Effect of separation material particle size on pressure drop and process efficiency in continuous chromatographic separation of glucose and fructose
Accepted Manuscript Effect of separation material particle size on pressure drop and process efficiency in continuous chromatographic separation of gl...
Accepted Manuscript Effect of separation material particle size on pressure drop and process efficiency in continuous chromatographic separation of glucose and fructose Jari Heinonen, Quentin Sanlaville, Henna Niskakoski, Tuomo Sainio PII: DOI: Reference:
Please cite this article as: J. Heinonen, Q. Sanlaville, H. Niskakoski, T. Sainio, Effect of separation material particle size on pressure drop and process efficiency in continuous chromatographic separation of glucose and fructose, Separation and Purification Technology (2017), doi: https://doi.org/10.1016/j.seppur.2017.10.049
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EFFECT OF SEPARATION MATERIAL PARTICLE SIZE ON PRESSURE DROP AND PROCESS EFFICIENCY IN CONTINUOUS CHROMATOGRAPHIC SEPARATION OF GLUCOSE AND FRUCTOSE Jari Heinonen*, Quentin Sanlaville, Henna Niskakoski, Tuomo Sainio School of Engineering Science Lappeenranta University of Technology Skinnarilankatu 34, FIN-53850 Lappeenranta, Finland
ABSTRACT Effect of the particle size (250–320 µm) of gel type strong cation exchange resins in Ca2+ form on the pressure drop across the resin bed and on the efficiency of chromatographic glucose–fructose separation was investigated experimentally. Pressure drop caused by the resin bed was found to depend strongly on the particle size. Strong dependency of the pressure drop on the solution viscosity was also observed. No elastic deformation of the resin particles was observed as in each case the dependence of the pressure drop on flow rate was linear. Both material (water consumption) and energy aspects were addressed in the investigation of the effect of particle size on the efficiency of continuous glucose–fructose separation. An eight-column four-zone simulated moving bed (SMB) process was used. The water consumption was found to decrease by 53 % by change of particle size from 320 µm to 250 µm. This could be achieved without compromising specific productivities or product yields and purities by proper adjustment of the operating parameters. Lower water consumption results in more concentrated product fractions which decreases the energy consumption of the product concentration step (evaporation). Change in resin particle size from 320 µm to 250 µm decreased the energy consumption in evaporation by 18 %. Such a change in the particle size resulted in 50 % increase in the energy consumption of the pumps due to higher pressure drop. However, the total (pumps and evaporation) energy consumption decreased as the product concentration step is significantly more energy intensive unit operation.
Keywords:
chromatographic separation; pressure drop; material efficiency; energy efficiency; simulated moving bed; ion exchange resin; particle size
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1
INTRODUCTION
Chromatographic separation of glucose and fructose in the production of high fructose corn syrup (HFCS) is by volume one of the largest applications of chromatography [1]. HFCS is a widely used sweetener in food and beverage industries due to its higher sweetness, better color and hygroscopic characteristics, and due to reduced viscosity when compared to sucrose [2–4]. It is produced from corn starch by first hydrolyzing the starch into glucose which is subsequently converted partially into fructose enzymatically: a solution containing 42 % fructose (HFCS-42) is formed [2–5]. Part of HFCS-42 is processed further using chromatographic simulated moving bed process (SMB) to produce a 90 % fructose containing solution (HFCS-90). HFCS-90 is mixed with HFCS-42 to obtain the most commonly used commercial HFCS solution containing 55 % fructose (HFCS-55), or is sold as such [4]. HFCS is a relatively low value product and, thus, the SMB process must be carefully optimized. The chromatographic separation of glucose and fructose using SMB is well-proven technology and the effect of operating parameters and separation material on the separation efficiency has been investigated in detail by a number of authors [4,6–13]. However, there are no reports on the effect of particle size on the material efficiency and energy consumption of an SMB process for this application. In industry, the chromatographic separation of glucose and fructose is carried out using gel type strong cation exchange resins in Ca2+ form as the separation material with particle size in the range of 310-320 µm [14–16]. A decrease in the particle size could further increase the efficiency of this important separation task due to lower mass transfer resistance and bulk phase dispersion and should therefore be investigated in detail. An interesting aspect related to particle size is the energy consumption. With smaller particles, more concentrated product streams could be obtained, which results in lower energy consumption in the product concentration step following the separation step. However, a decrease in the particle size increases the pressure drop across the resin bed. This increases the energy consumption of the pumps needed in the process. In addition, existing process equipment usually have a certain maximum allowed pressure drop which further limits the usable particle size. Thus, the optimal particle size is a trade-off between separation efficiency and pressure drop.
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The effect of particle size on the separation efficiency of chromatographic processes and on pressure drop has been investigated by a number of authors. Ludemann-Hombourger et al. [17] have investigated the effect of particle size on the performance methyl mandelate enantiomer separation on Chiralcel OD stationary phase (dp < 50 µm) with simulated moving bed (SMB) process. Optimization of the particle size and column dimensions within set pressure drop limits showed that the smallest particle size enabling good column packing was found to be optimal [17]. Jo et al. [18] optimized the particle size of silica with respect to productivity in the separation of amino acids. The optimal particle size was found to be at the boundary between the pressure limiting (flow rate has to be limited due to the maximum allowed pressure drop) and mass transfer limiting (slow mass transfer due to large particle size) regions [18]. In the aforementioned studies [17,18] the effect of particle size on the separation performance and pressure drop was investigated with rigid particles. However, studies dealing with gel type particles have also been carried out. Houwing et al. [19] optimized the particle size of Sepharose Big Beads in the separation of bovine serum albumin and myoglobin as a trade-off between pressure drop and process performance in SMB. Adachi et al. [20] investigated the effect of resin particle size on the separation of a number of aromatic compounds, and on pressure drop. Mohammad et al. [21] investigated the effect of particle size of Sephadex G-100 size exclusion gel on the pressure drop and compression of the particles (elastic deformation) at varying flow rates, and found out that with smaller particles, the compression occurs at lower flow rates. The elastic deformation is caused by axial compression [22], and the velocity at which it occurs depends on the stationary phase [23]. Sainio et al. [13] investigated the effect of particle size on the separation of glucose and fructose and on sugar beet molasses fractionation using batch chromatography. In both systems, the separation performance increased with decreasing particle size [13]. It is clear that decrease in the particle size results in higher separation efficiency, but also in higher pressure drop across the resin bed. However, the effect of particle size on eluent consumption or energy efficiency was not investigated in the aforementioned studies. Here we present an experimental investigation of the effect of the particle size of a gel type strong cation exchange resin in Ca2+ form on the efficiency of chromatographic separation of glucose and fructose in an SMB process. Both material and energy efficiency of the SMB process for glucose–fructose separation are addressed. The effect of particle size on the specific 4
productivities and product yields in this separation task has been investigated earlier [13], but only in a batch mode making the application of the results to a continuous SMB process challenging. In addition, eluent (water) consumption related to the separation task was not addressed in [13]. This aspect was also ignored in [17–20] in which the effect of particle size on the performance of SMB processes was studied. Eluent consumption is an important process performance indicator and should be taken into account. Here, minimization of eluent (water) consumption in glucose–fructose separation is of particular interest. The effect of particle size of the gel type cation exchange resins on the pressure drop across the resin bed is also investigated. The investigation is carried out using single-column batch chromatography setup. This data is utilized in the evaluation of the energy efficiency of the chromatographic separation of glucose and fructose.
