3096
Ind. Eng. Chem. Res. 1999, 38, 3096-3104
A Novel Process for Diethanolamine Recovery from Partially Degraded Solutions. 1. Process Description and Phase Equilibria of the DEA-BHEP-THEED-Hexadecane System Majid Abedinzadegan Abdi and Axel Meisen* Department of Chemical and Bio-Resource Engineering, The University of British Columbia, Vancouver, British Columbia V6T 1Z4, Canada
A two-part study is presented on a novel multistage distillation process for the recovery of diethanolamine from contaminated solutions. Such solutions, which pose serious problems in gas plants and oil refineries, cannot be easily purified because of the close volatilities of some of their constituents. The new process utilizes hexadecane as an inert carrier liquid to ensure proper fluid distribution in the distillation column, to avoid thermal degradation and fouling, and to facilitate separation of nonvolatile constituents. In part 1, the process is described together with new phase equilibrium data for mixtures of diethanolamine, bis(hydroxyethyl)piperazine, tris(hydroxyethyl)ethylenediamine, water, and hexadecane. The desired physical properties of the inert carrier are discussed and guidelines are presented for its selection. The nonrandom two liquid activity coefficient model was successfully used to represent the equilibrium data. Introduction Natural and refinery gases typically contain acid gases in concentrations ranging from a few parts per million to tens of volume percent.1 The major acid gases are hydrogen sulfide (H2S) and carbon dioxide (CO2). Because of the corrosivity of acid gases in the presence of water, the toxicity of H2S, and the lack of heating value of CO2, the gases must be purified prior to use. Aqueous solutions of alkanolamines react reversibly with acid gases and therefore are widely used to remove them.1 The alkanolamines of primary significance include monoethanolamine (MEA), diethanolamine (DEA), methyldiethanolamine (MDEA), di-2-propanolamine (DIPA), and diglycolamine (DGA). Although the acid gas-amine reactions are reversible, irreversible reactions may also occur, resulting in products from which the amines are not easily recovered. This phenomenon is called amine degradation.2 Partially degraded DEA solutions usually contain hydroxyethyl oxazolidone (HEOD), bis(hydroxyethyl)piperazine (BHEP), and tris(hydroxyethyl)ethylenediamine (THEED),2 together with small concentrations of higher boiling compounds.3 In addition to impurities resulting from acid gas-induced amine degradation, industrial amine solutions frequently contain other contaminants such as suspended solids, dissolved hydrocarbons, sodium and chloride ions, iron sulfide, foaming agents and foam suppressants, corrosion inhibitors, and heat stable salts. The latter are formed as a result of amine protonation by acids stronger than CO2 and H2S. The accumulation of contaminants in amine solutions creates major operational problems including foaming, fouling, corrosion, a reduction in active amine content, and an increase in solution viscosity. In the United States alone, solvent losses in alkanolamine gas- and liquid-treating plants amount to 95 million lbs/year, valued at more than $60 million/year.4 To maintain * Corresponding author. Tel.: 604-822-6708. Fax: 604-8226003. E-mail:
[email protected].
