Economic Comparison of Crystallization Technologies for Different

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Economic comparison of crystallization technologies for different chemical products Kwan-Ling Wu, Hsing-Yu Wang, and Jeffrey Daniel Ward Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b01801 • Publication Date (Web): 17 Aug 2018 Downloaded from http://pubs.acs.org on August 18, 2018

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Industrial & Engineering Chemistry Research

Economic comparison of crystallization technologies for different chemical products Kwan-Ling Wu, Hsing-Yu Wang and Jeffrey D. Ward* Dept. of Chemical Engineering, National Taiwan University, Taipei 106-07, Taiwan

Abstract Crystallization is an ancient unit operation that remains vital for the chemical process industry. Traditional single-effect evaporation consumes a great deal of energy and various alternatives to this method have been proposed. In this work the total cost of producing a fixed quantity of different solid chemicals by crystallization from water is determined for several different technologies: evaporative crystallization (EC), membrane distillation (MD) with porous hydrophobic membranes, reverse osmosis membrane assisted crystallization, and eutectic freeze crystallization (EFC). Among the solute properties, the solubility has the greatest effect on the cost of the process since it determines the amount of water that must be removed per unit product produced. If waste heat is available at a unit price lower than that of low-pressure steam produced using coal or natural gas, then the assumed price of waste heat also has a significant effect on the economics. The results indicate that, if feasible, RO has the lowest TAC. On the other hand, for high osmotic pressure cases, EFC has the

*

Correspondence concerning this article should be addressed to Jeffrey. D. Ward at

[email protected].

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lowest cost when cheap heating energy (waste heat) is not available. On contrary, a conventional single or multiple effect evaporative crystallization has the lowest cost among the remaining options when heating energy is cheap or solubility is high. Moreover, eutectic temperature plays important role when determining the best technology when solubility is between 0.2–0.6 kg/kg solution.

1. Introduction In the chemical industry, crystallization is a commonly used technique to separate solute from solution.1–4 Usually the system is operated at a solubility higher than saturation but lower than the metastable zone5 limit to prevent primary nucleation. The driving force for crystal growth and nucleation is supersaturation.6 There are many ways to create supersaturation, including changing the temperature or adding anti-solvent. However, for large-scale continuous processes, the most common way to generate supersaturation is to remove the solvent selectively. In this way no residual material is created that requires further processing. Several methods have been developed for selectively removing solvent from a solution. By far the oldest, and still the most commonly used method is evaporation.1,7 Single-effect evaporation consumes a great deal of energy, and so many evaporation processes operate with multiple effects at different pressure, and the steam from each


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effect other than the first is used to power the subsequent effect. Recently, porous hydrophobic membranes have been proposed for removing the water from the solution as vapor8–10 while saving energy by utilizing low-grade waste heat sources.11 In such membrane-assisted processes, also called membrane distillation (MD), solvent is removed by creating a driving force across the membrane. MD processes can be classified according to the way they create this driving force. The options that show the most promise for membrane-assisted crystallization are vacuum membrane distillation (VMD) and direct contact membrane distillation (DCMD). In VMD process, the driving force for mass transfer across the membrane comes from the vacuum pressure at permeate side and high temperature on the retentate side, while in DCMD, the driving force results from low temperature at permeate side and high temperature at retentate side. These methods provide the greatest water vapor flux through the membrane, so they are the only type of porous hydrophobic membrane processes that are considered in this work. The second kind of membrane considered in this work is a reverse osmosis membrane. For systems with low osmotic pressure, reverse osmosis is an efficient way to remove water from solution.12,13 Because liquid water passes through the RO membrane, there is no need to supply the enthalpy of vaporization. Another method for removing solvent from solution is by freezing. Since



freezing the solvent will drive the solution composition to the eutectic point, where solute crystals and ice form simultaneously, this method is commonly called eutectic freeze crystallization.14–17 Since the ice is usually less dense than the solution and the solid crystals more dense than the solution, the two product streams can often be collected separately. Previously, cost analysis and comparison of evaporative and membrane assisted crystallization process for two chemical systems were completed.18 Since only two chemical systems were considered, it was not possible to draw conclusions about the effect of solute properties on the choice of crystallization technology. Hence, in this study, many more chemical systems are considered which permits determination of the effect of solute properties on the results. Furthermore, eutectic freeze crystallization is included as a technology alternative. Each of the processes described above in this section is designed for each of a number of selected solutes crystallized from water. Cost equations are used to estimate the total annual cost (TAC) for each solute for each method. Then the best process alternative can be identified depending on the values of design parameters (e.g. the cost of energy) and solute properties. 2. Model development 2.1 Crystallization process design For each process design, it is assumed that the feed condition is saturated solution at