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EXPERIMENTAL METHODS
2.1
Materials
Gel type strong cation exchange resins CS11GC (Finex Oy, Kotka, Finland) with PS–DVB matrix (cross-link density 5.5 wt.% DVB) with four different particle sizes (Table 1) were used as stationary phase materials in the experiments. The resins were used in Ca2+ form. Ultrapure water produced with CENTRA R 60/120 (ELGA LabWater, Aquaflow Ltd.) water purification system was used in all experiments and in preparation of all solutions. Analysis grade D-(+)-glucose (≥ 99.5 %, Sigma-Aldrich) and D-(-)-fructose (≥ 99.0 %, Sigma-Aldrich) were used in the experiments. Blue Dextran 2000 (Pharmacia Biotech/GE Healthcare) was used in resin bed porosity measurements. <
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2.2
Effect of particle size on pressure drop
The effect of particle size on the pressure drop across the resin bed in the glucose–fructose separation was studied using a batchwise chromatographic setup consisting of one dual-piston HPLC pump (Intelligent pump AI-12-33, Flom Inc.) for liquid feeding; two glass columns in series (with water-heating jackets; YMC ECO SR 15/750, d = 1.5 cm; YMC Europe GmbH) connected by pipe; water circulation thermostat (C6 CS, Lauda Dr. R. Wobser GmbH & Co. KG,); and a pressure detector (PM10A, Nokeval Oy) with pressure sensors at column inlet and outlet (NAT 10.0A, Trafag AG). The flow rate range was 1-8 m/h given as superficial velocity (Q = 2.9-23 mL/min). This was chosen as it is on the level used in the industrial scale glucose– fructose separation. The experiments were carried out at 50 °C temperature. The bed height of one column was 70 cm (total bed height = 140 cm). The data from the pressure detector was collected using Labview software (version 14.0, National Instruments). Each pressure drop measurement was carried out twice to have reliable results. Pure water, 28 wt.% glucose solution, and 56 wt.% glucose solution were used as the feed solutions. 56 wt.% glucose solution has the same dynamic viscosity at 50 °C temperature as authentic high fructose corn syrup (59-60 % dry solids of which 42-45 % fructose and 50-53 % glucose) at 60 °C temperature (typical process temperature). The dynamic viscosity of the glucose solution corresponding to the authentic HFCS42 was adjusted with a digital viscometer (DV-II, Brookfield AMETEK Inc.). The exact values of the dynamic viscosities of the sugar solutions were measured at the temperatures used in the experiments with a rheometer (MCR 302, Anton Paar). The densities of the solutions were measured at the temperatures used in the pressure drop measurements using DMA 4500 density meter (Anton Paar). In addition to the measurements done with the glucose solutions and water, the pressure drop caused by the 320 µm resin was measured at a temperature of 60 °C with synthetic HFCS42 solution with 50 wt.% dry solids (58 % glucose, 42 % fructose) as the feed solution.
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The pressure drop caused by the equipment (pipes, column adapters, etc.) was measured with each of the feed solutions using columns filled only with the feed solutions (no resin in the columns). The length of the fluid beds in the columns was 140 cm in total (70 cm + 70 cm). The resin bed porosity was determined from the retention times of 0.5 mL pulses of 1.5 g/L Blue Dextran 2000 solutions with water as eluent (flow rate = 1 mL/min) after the pressure drop measurements with each resin.
2.3
Effect of particle size on separation process performance
The effect of particle size on the material and energy efficiency of chromatographic separation of glucose and fructose was studied using a 4-zone simulated moving bed (SMB) process (Fig. 1). The setup consisted of four dual-piston HPLC pumps (Intelligent pump AI-12-33, Flom, Inc.) for eluent and feed introduction and extract and raffinate outtake, one single-piston HPLC pump (Series III, SMI-Labhut Ltd.) for recycle stream pumping, eight glass columns with water-heating jackets (YMC ECO SR 15/750, d = 1.5 cm, YMC Europe GmbH), water circulation thermostat (Alpha A24, Lauda Dr. R. Wobser GmbH & Co. KG), six 12-port motor valves (RV-750-116, IDEX Health & Science Group), and eight 3-port solenoid valves (MLV-3-1-1/4UKGH-3, Takasago Electric Inc.). The bed height of one column was 70 cm (Vcol ≈ 123.7 mL). The column configuration of the SMB system (2:3:2:1; Fig. 1) was chosen on the basis of an industrial scale SMB process used in glucose–fructose separation. The recycle stream was pumped from the last column to the first column through a small (15 mL) buffer tank (Fig. 1) to ensure stable operation of the recycle pump. The experiments were carried out at a temperature of 50 °C. The SMB system was controlled using Labview software (version 14.0; National Instruments).
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The minimum purity requirements of glucose and fructose in the outlet streams were set to 89 % and 90 %, respectively. Thus, the extract stream had the same target purity for fructose as the common commercial grade high fructose corn syrup with 90 wt. % fructose on dry solids (HFCS90). SMB runs were conducted with switch times of 5 min (all particle sizes) and 5.79 min (250 µm particle size). Due to this, the linear flow velocities are comparable with those used in the industrial scale processes. The SMB runs were done using a feed solution with approximately 40 wt.% sugar solution (250 g/L glucose and 180 g/L fructose; HFCS-42) was used as the feed solution in the SMB runs. 40 wt.% concentration was used due to limits set by the laboratory scale SMB setup. The total bed porosity at a temperature of 50 °C was determined for each resin particle size with 3 mL pulses of 3 g/L Blue Dextran 2000 solutions. The pulses were eluted through all the columns with a flow rate of 3 mL/min. Each SMB experiment consisted of two phases: start-up and processing. In the start-up phase, the process was run unsupervised with the same m-parameter values as in the processing phase, but with lower (1/3) flow rates and longer (3x) switch times. During the start-up and processing phases, approximately 9-11 and 10-12 full SMB cycles were run, respectively; the total amount being 21. During the processing phase, samples were collected on each cycle from extract and raffinate streams. In addition, at the end of each run, samples were collected from the outlet of each column in order to obtain approximation of the spatial profiles inside the columns. These samples were collected during two minutes with a flow rate of 5 mL/min.