process efficiency, the circulating amine solutions must be partially or completely replaced with fresh amines. Typically, spent contaminated solutions contain over 80% of the original active amine.5 The current practices of deep well injection and incineration of the spent amine solutions not only are environmentally unattractive but also constitute a loss of valuable amine. One solution to these problems is to reclaim the contaminated solutions. However, the chemical and physical nature of the contaminants makes the development of purification schemes inherently difficult. Ideally, a purification scheme should have the following features:5 (i) high versatility, i.e., the process should be suitable for a wide range of amines; (ii) high purity, i.e., the recovered amines should be sufficiently pure so that they can be returned to the absorption process for normal service; (iii) low environmental impact, i.e., the byproducts should be considerably less in volume and less hazardous than the contaminated amine solution; (iv) low cost, i.e., the overall cost should be less than purchasing fresh amine and disposing of the contaminated amine solutions. Several techniques for the purification of degraded DEA solutions have been proposed and/or patented.6 However, most of them deal only with the removal of heat stable salts. Some preliminary work has also been reported on recovering DEA by reversing the degradation reactions.7 The purification technique, which has been the most successful so far, is distillation. However, the tendency of amines to degrade and their close boiling points with some degradation compounds preclude separation by conventional atmospheric or vacuum distillation. A lack of vapor-liquid equilibrium data for degradation compounds and mixtures including amines also hampers the design of distillation units. Canadian Chemical Reclaiming Ltd. (CCRL) has developed a process to reclaim degraded amine solutions5,8 which consists of two basic steps: caustic pretreatment followed by vacuum distillation in a flash separator. Despite operating under vacuum, there is the possibility of further thermal degradation of DEA and
10.1021/ie980454t CCC: $18.00 © 1999 American Chemical Society Published on Web 07/13/1999
Ind. Eng. Chem. Res., Vol. 38, No. 8, 1999 3097
Figure 1. Flow diagram of the novel process for purifying contaminated amine solutions.
significant BHEP concentrations in the product. Other degradation compounds such as THEED and HEOD are well-separated by the process. The lack of separation of BHEP may be explained by the fact that their volatilities are similar. Novel Process The capabilities of the CCRL process can be enhanced by replacing its single-stage flash separator with a distillation column containing several theoretical stages. Additional stages should lead to overhead and underflow products that are richer and leaner in amine, respectively. However, the concentrations of high-boiling impurities in the feed are usually too low to provide adequate liquid and vapor flows in the lower section of the column. Apart from destabilizing the column operation, this may also lead to fouling and plugging of the lower stages with nonvolatile materials. The CCRL process leaves part of the amine to carry the solids and heavier degradation products out of the flash separator. This can result in up to 20% of active amines being lost. The problems associated with multistage distillation can be overcome by introducing a suitable inert, highboiling carrier liquid. This liquid is blended with the contaminated amine feed and sent to the distillation column. By virtue of its high boiling point, the inert liquid descends to the bottom section of the column, thereby taking nonvolatile contaminants with it. If the inert liquid is selected properly, the solubility of amine contaminants is low and separation can occur gravimetrically outside the column. Upon separation, the inert liquid is reheated and mixed with the incoming contaminated amine, thereby creating a fully cyclical process. Figure 1 shows a flow diagram of the proposed process. The contaminated amine feed solution is contacted and mixed thoroughly with the hot inert liquid and sent to a separator in which the amine, water, and some of the volatile degradation products are flashed under vacuum conditions. The feed/inert liquid ratio and temperature of the flash separator are set so that a portion of the inert liquid is also volatilized. The entire solids content of the feed solution is precipitated in the flash separator and leaves with the inert liquid. The
vapor leaving the separator is partially condensed before it enters the distillation column. The inert liquid and degradation products are separated from the amine and leave the distillation column at the bottom, whereas DEA, water, and small amounts of the inert compound pass overhead. Since the solubility of the inert liquid in the amine solution is very small at the condenser temperature, it is easily separated. The inert liquid leaving the bottom of the distillation column and containing the impurities are separated in a gravity separator. The inert liquid is then recirculated. These major process modifications lead, in principle, to a high degree of amine recovery and purity. Furthermore, no significant amounts of amine come in contact with hot heat-transfer surfaces. The underflow from the gravity separator consists of the impurities present in the original contaminated amine solution together with very small amounts of inert liquid. The inert liquid should have the following principal characteristics: high chemical inertness, high boiling point, low melting point, low specific gravity, low viscosity, low solubility for amines and amine contaminants, high thermal stability, good heat-transfer characteristics, low corrosivity, low cost, and no environmental hazard.8,9 The high boiling point characteristic of the inert liquid prevents its presence in the top product of the distillation column. The low specific gravity and low mutual solubility of amine and other contaminants present in the contaminated solutions as well as chemical and thermal stability facilitate the recovery of the inert liquid. The other characteristics are typical for solvents associated with environmentally friendly processes. A knowledge of vapor pressures and liquid-liquid equilibrium data is essential for developing the process and has been reported by Abedinzadegan Abdi and Meisen.10,11 Vapor-liquid equilibrium data are also required and are presented here. Vapor-Liquid Equilibrium (VLE) Studies Two methods were used by Abedinzadegan Abdi9 to obtain the vapor-liquid equilibrium properties of mixtures containing water, DEA, and its degradation products. Devices used for determining VLE differ considerably in design, but all are based on the common principle of bringing the liquid and vapor phases into intimate contact to achieve equilibrium and ascertaining the temperature, pressure, and concentrations of one or both phases. Thermodynamics of Vapor-Liquid Equilibria Vapor and liquid phases are in equilibrium when both are at the same temperature and when the fugacity of any component i in the two phases are equal. The fugacity is usually related to other parameters such as the fugacity coefficients in the gas phase and the activity coefficients in the liquid phase. At low pressures (up to a few bar), the composition in the vapor and liquid phases can be related by the following equation:
yiP ) xiγipsi
(1)
Tables of experimental vapor-liquid equilibrium data are frequently based on eq 1.
