid solution (50%) at 50 °C at 20 °C viscosity of nitric acid solution at 50 °C at 20 °C heat capacities adipic acid nitric acid solution (50%) air latent heat of vaporization for water at 50 °C at 20 °C diffusion parameters13

Figure 9. Effect of retrofit on crystallizer pressure and solvent evaporation.

237 m3 26 m2 2.2 min 20 000 N/m2 35% 1.0 × 109 m-1 175 °C 0.47 51 kcal/kg i ) 3.5, j ) 0.4 kr ) 9.1 × 1032 1360 kg/m3 1424 kg/m3 1572 kg/m3 1287 kg/m3 1316 kg/m3 0.0011 kg/(m s) 0.0019 kg/(m s) 0.38 kcal/(kg °C) 0.70 kcal/(kg °C) 0.24 kcal/(kg °C) 570 kcal/kg 587 kcal/kg kd ) 1.3 × 10-13 c1 ) 1.52, c2 ) -1.05

each temperature level, according to vapor-liquid equilibrium data of nitric acid solutions (Perry’s Chemical Engineers Handbook, 6th ed.). To operate at the new conditions, the existing vacuum pump or steam ejector will probably have to be replaced. Figure 9 also shows that more solvent is evaporated from the crystallizer at lower temperatures because of the effect of the sensible heat released by the feed stream. Care should be taken to ensure that excessive solids entrainment is not a problem.14 Also, temperatures below 30 °C will require recompression of the solvent vapors, to avoid the use of refrigeration in the condensation equipment. Figure 10 illustrates the effect of the purge ratio on the median crystal size, according to the MSMPR (mixed suspension mixed product removal) model. When the purge ratio is increased, the flow rates of all streams in

Figure 10. Effect of retrofit on crystallizer feed flow rate and median crystal size.

the mother liquor recycle loop are reduced, including the crystallizer feed. As a result of higher residence times, the median size of the crystals is increased. The effect is more pronounced at purge ratio values between 1.5 and 6%, which correspond to crystallization temperatures between 50 and 30 °C. Within this range of process conditions, the mother liquor composition changes more significantly, which explains the higher variations of flow rate observed. The influence of the new operating conditions on the filtration step is demonstrated in Figure 11. As mentioned, the mother liquor flow rate decreases at higher purge ratio values, especially in the range of 1.5-6%. Because the production rate of crystals is fixed, the magma density in the crystallizer increases accordingly, which explains the initial drop of the required cake formation time. Part of this effect also results from the reduction in cake resistance due to the presence of larger crystals. Because we are simultaneously reducing the crystallization temperature, the filterability of the crystals is also affected by the higher viscosities of the mother liquor. At purge ratios above 6%, the median crystal size and the concentration of solids do not change


Figure 11. Effect of retrofit on concentration of solids and cake formation time.

Figure 13. Effect of retrofit on residual cake moisture and air for dryer.

the residual wash liquor from the cake. As a consequence, we need to increase the air flow rate in the dryer to achieve the moisture specification for the final product. Economics of Retrofit and Equipment Replacements

Figure 12. Effect of retrofit on impurity concentration and required wash ratio.

significantly. Thus, the increase in viscosity tends to be the dominant effect and higher filtration times are required. It is interesting to note that the existing filter will not have to be modified to operate at the final retrofit conditions because the initial and final filtration times are practically the same. Figure 12 shows the effect of the retrofit modifications on the cake washing step. An increase in the purge ratio decreases the concentration of impurities in the mother liquor. As a consequence, less wash solvent is required to achieve the purity specification for the final product. This also results in a load reduction in the concentration column, which receives the wash effluent as a feedstream. Finally, the influence of the new process conditions on the dewatering and drying steps is illustrated in Figure 13. Similar to the filtration case, the cake moisture after dewatering is affected by changes in crystal size and wash liquor viscosity. However, because of the combination of opposite effects, the residual cake moisture does not change significantly. At lower temperatures, more energy is required for evaporation of

The flow rate and composition of each stream in the new process, after simultaneous reduction of the crystallization temperature to 20 °C and increase of the purge ratio to 12%, are compared to those of the existing plant to identify the necessary equipment retrofit modifications. The economic incentive for the retrofit project is based on the incremental investment costs and savings in operating costs associated with the proposed modifications. The major process bottleneck for the retrofit modifications is found to be the purge evaporation column. Because the new purge flow rate before evaporation (stream 16) is almost 5 times as high as that in the existing plant, the existing column will have to be replaced. Only the major equipment costs are considered, which are annualized using a capital charge factor of 1/3. Table 3 shows the results of this screening economic analysis using cost models available in the literature. Clearly, the savings in product recovery exceeds by far the costs of installing a new purge evaporation unit. In addition, the incremental energy costs associated with purge evaporation and drying are compensated for by energy savings in the concentration column. As a result, an economic potential of $1 437 000/ yr can be realized. This estimate shows a large incentive for modifying the process according to the proposed strategy. Therefore, it is advisable to examine secondary equipment replacements omitted in this study. In the reactor, for example, the total flow rate is reduced by 30%. Because the mother liquor recycle helps to moderate the temperature change through the reactor, additional heat removal capacity is probably needed. Depending on the diameter of the existing concentration column, the reduced flow rate may lead to minimum vapor rate constraints (column weeping). A retrofit alternative in this case is to plug up sieve tray holes in the existing column to reduce area available for vapor flow. Another