2.4
Carbohydrate analyses
The carbohydrate concentrations were determined with an offline HPLC (HP/Agilent 1100, Agilent Technologies, Inc.) equipped with a refractive index detector. A sugar analysis column 8
Shodex SP-0810 was used with a Shodex SP-G guard column (both from Waters Corporation). Purified water was used as eluent with 0.8 mL/min flow rate, the temperature was set to 80 °C, and the injection volume was 10 µL.
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CALCULATIONS
3.1
Pressure drop
i The pressure drop ptotal was calculated as the difference in the readings of the pressure sensors
1 (before the columns) and 2 (after the columns) i ptotal p2i p1i .
(1)
i where ptotal is the total pressure drop caused by the resin bed and equipment in solution i. i presin can be calculated from
i i i presin ptotal pequipment ,
(2)
i where pequipment is the pressure drop caused by the equipment in solution i. The pressure drop
caused by the equipment was correlated with a second-order polynomial i i i i pequipment equipment v2 equipment v equipment ,
(3)
i i where equipment and equipment are the flow resistance coefficients of the equipment in feed solution i i and equipment describes possible offset in the measurement for this particular feed solution.
i Here, the total pressure drop ptotal , caused by the resin bed and the equipment, and the pressure i drop presin , caused by the resin bed, were scaled to unit length of the bed by dividing the
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pressure drop value with the bed length L. In the case of the pressure drop caused by the resin bed, the scaled values were found to depend linearly on flow velocity with each feed solution, and were therefore correlated with linear models i presin i i resin v resin , Lbed
(4)
i i where resin is the flow resistance coefficient of the resin bed in solution i and resin is
measurement error term. In the case of the total pressure drop, nonlinear dependency of the scaled values on the flow rate was observed with pure water whereas with the other feed solutions linear behavior was observed. Therefore, the scaled values were correlated with a second-order polynomial i ptotal i i i total v 2 total v total , Lbed
(5)
i i i where total and total are the total flow resistance coefficients in solution i and total is i measurement error term. For the sugar solutions, the values of total were zeros.
3.2
Operating parameters of the SMB process
The operating parameters of an SMB process can be given as dimensionless m-parameters that describe the ratio of the flow rate of the liquid phase and the effective flow rate of the solid phase in zone k (k = 1–4) [24–27] mk
Qk tswitch bedVcol , Vcol ( 1 bed )
(6)
where Qk is the volumetric flow rate in zone k, tswitch is the duration of one switch in an SMB, and Vcol is the volume of one column. The initial values of the m-parameters (for 280 µm resin; reference) were calculated from the operating parameters of an industrial scale SMB process for glucose–fructose separation where 10
the CS11GC resin in Ca2+ form is currently used as the separation material. The feasible operating area was determined with the 280 µm resin by trial-and-error method by taking the mparameters calculated on the basis of the industrial process as initial guess. The m2 and m3 parameters, i.e. the flow rate in zone II and the amount of fresh feed to zone III, were then varied in the subsequent SMB runs until the feasible operating area was found. With the other particle sizes (250 µm and 280 µm), the first SMB runs were done using the feasible operating points of the 280 µm resin. Use of this approach is justified by the fact that only the particle size, not the actual separation material, was changed. Particle size affects only the size of the feasible operating area due to changes in apparent dispersion. Only change in the separation material would affect the position of the feasible operating area making the determination of the feasible operating parameters by trial-and-error method very laborious. The amount of feed to the SMB process in one switch can be given as a dimensionless parameter defined as the difference between the m-parameters of the second and third SMB zones
mFF m3 m2 ,
(7)
where superscript FF stands for fresh feed. 3.3
Process performance
The separation performance of the SMB process was evaluated with product purities and yields, and with specific productivities, water consumption, water-to-feed ratio, and energy consumption of product concentration step. Purity of component i in process outlet stream x (extract for x fructose and raffinate for glucose) Pui was calculated from
x
Puix
Ci n
C
100% ,
(8)
x j
j 1
where C is the average liquid phase concentration and the sum in the nominator stands for the x total concentration. Yield of component i Yi was calculated with
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x
Ci Q x Yi FF FF 100% , Ci Q x
(9)
Specific productivity of i with respect to the total resin bed volume of the system Pri was calculated with x
Ci Q x , Pri ncolVcol
(10)
where ncol is the total number of columns in the SMB train. Specific water (eluent) consumption needed to produce one mass unit of product ECi was calculated with
ECi
Q water x
Ci Q x
,
(11)
where superscript water stands for the water (eluent) stream. Water-to-feed ratio was calculated with
Q water W /F FF . Q
(12)
The effect of resin particle size on the energy consumption related to pumping of the liquids in the SMB and on the steam and heat consumption of the product concentration step were also evaluated on the basis of the data obtained in the experiments. In order to obtain more realistic understanding on the energy consumption, the SMB process was scaled up so that the volume of a single column in the SMB was 6.3 m3 instead of 120 mL used in the experiments. The flow rates of the SMB process were scaled up by keeping the reduced flow rate (bed volumes/h) as a constant. Evaluation of the heating power and steam consumption was done for a single-effect evaporator. In this evaporator type, when the boiling point elevation of the concentrate is negligible, the heat transfer rate on the liquid side q can be obtained from [28] 12
where Cp,F is the specific heat of the feed, Tb is the boiling temperature, TF is the temperature of
the feed, w is the mass flow of the concentrated product stream, w F is the mass flow of the evaporator feed stream, and λ is the latent heat of vaporization of water at the pressure of the vapor space in the evaporator. In the calculation of q following values were used and assumptions were made. The total concentration of the final product stream was the same as the feed to the SMB process, i.e., 430 g/L. The boiling point elevation of the product stream was assumed to be negligible. The pressure of the saturated steam was 2.0 bar which yields a λS value of 2202 J/g [28]. The absolute pressure in the vapor space was 0.1 bar. At this pressure, the boiling point of water Tb,water is 45.82 °C [29] and the latent heat of vaporization of water is 2393 J/g [28]. The boiling point of the evaporator feed stream can be calculated with [29]
Tb Tb,water Tb Tb,water kbmF ,
(14)