3098 Ind. Eng. Chem. Res., Vol. 38, No. 8, 1999 Table 1. Total Pressure Measurements for the DEA-H2O System (σT ) 0.2 °C, σP ) 0.07 kPa, σx ) 0.0005) xDEA ) 0.3393
xDEA ) 0.5851
xDEA ) 0.7199
xDEA ) 0.8747
xDEA ) 0.9233
xDEA ) 0.9475
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T,°C
P, kPa
T, °C
P, kPa
35.9 44.7 54.4 64.1 71.5 79.4 84.6
3.57 6.13 9.46 14.52 19.59 26.26 31.86
36.1 45.0 54.5 64.3 74.3 84.6 94.7
1.87 3.33 5.46 8.53 12.93 19.73 27.59
36.6 45.4 54.6 64.1 74.3 84.3 94.2 104.3
1.33 2.53 4.00 6.13 9.20 13.46 19.46 25.99
39.6 58.6 74.1 89.3 104.2 116.0
0.67 2.27 4.40 7.46 12.13 17.33
49.1 59.6 69.4 79.5 89.4 99.4 109.6 119.5 129.0
1.07 2.00 3.07 4.40 6.00 8.13 10.93 14.40 18.66
55.1 64.5 74.2 84.3 94.3 132.2
0.40 0.67 1.33 2.00 3.20 13.46
For systems in which the components associate in the vapor phase, especially compounds with high intermolecular forces such as hydrogen bonds, deviations from ideal gas behavior result. Typical examples are mixtures containing water, DEA, and its degradation products, which include polar compounds undergoing hydrogen bonding. Since the compounds of interest in this study boil at elevated temperatures, and their chemical stability is of importance, the distillation should be conducted under vacuum conditions as discussed before. Under such conditions, the intermolecular forces in the vapor phase are low. Therefore, ideal gas conditions can be assumed for the vapor phase. However, because of the severe nonideality of the liquid phase, activity coefficient models should be used to predict or correlate VLE data. Many equations have been used for correlating equilibrium data (e.g., Margules, van Laar, Wilson, NRTL, and UNIQUAC). There are also major compilations of both isobaric and isothermal data.12-16 A thorough survey of the literature, including these major compilations, did not produce any vapor-liquid equilibrium data for the systems containing water, alkanolamines, and their amine degradation compounds. Systems of particular interest are aqueous DEA solutions containing one or more of its degradation compounds, such as BHEP and THEED. Experimental Apparatus and Procedures. Because of the physical and chemical nature of the compounds under study and temperature and pressure range of interest, two types of VLE cells were used. A detailed description of the experimental aspects together with vapor pressures of pure compounds are reported by Abedinzadegan Abdi9 and Abedinzadegan Abdi and Meisen.10,11 For the DEA-H2O system, a dynamic equilibrium cell was used. The equipment is described in Appendix A. For other binary systems, a static system was employed. Details of the static cell have been given by Abedinzadegan Abdi9 and Abedinzadegan Abdi and Meisen;10,11 they are briefly explained in Appendix A. Results and Discussion Binary Mixtures. The results of total pressure measurements for binary mixtures are presented in Tables 1-6 and were used to determine the nonrandom two-liquid (NRTL) adjustable parameters in the NRTL activity coefficient model by means of the ASPEN process simulator. The latter incorporates the generalized least-squares method based on the maximum likelihood principles and utilizes the Britt-Luecke algorithm.16 The nonrandomness parameter R was set to 0.3. The resulting NRTL parameters are shown in Table 7 and can be used to generate the vapor-liquid equilibrium relationships. Figure 2 shows a T-xy diagram for the DEA-H2O system.