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Table 3. Economic Incentive for Retrofitting the Adipic Acid Plant savings in product recovery annualized cost of a new purge evaporation unit (column + auxiliaries) incremental energy costs (steam) purge evaporation drying energy savings in a concentration column (steam) economic potential of retrofit unit prices adipic acid steam at 150 psig steam at 15 psig Marshall and Swift (M&S) index, 1995

possibility is to increase the vapor rate in the column by increasing the reflux ratio, at the expense of higher operating costs in the reboiler and condenser. Furthermore, this retrofit study has not included optimization analysis. This means that the selected values for crystallization temperature and purge ratio may not correspond to the maximum economic potential. However, according to the solubility behavior depicted in Figure 7, the incremental savings in product recovery for crystallization temperatures below 20 °C is unlikely because of the need for refrigeration. Conclusions A procedure is presented for the retrofit of a fractional crystallization process to improve product recovery. It consists of three main tasks. First, the solid-liquidphase behavior is examined to determine the retrofit incentives, if any, with known techniques in fractional crystallization. After the necessary changes in flowsheet configurations are identified, except for new streams, the flow rates and compositions of each stream of the new plant and existing plant are compared. Then, the equipment modifications or replacements are determined. Third, the final retrofit design is determined by taking into account the process constraints and the investment costs for the perceived retrofit benefits. The procedure is applied for a simplified adipic acid plant. In the existing plant, the crystallization temperature and purge ratio are fixed at 50 °C and 1.5%, respectively, which lead to an overall adipic acid recovery of 98.6%. According to the solubility data generated by UNIFAC, the solution of dibasic acids behaves as a type I system. With reduction of the crystallization temperature to 20 °C with a simultaneous increase of the purge ratio to 12%, the process can operate at 99.3% of adipic acid recovery. With short-cut unit models and approximate cost models, the estimated economic potential of the retrofit project is $1 437 000/yr, signaling sufficient incentives for a more detailed study. Acknowledgment We acknowledge a Fulbright Fellowship awarded to M.A.B.C. and the leave granted by Rhone Poulenc,

$2 060 000/yr $341 000/yr $1 142 000/yr $126 000/yr $986 000/yr $1 437 000/yr $0.7/lb $3/1000 lb $2/1000 lb 1027.5

Brazil, for his graduate studies. Additional financial support was provided by the National Science Foundation (Grant CTS-9211673). Literature Cited (1) Grossmann, I. E.; Westerberg, A. W.; Biegler, L. T. Retrofit Design of Processes. In Computer Aided Process Operations; Reklaitis, G. V., Spriggs, H. D., Eds.; CACHE, Elsevier: New York, 1987. (2) Gundersen, T. Retrofit Process DesignsResearch and Applications of Systematic Methods. In Foundations of Computer Aided Process Design; Siirola, J., Grossmann, I., Stephanopoulos, G., Eds.; Elsevier: New York, 1990. (3) Rajagopal, S.; Ng, K. M.; Douglas, J. M. A Hierarchical Procedure for the Conceptual Design of Solids Processes. Comput. Chem. Eng. 1992, 16, 675. (4) Hill, P. J.; Ng, K. M. Simulation of Solids Processes Accounting for Particle Size Distribution. AIChE J. 1997, 42, 7715. (5) Gruhn, G.; Rosenkranz, J.; Werther, J.; Toebermann, J. C. Development of an Object-Oriented Simulation System for Complex Solids Processes. Comput. Chem. Eng. 1997, 21, S187. (6) Fitch, B. How to Design Fractional Crystallization Processes. Ind. Eng. Chem. 1970, 62 (12), 6. (7) Cisternas, L. A.; Rudd, D. F. Process Designs for Fractional Crystallization from Solution. Ind. Eng. Chem. Res. 1993, 32, 1993. (8) Dye, S. R.; Ng, K. M. Fractional Crystallization: Design Alternatives and Trade-Offs. AIChE J. 1995, 41, 2427. (9) Berry, D. A.; Ng, K. M. Separation of Quaternary Conjugate Salt Systems by Fractional Crystallization. AIChE J. 1996, 42, 2162. (10) Walas, S. M. Phase Equilibria in Chemical Engineering; Butterworth: Boston, 1985. (11) Cesar, M. A. B. Retrofit Design of Solids Processing plants. M.S. Thesis, University of Massachusetts, Amherst, MA, 1997. (12) Liu, C. H.; Zhang, D. H.; Sun, C. G.; Shen, Z. Q. The Modeling and Simulation of a Multistage Crystallizer. Chem. Eng. J. 1991, 46, 9. (13) Reid, R. C.; Prausnitz, J. M.; Poling, B. E. The Properties of Gases and Liquids; McGraw-Hill: New York, 1987. (14) Farrell, R. J.; Outland, J. Opportunities for Steady-State Optimization of Industrial Vacuum Evaporative Crystallizers. Paper presented at the AIChE Annual Meeting, Los Angeles, Nov 1997; Paper 16 g.

Received for review June 10, 1998 Revised manuscript received October 20, 1998 Accepted October 22, 1998 IE9803671

Improving Product Recovery in Fractional Crystallization Processes

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