where kb is the ebullioscopic constant of the solvent (water), 0.513 K kg/mol [29], and mF is the molality of the evaporator feed stream. Tfeed of 50 °C was assumed for the evaporator feed stream. Cp of water at a temperature of 50 °C, 4.1813 kJ/(kg K) [29] was used as Cp,feed. The energy consumption of the pumps in the SMB process was calculated as a sum of the power requirements of the eluent, feed, and recycle pumps. It was assumed that the extract and raffinate streams were regulated by adjustable flow restrictors instead of pumps. Only the pressure drop caused by the resin bed was taken into account in the calculations as the pressure drop caused by the pipes and other process equipment (column adapters, etc.) are not affected by changes in the particle size of the separation material. Thus, the pressure drop was calculated using the flow resistance coefficient of the resin bed only (Eq. (4)). In addition, the flow resistance coefficient obtained with 28 wt.% glucose solution was used. Although the feed of the SMB process is more concentrated, it is diluted immediately when it enters into the SMB process. In addition, zones 1 and 4 in the SMB process are only partially filled with the (diluted) feed solution. Due to these 13
facts, 28 wt.% glucose concentration is closer to the average sugar concentration in the SMB unit than 56 wt.% concentration. The power P supplied to a pump from an external source can be calculated from [28]
P
wi H
,
(15)
where wi is the mass flow of stream i, η is pump efficiency, and ΔH is the total pump head [28]
H H b Ha ,
(16)
where H is the head, and subscripts a and b stand for suction and discharge, respectively. The head H can be calculated from [28] H
p
gZ
v2 ,
(17)
2
where p is pressure, ρ is density of the solution, g is gravitational acceleration, Z is height, α is kinetic energy correction factor, and v is the superficial velocity. It was assumed that in case of eluent and recycle pump the liquid had the density of pure water throughout the SMB system (ρ = 999.82 kg/m3 [29]). With the feed pump, the density of a glucose solution with 430 g/L concentration, 1157.17 kg/m3 [29] was used. The height difference between the pump suction head and discharge head (ΔZ = Zb–Za) was assumed to be zero [28]. η and α values were 0.7 and 1.0, respectively. 4
RESULTS AND DISCUSSION
4.1
Effect of particle size on pressure drop
In order to calculate the pressure drop caused by the resin bed from the total pressure drop i measured, the pressure drop caused by the equipment ( pequipment ) was determined (Fig. 2).
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i pequipment increased with increasing dynamic viscosity of the feed solution: the lowest values
were obtained with pure water and the highest with 56 wt.% glucose solution. Interestingly, the i dependence of pequipment on the flow rate was found to be somewhat nonlinear in pure water,
whereas with the other feed solutions nonlinear behavior could not be detected beyond experimental accuracy. The flow resistance coefficients for Eq. (3) determined from the data shown in Fig. 2 are given in Table 2.
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The reliability of the results was assessed by determining the pressure drop caused by the 320 µm resin with a synthetic HFCS42 solution with 50 wt.% dry solids at a temperature of 60 °C (Fig. 3) and by comparing the results with the pressure drop caused by Dowex Monosphere 99 in same conditions [14,15]. The pressure drop data measured here for the 320 µm resin with HFCS42 HFCS42 (50 wt.% d.s.) was on the same range as estimated for the Dowex Monosphere 99. resin of
320 µm resin in HFCS42 (50 wt.% dry solids) was 0.103. Two different values, with some 20% HFCS42 difference, for resin of Dowex Monosphere 99 (Ca2+ form) with 320 µm particle size could be
estimated from the data found in literature: 0.108 [14] and 0.121 [15]. As these values are close to that obtained in this work, the reliability of the measurements was deemed to be good.
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Porosity of the resin bed affects the pressure drop caused by the bed. In this study, the pressure drop was measured for each material 2 or 3 times with 7 flow rate values (typically 14 to 21 measurements), and the flow resistance coefficients were estimated from the data as detailed above. The porosities measured after packing in water varied somewhat between the beds of different particle sizes (Table 3) and may also have changed during the repeated measurements due to compression of the bed under pressure. It should thus be borne in mind that the differences in bed porosity are included in the values of the flow resistance coefficients reported here. i was found to depend significantly on the particle size of the CS11GC resin in Ca 2+ form presin
(Fig. 4, Table 3). For example, with 56 wt.% glucose solution, a change in particle size from 320 µm to 310 µm increased Δpresin by 32%, from 320 µm to 280 µm by 115 %, and from 320 µm to 250 µm by 250 % (Table 3). These changes are large but seem to be on a realistic level. It has been reported in [15] for Dowex Monosphere 99 (Ca2+ form) that a change in the particle size from 320 µm to 310 µm increases the flow resistance coefficient by approximately 66 % , from 0.121 to 0.201 (HFCS42, 50 wt.% d.s., 60 °C). The cause of the higher increase in pressure drop with Dowex Monosphere 99 is unknown, but might be caused by different particle size distribution. Although the effect of particle size on Δpresin is significant, the increase in Δptotal was considerably smaller (Fig. 4, Table 3). In the aforementioned examples, the increase in Δptotal with 56 wt.% glucose solution were 7%, 24, % and 51 % with the equipment used in these measurements. The higher the pressure drop caused by the process equipment, the lower the increase in Δptotal with decreasing particle size. It is thus expected that in an SMB process, the relative increase in total pressure drop when switching to smaller particles is lower than estimated directly from the experimental Δpresin values. The effect of the particle size on pressure drop decreased with decreasing dynamic viscosity of the feed solution (Fig. 4, Table 3). For example, a change in particle size from 320 µm to 280 µm increased Δpresin by 115 % in 56 wt.% glucose solution, 46 % in 28 wt.% glucose solution, and 37 % in pure water. The significantly large difference between 56 wt.% glucose solution and 28 wt.% glucose solution can be explained in the large increase in the dynamic viscosity of glucose solutions at this concentration range: dynamic viscosity of 28 wt.% glucose solution was 16
1.35 mPa s whereas that of 56 wt.% glucose solution was 7.19 mPa s. However, when an SMB process is considered, the pressure drop increase caused by changing to smaller particles is not the same as measured with the feed solution. It is expected to be lower because the average concentration in the beds of SMB unit are lower than the feed concentration. i With each particle size, the presin was found to depend linearly on the flow rate with each feed
solution (Fig. 3; Fig. 4 B, D, F). This indicates that deformations of the resin particles were negligible in the flow rate range 1–8 m/h. An elastic deformation resulting from axial compression would have been seen as concave upward behavior of the Δpresin/Lbed vs. v curves [21–23]. Linear behavior has also been observed in case of Dowex Monosphere 99 resin in Ca2+ form [14,15]. i Linear dependency of ptotal on the flow rate was observed in the case of 28 wt.% and 56 %
glucose solutions (Fig. 4 A, C) and HFCS42 with 50 wt.% dry solids (Fig. 3). On the other hand, slightly nonlinear behavior of Δptotal was observed in pure water (Fig. 4E). This was caused by the nonlinearity of the Δpequipment curve (Fig. ) and Δpresin was found to increase linearly with flow velocity in this case, too.