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40 °C. In order to consider a wide variety of solutes, process designs were completed based on material and energy balances, but crystal population balances were not considered because reliable kinetic models for the crystal growth and nucleation rates in an industrial scale crystallizer are not available for many species. Therefore, the product crystal size distribution and the residence time required to achieve a given crystal size were not determined. Furthermore, it is assumed that the cost of downstream processing of the product crystals (filtration, washing, deliquoring and drying) are the same for each technology choice and therefore do not affect the relative competitiveness of one technology over another. A schematic diagram of a continuous, single-effect EC process is shown in Figure 1. The fresh feed enters the crystallizer, which is operated at 0.07 bar so that water is evaporated at the inlet temperature. After the water vapor is removed, it is condensed in a condenser. For the design of multiple-effect crystallization process, a quintuple-effect EC process is used as an example. Following Wang et al,18 as shown in Figure 2, the first and last effect are operated at 100 and 40 °C respectively, and the temperature difference between each effect is assumed to be the same. A vacuum system is utilized to achieve the corresponding desired operating pressures. After water is evaporated in the first effect, a fraction of the steam is utilized in a feed preheater and the remainder is



utilized to power the second effect. It is further assumed that there will be an energy loss of 10% between each effect, so the first effect must produce the greatest amount of steam and therefore also the greatest amount of product. Designs for membrane assisted crystallization processes are shown in Figures 3– 5. To avoid fouling by crystal growth on the membrane surface, solution is heated in a buffer tank before it enters the membrane module. As shown in Figure 3, in the VMD process, the buffer tank is operated at 80 °C, and the crystallizer is operated at the feed temperature. Vacuum is maintained using a steam jet ejector to keep the permeate side pressure at 0.04 bar. The only difference between Figure 3 and 4 is that in Figure 4 the permeate side of membrane in the DCMD process is kept at 20 °C using chilled water. As for the RO process, as suggested by Kuhn et al.,19 the buffer tank is operated at 62.3 °C. The higher temperature is necessary for RO-assisted crystallization in order to ensure that solution contacting the membrane is undersaturated to prevent crystals from forming on the membrane surface. This is in contrast to seawater desalination RO processes where the retentate concentration is less than the saturation concentration and membrane fouling by crystal formation is not a concern. It is further assumed that the maximum pressure that RO membrane can withstand is 40 bar and that the flow rate into the membrane module is ten times that of the permeate


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flow so that membrane performance will not be compromised and fouling will be reduced. Moreover, pressure exchangers are utilized that can transmit the pressure of the retentate flow to the inlet flow with 90% efficiency. A booster pump is used to provide the additional mechanical work required to raise the feed pressure to the desired value. Finally, the design for the EFC process is based on descriptions found in the literature.20–23 A schematic diagram of the process is shown in Figure 6. The crystallizer is operated at the eutectic temperature, and ice and solute product crystals are collected separately due to the difference in their density. In In order to produce pure water and avoid loss of product with the water stream, the ice crystals must be filtered and washed. The ice slurry is used to precool the inlet stream and in one design alternative, remaining ice-water slurry is used as a heat sink in a refrigeration cycle to reduce the cost of cooling. Finally, the ice mass ratio in slurry is assumed to be 30%, which is the maximum value studied by several researchers.24,25 2.2 Performance of unit operations In order to design the crystallization processes considered in this work, it is necessary to make some assumptions about the performance of various units. These assumptions are described in this section. For the design of processes based on porous hydrophobic membranes, the most