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4.2
Effect of particle size on the separation performance
The separation of glucose and fructose on cation exchange resins is based on ligand exchange. Fructose is more strongly adsorbed to the resin than glucose due to stronger complexes formed between fructose molecules and Ca2+ ions in the resin [30,31]. Due to stronger adsorption, fructose elutes towards zone 1 in SMB and is taken out of the process through the extract port. Glucose as the less strongly adsorbed component is taken out through the raffinate port. 17
The effect of particle size on the separation performance in glucose–fructose separation was studied with 250 µm, 280 µm, and 320 µm resins using a four-zone SMB process (Table 4). 280 µm resin was taken as a reference against which the other resins were compared. With the other resins, the first SMB runs were conducted using the same operating parameters (mk, tswitch) with which the target purities were met with 280 µm resin. In total, 14 SMB runs were conducted. Of these, six were initial runs done with the 280 µm resin in order to find the feasible operating area for the CS11GC resin in Ca2+ form (data not shown).
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With each resin, the cyclic steady state conditions were attained in 10 cycles after the start-up phase. Typical evolution of the product purities in the outlet streams is shown in Fig. 5A for 320 µm resin (similar trends were observed with each resin). Purity of fructose in the extract stream increased during the runs due to a decrease in glucose concentration and increase in the fructose concentration at the end of zone 1 (collection point of the extract stream). On the other hand, the purity of glucose in the raffinate stream decreased due to the increased fructose concentration and decrease in the glucose concentration at the end of zone 3 (collection point of the raffinate stream). Typical spatial profiles of glucose and fructose inside the SMB train at steady state conditions are shown in Fig. 5B for the resin with 320 µm particle size. The profiles obtained in each SMB run showed similar characteristics. The front of the glucose profile in zone 3 are clearly less dispersed than that of the rear of the profile (Fig. 5). This results from low flow rate in zone 4. On the other hand, the rear of the fructose profile was less dispersed than the front due to low average sugar concentration in zone 1 resulting from the high flow rate in zone 1. The concentrations of both glucose and fructose at the end of column eight were very low in each run (Fig. 5). Thus the recycle stream did not contain neither of these in significant amounts.
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Increase in the particle size increases the effective dispersion and thus leads to decreased separation performance. Due to this, the SMB run with 320 µm resin using the same operating parameters (mk, tswitch) as used with the 280 µm resin resulted in product streams that did not fulfill the purity requirements (run 2 in Table 4). In order to meet the target purities, the operating parameters (mk) had to be adjusted (larger m1 and m2, smaller m3) in order to meet the target purities. This resulted in 30 % decrease in the amount of feed introduced to the system when compared to the 280 µm resin (run 3 in Table 4). Due to larger m1, the amount of water used increased, and, thus, the water to feed ratio with the 320 µm resin increased by approximately 47 % from that of 280 µm resin (Table 4).
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Increase in the amount of water and decrease in the amount of feed introduced to the SMB with 320 µm resin resulted in more diluted product streams. This was seen in lower Pri and C i , and in higher ECi with 320 µm resin over 280 µm resin (runs 1 and 3 in Table 4). More dilute product fractions increase the energy consumption of the product concentration step as more water has to be removed to obtain the final product concentrations. The obtained results demonstrate that the separation efficiency can be increased considerably by changing the resin particle size from 320 µm to 280 µm. Decrease in the particle size from 280 µm to 250 µm resulted in increased separation efficiency due to weaker effective dispersion. When the same operating parameters (sk, tswitch) were used with 250 µm resin as with 280 µm resin, notably higher product purities were obtained with the smaller particle size (run 4 in Table 4). Due to this the m-parameters could be adjusted to allow treatment of increased amount of feed: 24 % more feed could be treated (larger m3, smaller m2) with 250 µm resin than with 280 µm resin without compromising the target purities (run 5 in Table 4). Increase in mFF led to considerable improvements in the Pri, ECi, C i , and W/F compared to the 280 µm resin. Higher product concentrations lead to decrease in the energy 19
consumption of the product concentration step as less water has to be removed to obtain the final product concentrations. It was shown that decrease in the particle size increases the separation performance of the SMB process in the glucose–fructose separation. However, when the particle size is decreased the pressure drop across the resin bed increases. As is seen in Table 3, decrease in the particle size causes notable increase in Δptotal and Δpresin regardless of the feed solution. However, the increase in pressure drop caused by smaller particle size can be compensated to a certain limit by use of lower flow rates. In fact, this should increase the separation performance as effective dispersion decreases with decreasing flow rate. Here this effect was demonstrated with the 250 µm resin by decreasing the flow rates by 14 %. This decrease results in 14 % decrease in Δptotal and Δpresin due to linear dependency of the pressure drop on the flow rate (see Fig. 4). In order to keep the operating parameters constant, i.e., constant Qk*tswitch term in Eq. (7), tswitch had to be, thus, increased by 16 % due to decrease in flow rates. Due to the decrease in the flow rates, a small increase in Puglucose could be observed (run 6 in Table 4), but Pufructose was not changed when same m-parameter values were used as in run 5. However, this was most likely due to analytical errors, and in practice also Pufructose was increased. Slight decrease in Pri was observed due to decreased flow rates when compared to run 5. As the decrease in the flow rates leads to increased separation performance, higher amount of feed could be treated in the SMB without compromising the target purities. Here, increase of mFF from 0.23 to 0.26 (run 7 in Table 4) led to a notable increase in the process performance (compared to run 6 in Table 4) with respect to Pri, ECi, and C i without compromising the target purities. In fact, Pri and ECi values were on the same level as in run 5 with the 250 µm resin. In addition, a 12 % decrease in W/F was obtained. Thus the decrease of Pri in run 6 due to lower flow rates could be compensated by proper adjustment of the operating parameters. With mFF = 0.3 (run 8 in Table 4), Puglucose was above the set purity limit, but Pufructose was slightly lower than the target value. Thus, it can be concluded that the minimum W/F with which the set target purities could be reached with 250 µm resin is 1.21 (run 7 in Table 4). This value is 20
over two times smaller than the value reached with the 320 µm particle size and over 30 % smaller than with the 280 µm particle size. These results clearly show that decrease in particle size in the chromatographic glucose–fructose separation lead to significant improvements in the process performance.
4.3
Energy efficiency
The effect of the particle size on the power requirements of the pumps of the SMB process and the heat and steam consumption of the product concentration step was investigated using the data shown in Table 4. Decrease in the particle size results in more concentrated product streams due to lower effective dispersion. This lowers the energy consumption and costs related to the final concentration of the product streams after the separation. For example when a single-effect evaporator is used, decrease in the resin particle size from 280 µm (run 1 in Table 4) to 250 µm (run 7 in Table 4) decreases the required heating power and steam consumption by approximately 10 % (Table 5). On the other hand, a 10 % increase in the heating power and steam consumption is obtained if 320 µm resin (Table 5) is used instead of 280 µm resin. Only single-effect evaporator was considered in the energy efficiency calculations. Even better energy efficiency would be obtained with a multiple-effect evaporator [28], but this would not affect the relative changes in the heating power requirements. The energy required by the boiler used to generate the steam was not evaluated here, but the changes in the particle size should have the same effect on the energy needed by the boiler as on the steam and heat consumption. Decrease in the particle size increases the pressure drop caused by the resin beds. This increases the energy consumption related to pumping of the solution in the SMB process (Table 5). Change in the particle size from 280 µm to 250 µm particle size increases the energy consumption of the pumps by 36 %. On the other hand, change of the particle size from 280 µm to 320 µm decreases their energy consumption by 30 %. Although the relative increase in the pumping power demand is higher than the relative decrease in the heating power demand (Table 5), the overall power requirement is dominated by the evaporation. Therefore, smaller particle size is recommended as the product stream is more concentrated and the heating for the evaporation is significantly more energy intensive than pumping of solutions. 21
<
>
22
5
CONCLUSIONS
An experimental investigation of the effects of ion exchange resin particle size on the pressure drop and on the efficiency of the chromatographic glucose–fructose separation was carried out. The pressure drop measurements were carried out using a batchwise chromatographic separation unit whereas the efficiency of the glucose–fructose separation was investigated using a four-zone simulated moving bed process. The efficiency of the separation process was investigated with respect to material efficiency (water consumption) and energy efficiency (energy consumption of pumps in the SMB process and heat and steam consumption of product concentration step). Gel type strong cation exchange resins in Ca2+ form with varying particle sizes were used as stationary phases. Pressure drop caused by the resin bed was found to depend considerably on the particle size of the separation material. In addition, strong dependency of the pressure drop on the solution viscosity was also observed. In each case, however, the dependence of the pressure drop on flow rate was linear. This clearly indicates that no elastic deformation of the gel type cation exchange resin particles occurred at the flow rate range used. The material efficiency of the glucose–fructose separation with respect to water consumption was found to increase considerably when smaller resin particles were used. Minimization of the water consumption did not affect the productivity of the process when the operating parameters were properly adjusted. The lower water consumption resulted in more concentrated product streams which results in lower energy consumption of the product concentration step. On the other hand, as expected, the pressure drop across the resin bed increased considerably with decreasing particle size. This results in higher energy consumption related to the pumping of solutions. However, the concentration of the product streams is usually carried out by evaporation which is a more energy intensive unit operation than pumping of solutions. Thus, the total energy consumption of the process decreases due to lower energy required by the product concentration step. The results shown here demonstrate that the economics of the chromatographic glucose–fructose separation can be increased considerably by using smaller resin particles. Although the pressure 23
drop increases with decreasing particle size, both the material and energy efficiency of the separation process can be improved by this, and by proper adjustment of the operating parameters. Acknowledgements The authors are grateful for Finex Oy (Finland) for providing the ion exchange resins used in this study. Financial support from Academy of Finland (grant SA/298548) is gratefully acknowledged.
NOMENCLATURE Ci Cp ECi g H ΔH kb Lbed mF mFF mk ncol P p Δp Pri
Puix Qk q Tb TF t tswitch Vcol V W/F
average liquid phase concentration of i, g/L specific heat, j/(kg K) specific water (eluent) consumption needed to produce 1 ton of product, m3/ton gravitational acceleration, m/s2 pump head, m2/s2 total pump head, m2/s2 ebullioscopic constant of the solvent, K kg/mol length of the resin bed, m molality of the evaporator feed stream, mol/kg dimensionless figure for the amount of fresh feed fed to SMB process, operating parameter of SMB process, number of columns in the SMB train, power supplied to a pump from an external source, W pressure, Pa or bar pressure drop, bar specific productivity of i with respect to the total resin bed volume, kg/(m3 (bed) h) purity of component i in outlet stream x, % volumetric flow rate in zone k, mL/min heat transfer rate, W boiling temperature, K temperature of the feed, K time, h duration of one switch in an SMB cycle, min column volume, mL superficial velocity, m/h or m/s water to feed ratio, -
w
mass flow of the evaporator concentrate, kg/s 24
wF
mass flow of the evaporator feed stream kg/s
wi
Yi
x
Z
mass flow of stream i, kg/s yield of component i in outlet stream x, % height, m
Greek letters α β εbed η λ λS ρ χ