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important design parameter is flux of water through the membrane. For porous hydrophobic membranes, Table 1 shows the measured flux of vapor through the membranes under different conditions by different researchers.26–34 According to these results, in most cases, a design with retentate side at 80 °C and permeate side at 20 °C or 0.05 bar can provide a maximum flux of about 40 kg/m2-h. Therefore, these operating conditions are assumed for VMD and DCMD processes, and it is further assumed that the flux of vapor through the membrane will be 40 kg/m2-h. For highly selective reverse osmosis membranes, the flux through the RO membrane N w (in kg/h) is given by Kuhn et al.:19 =

,

(

)(∆ −

)

(1)

where P0,weff is the effective pre-exponential factor which is 7.9 ×107 kg/m2-h-bar, Eapp is apparent activation energy which is 2.92×104 J/mol, and η is polarization effect parameter which is 1.08 here. T is the operating temperature, R is the ideal gas constant, Δ is pressure imposed on the membrane, c stands for mass concentration (kg/kg solution) on the retentate side, and M stands for the molar mass of the solute (kg/mol). It is further assumed that the membrane has very high selectivity so that only a negligible amount of salt passes through the membrane. The reverse osmosis processes design includes a pressure exchanger to recover energy from the high-pressure retentate stream. Stover35 defined the efficiency of pressure


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exchanger, ηpx, as =

∑( )!"# ∑( )

%$× 100%

(2)

where P stands for pressure of fluid and Q stands for volumetric flow rate. Such devices have been reported36 to achieve an efficiency of more than 0.9 for flow rates from 4.54 to 68.1 m3/h, so in this study the efficiency of the pressure exchanger is assumed to be 0.9. Finally, filtrate flow rate per unit filter area must be known to determine the required filter area. As suggested by Seider et al.37 for coarse particles, the specific filtrate rate for a continuous rotary-drum vacuum filter was assumed to be 6000 lb/day-ft2. 2.3 Chemical systems For this study, 14 different chemical systems38,39 were considered. For simplicity and in order to ensure a fair comparison of methods, only substances that were not expected to form crystal hydrates were considered. Among these systems, only three are feasible for RO process based on the requirement that the osmotic pressure of a saturated solution at 62.3 °C must be less than 39 bar. Some important physical properties of the chemicals considered for crystallization are shown in Table 2. The eutectic point for each chemical was estimated based on solubility data (from Mullin38 and Aspen Plus) and the freezing point depression line. 2.4 Cost equations



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In this section, the cost equations used for economic calculations are presented for all units considered in this work. Unless otherwise mentioned, cost equations are taken from Seider et al.37 The capital cost for a continuous crystallizer (CCCC) is determined by:40 **++ = ,

+

-..

/ (2.4422 × 103 4 3 − 5.314648 + 5621.34 + 237830)

(3)

The cost of vessels other than crystallizers (CCvessel) is given by: **?@AA@B = (

+

-..

..-D )265C?@AA@B )3 − 0.0035546(NO> )V P

(9)

where FTm is the motor type factor, which is 1.8 for motor operating at 3600 rpm. Pc is the power consumption of the motor, which is determined by following equations. = −0.316 + 0.24015(NOI ) − 0.01199(NOI )8

(10)

= 0.8 + 0.0319(NO= ) − 0.00182(NO= )8

(11)

= = I JW/33000

(12)

+ = = /

(13)

where ηP is the pump efficiency, η M is the motor efficiency, PB means the brake



horsepower of pump, and

ρ

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is the liquid density in pounds per gallon. Finally, the

total capital cost for high pressure pump is given by: **E = Y** + **R Z

(15)

where P is the operating pressure in Torr (1–760) and V vac is the vessel volume in ft3. Then the cost of vacuum system (CCvac) can be calculated by: < = = *^aR )aR -..

(19)

+

**>>aR = ( ) (7.5800 + 0.80NO>aR )aR -..

(20)

where FBMcom is the bare module factor for a compressor, which is 2.15 and Pcom stands for the power consumption of the compressor in hp (10–200 hp for screw compressor and 200–30000 hp for a centrifugal compressor). The cost of a rotary-drum vacuum filter (CCf) is determined by: **b = (

+

-..

) (11.670 – 0.1095NOFb + 0.0554MNOFb P8 )

Economic Comparison of Crystallization Technologies for Different

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