flow resistance coefficient, bar h/m flow resistance coefficient, bar h2/m2 resin bed porosity, pump efficiency, latent heat of vaporization of water, J/g latent heat of condensation of steam, J/g solution density, kg/m3 measurement error, bar
Subscripts and superscripts a b equipment F FF water resin total
pump suction head pump discharge head process equipment feed stream fresh feed water (eluent) stream resin bed resin bed and the surrounding process equipment
25
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G. Ganetsos, P.E. Barker, Semicontinuous countercurrent chromatographic refiners, in: G. Ganetsos, P.E. Barker (Eds.), Prep. Prod. Scale Chromatogr., Marcel Dekker, Inc., 1993: pp. 233–256. V. Bravo, E. Jurado, G. Luzón, N. Cruz, Kinetics of fructose-glucose isomerization with sweetzyme type A, Can. J. Chem. Eng. 76 (1998) 778–783. doi:10.1002/cjce.5450760413. A.M. Dehkordi, M.S. Tehrany, I. Safari, Kinetics of glucose isomerization to fructose by immobilized glucose isomerase (Sweetzyme IT), Ind. Eng. Chem. Res. 48 (2009) 3271– 3278. doi:10.1021/ie800400b. J.S. White, HFCS, sucrose, and fructose: history, manufacture, applications, and production, in: J.M. Rippe (Ed.), Fruct. High Fruct. Corn Syrup, Sucrose Heal., Springer Science+Business Media, New York, 2014. E. Palazzi, a Converti, Generalized linearization of kinetics of glucose isomerization to fructose by immobilized glucose isomerase., Biotechnol. Bioeng. 63 (1999) 273–284. K. Hashimoto, S. Adachi, H. Noujima, H. Maruyama, Models for the separation of glucose/fructose mixture using a simulated moving-bed adsorber, J. Chem. Eng. Japan. 16 (1983) 400–406. doi:10.1252/jcej.16.400. Y.L. Cheng, T.Y. Lee, Separation of fructose and glucose mixture by zeolite Y, Biotechnol. Bioeng. 40 (1992) 498–504. V. Viard, M.-L. Lameloise, Modelling glucose-fructose separation by adsorption chromatography on ion exchange resins, J. Food Eng. 17 (1992) 29–48. D.C.S. Azevedo, A.E. Rodrigues, Fructose-glucose separation in a SMB pilot unit: Modeling, simulation, design, and operation, AIChE J. 47 (2001) 2042–2051. doi:10.1002/aic.690470915. M.S. Coelho, D.C.S. Azevedo, J.A. Teixeira, A. Rodrigues, Dextran and fructose separation on an SMB continuous chromatographic unit, Biochem. Eng. J. 12 (2002) 215– 221. doi:10.1016/S1369-703X(02)00071-2. H.J. Subramani, K. Hidajat, a. K. Ray, Optimization of Simulated Moving Bed and Varicol Processes for Glucose–Fructose Separation, Chem. Eng. Res. Des. 81 (2003) 549– 567. doi:10.1205/026387603765444500. K. Lee, Continuous separation of glucose and fructose at high concentration using twosection simulated moving bed process, Korean J. Chem. Eng. 20 (2003) 532–537. doi:10.1007/BF02705561. T. Sainio, A. Kärki, J. Kuisma, H. Mononen, Influence of Ion Exchange Resin Particle Size and Size Distribution on Chromatographic Separation in Sweetener Industries, Ion Exch. Solvent Extr. (2011) 145–169. Dow, Dow Water Solutions DOWEX TM MONOSPHERE TM Ion Exchange Resins Chromatographic Separation of Fructose and Glucose with DOWEX MONOSPHERE Ion Exchange Resins Technical Manual Chromatographic separation of fructose and glucose with DOWEX TM MONOSPHERE TM resins, (n.d.). P.D. Sheet, DOWEX TM MONOSPHERE TM 99 Ca / 310 chromatographic separation resin, (n.d.) 1–2. PRINCIPAL APPLICATIONS TYPICAL PACKAGING Purolite ®, (2017) 31385. http://www.purolite.com/product-pdf/pcr642ca. O. Ludemann-Hombourger, M. Bailly, R.M. Nicoud, Design of a simulated moving bed: 26
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Optimal particle size of the stationary phase, Sep. Sci. Technol. 35 (2000) 1285–1305. doi:Doi 10.1081/Ss-100100225. S. Jo, Q. Han, Y. Suh, J. Ryu, S.C. Yi, K.B. Lee, S. Mun, Particle-Size Optimization for a Polymer Coated Silica Gel in SMB Chromatography for Amino Acid Separation, J. Liq. Chromatogr. Relat. Technol. 6076 (2009) 2822–2838. doi:10.1080/10826070903288839. J. Houwing, H.A.H. Billiet, L.A.M. Van der Wielen, Mass-transfer effects during separation of proteins in SMB by size exclusion, AIChE J. 49 (2003) 1158–1167. doi:10.1002/aic.690490509. T. Adachi, S. Ando, J. Watanabe, Characterization of synthetic adsorbents with fine particle sizes for preparative-scale chromatographic separation, J. Chromatogr. A. 944 (2002) 41–59. doi:10.1016/S0021-9673(01)01370-X. A.W. MOHAMMAD, D.G. Stevenson, P.C. Wankat, Pressure drop correlations and scaleup of size exclusion chromatography with compressible packings, Ind. Eng. Chem. Res. 31 (1992) 549–561. doi:10.1021/ie00002a016. A. V. Danilov, I. V. Vagenina, L.G. Mustaeva, S.A. Moshnikov, E.Y. Gorbunova, V. V. Cherskii, M.B. Baru, Liquid chromatography on soft packing material, under axial compression. Size-exclusion chromatography of polypeptides, J. Chromatogr. A. 773 (1997) 103–114. doi:10.1016/S0021-9673(97)00034-4. R.N. Keener, E.J. Fernandez, J.E. Maneval, R.A. Hart, Advancement in the modeling of pressure-flow for the guidance of development and scale-up of commercial-scale biopharmaceutical chromatography, J. Chromatogr. A. 1190 (2008) 127–140. doi:10.1016/j.chroma.2008.02.113. M. Mazzotti, G. Storti, M. Morbidelli, Optimal operation of simulated moving bed units for nonlinear chromatographic separations, in: J. Chromatogr. A, 1997: pp. 3–24. doi:10.1016/S0021-9673(97)00048-4. M. Mazzotti, Equilibrium theory based design of simulated moving bed processes for a generalized Langmuir isotherm, J. Chromatogr. A. 1126 (2006) 311–322. doi:10.1016/j.chroma.2006.06.022. J. Siitonen, T. Sainio, Unified design of chromatographic separation processes, Chem. Eng. Sci. 122 (2015) 436–451. doi:10.1016/j.ces.2014.10.004. T. Sainio, Unified Design of chromatographic processes with timed events: Separation of ternary mixtures, Chem. Eng. Sci. 152 (2016) 547–567. doi:10.1016/j.ces.2016.06.038. W.L. McCabe, J.C. Smith, P. Harriott, Unit Operations of Chemical Engineering, 7th ed., McGraw Hill, 2005. W.M. Haynes, PHYSICAL CONSTANTS OF ORGANIC COMPOUNDS in the CRC Handbook of Chemistry and Physics, 2014. K. Dorfner, Introduction to ion exchange and ion exchangers, in: K. Dorfner (Ed.), Ion Exch., Walter de Gruyter, Berlin, 1991: pp. 7–189. F. Helfferich, Ion exchange, Dower Publications Inc., Mineola, 1995.
27
FIGURE CAPTIONS
Figure 1.
Schematics of the 4-zone SMB (2:3:2:1 column configuration) used in the investigation of chromatographic separation of glucose and fructose. Roman numerals = zone number; Arabic numerals = column number.
Figure 2.
Pressure drop caused by the experimental setup (pipes, column adapters, connectors). Symbols: black ●, water (T = 50 °C); brown ■, 28 wt.% glucose (T = 50 °C); red ▲, HFCS42 with 50 wt.% dry solids (T = 60 °C); blue ◆, 56 wt.% glucose (T = 50 °C). Lines: fit of Eq. (3). Error bars mark the standard error.
Figure 3.
Pressure drop caused by a bed of CS11GC resin in Ca2+ form with 320 µm mean particle size with HFCS42 solution (50 wt.% dry solids) at a temperature of 60 °C. HFCS42 HFCS42 Symbols: red ▲ = ptotal ; black ● = presin . Lines: red and black = fit of the HFCS42 Eqs. (4) and (5); blue = presin for Dowex 99 Monosphere with 320 µm particle size, estimated from data presented in [14] (dashed) and [15] (solid). Error bars mark the standard error.
Figure 4.
Effect of particle size of CS11GC resin in Ca2+ form on pressure drop with 56 wt.% glucose solution (A-B), 28 wt.% glucose solution (C-D), and pure water (E-F) at a temperature of 50 °C. Subfigures: A,C,E = Δptotal; B,D,F = Δpresin. Symbols: black ●, dp = 320 µm; orange ■, dp = 310 µm; red ▲, dp = 280 µm; green ▲; blue ◆, dp = 250 µm. Lines: fit of the model equations (Eqs. (4) and (5). Error bars mark the standard error.
Figure 5.
Evolution of product purities upon reaching steady state conditions (A) and outlet concentrations at each column at steady state conditions (B) in chromatographic glucose–fructose separation with a 4-zone SMB process. Resin: CS11GC in Ca2+ form. Particle size: 320 µm. Operating conditions: run 3 in Table 4. Ex = extract outlet; R = raffinate outlet. Symbols: black ● = glucose; red ◆ = fructose. Lines in A: black dash-dot = target purity of glucose (89 %); red dashed = target purity of fructose (90 %)
28
Water
I 1
Glucose-fructose feed
I 2
II 3
Extract Fructose rich
II 4
II 5
III 6
III 7
IV 8
Raffinate Glucose rich
Buffer tank
29
30
32
33
TABLES Table 1.
Particle size of the CS11GC resins (cross-link density 5.5 wt.%) in Ca2+ form used in this study. Nominal particle size, µm 250 280 310 320
Mean particle size, µm 251 282 312 322
34
Within ± 20 % 95.1 95.8 94.1 95.9
Table 2.
Values of flow resistance coefficients in Eq. (2) for the equipment determined from the pressure drop data in Fig. 2. βequipment, bar h2/m2 αequipment, bar h/m εi, bar
Pure water 28 wt.% glucose HFCS42 (50 wt.% d.s.) 56 wt.% glucose
3.63⋅10-3
7.78⋅10-2
3.00⋅10-2
0 0 0
0.207 0.455 0.941
6.47⋅10-2 0.117 0.358
35
Table 3.
Flow resistance coefficients for the total pressure drop (αtotal) and the pressure drop caused by the resin bed (αresin) at a temperature of 50 °C.
Resin bed equilibrated with purified water during measurement. Determined in this work. Determined in flow rate range 1-3 m/h where linear behavior of (v,Δptotal) curve was observed.
36
Table 4.
Operating parameters of the SMB process for glucose–fructose separation and obtained values of the performance parameters. Target purities: 89 % for fructose and 90 % for glucose. dp, µm
280 µm
320 µm
320 µm
250
250
250
250
Run
1
2
3
4
5
6
7
8 **
Reference
m(run 1)*
Target Pui reached
m(run 1)*
Target Pui reached
0.596 0.384 0.570 0.281 0.186 1.75 0.374
0.596 0.384 0.570 0.281 0.186 1.75 0.369
0.615 0.410 0.540 0.281 0.13 2.57 0.365
0.596 0.384 0.570 0.281 0.186 1.75 0.361
0.596 0.365 0.595 0.281 0.230 1.37 0.361
0.596 0.365 0.595 0.281 0.230 1.37 0.361
0.596 0.350 0.610 0.281 0.260 1.21 0.361
0.596 0.330 0.630 0.281 0.300 1.05 0.361
5
5
5
5
5.793
5.793
5.793
89.0 90.1 90.3 94.6
87.5 86.7 78.8 94.0
89.9 90.7 85.0 92.8
92.1 92.6 91.9 98.3
91.4 90.5 85.7 92.8
91.3 91.8 81.2 96.0
91.2 91.3 88.3 93.9
87.6 91.0 82.3 85.3
27.2
25.7
15.9
28.6
34.9
28.4
30.8
32.1
40.0
40.6
28.4
42.1
50.7
44.9
50.2
53.7
11.28
11.97
20.07
10.88
8.65
9.25
8.48
8.14
ECG, m3/ton 7.45 7.57 11.24 7.38 CF, g/L 130.1 122.7 81.3 135.1 CG, g/L 150.6 148.1 114.7 151.9 * The same m-parameter values were used as in run 1. ** Qk was decreased (14 %) and tswitch increased (16 %) from run 5. *** The same m-parameter values were used as in run 6.
tswitch, min 5 Process performance parameters PuF, % PuG, % YF, % YG, % PrF, kg/(m3(bed) h) 3
PrG, kg/(m (bed) h) 3
ECF, m /ton
37
Table 5.
Effect of resin particle size on the energy and consumption of saturated steam (p = 2 bar) in the continuous chromatographic separation of glucose and fructose. dp, µm 250 µm 280 µm* Run 7 1 Pump power requirement (η = 0.7) P, kW 5.24 3.85
320 µm 3 2.64
X/Xref, % 136.0 100.0 69.0 Heat and steam required by the product concentration step q, kW 8890 9810 10860
*
mS, ton/h
14.5
16.0
17.8
X/Xref, %
90.6
100.0
110.7
Reference.
38
HIGHLIGHTS -
Effect of particle size on the efficiency of chromatographic separation was studied
-
Continuous glucose–fructose separation in an SMB was used as model system
-
Material and energy efficiency were evaluated
-
Material and energy efficiency improved substantially with decreasing particle size