Table 2. Total Pressure Measurements for the BHEP-H2O System (σT ) 0.2 °C, σP ) 0.07 kPa, σx ) 0.0005) xBHEP ) 0.0374 xBHEP ) 0.0580 xBHEP ) 0.1266 xBHEP ) 0.1997 T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
45.1 46.3 53.4 60.5 66.5 74.5 79.6 84.3
9.33 9.86 14.00 19.59 25.72 36.25 44.79 51.98
25.5 31.5 37.4 44.9 49.8 54.6 59.6 64.6 69.6
3.33 4.67 6.26 8.93 11.46 14.53 18.39 22.39 27.72
42.2 45.5 51.2 56.1 61.1 64.4 66.6 73.1 79.1
7.20 8.53 11.46 14.53 18.26 20.39 22.53 28.92 37.72
61.3 66.3 72.8 79.4 89.1 99.4 108.8
16.93 20.66 26.39 34.39 49.32 73.31 103.97
Despite the industrial importance of the DEA-H2O system, little fundamental thermodynamic information can be found in the literature. In a recent paper, Chang et al.18 reported on the thermodynamics of alkanolamine-water systems using a freezing-point technique. The results are obviously more accurate for low temperatures. Activity coefficients for mixtures of alkanolamines including DEA and water were also correlated by Austgen et al.19,20 using the NRTL model. Very limited experimental data for the DEA-H2O system were, however, used to determine the thermodynamic parameters. Vapor-liquid equilibrium diagrams for the DEA-H2O system are also presented in the Dow Gas Conditioning Fact Book,21 but numerical data and the experimental methods are not given. Because of the uncertainty and temperature limitations inherent in previous work and the increase in confidence, it was decided to conduct some experiments on the DEA-H2O system. Figure 2 shows the experimental measurements (obtained with both static and dynamic cells) together with the data generated by the NRTL parameters reported by Austgen et al.19 Table 8 gives the numerical values of phase equilibrium measurements using the dynamic cell at 26.7 kPa. The discrepancy between the measured data in this work and those obtained from the literature is negligible. For the BHEP-H2O system, the experiments were limited to xBHEP < 0.2 (see Table 2). The main reasons were the high degree of uncertainty in accurately determining BHEP compositions and difficulties with preparing homogeneous samples at higher concentrations.9 THEED has very low vapor pressures, with pressures of less than 30 Pa at 200 °C having been reported by Abedinzadegan Abdi and Meisen.10,11 Because of the large differences between boiling points, the phase diagrams for binary systems containing THEED exhibit very facile separation by distillation.9 Figure 3 shows the phase diagram for the DEABHEP system. Unlike the other binary systems, the equilibrium lines are close to each other. BHEP and DEA can therefore not be separated efficiently in a single flash.
Ind. Eng. Chem. Res., Vol. 38, No. 8, 1999 3099 Table 3. Total Pressure Measurements for the DEA-BHEP System (σT ) 0.2 °C, σP ) 0.07 kPa, σx ) 0.0005) xBHEP ) 0.2115
xBHEP ) 0.3041
xBHEP ) 0.5299
xBHEP ) 0.7393
xBHEP ) 0.8324
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
149 158.5 169.2 178.8 189.2 198.8
1.20 1.87 2.93 4.40 6.66 9.73
149.2 159.1 168.9 178.8 189.2 198.3
1.07 1.73 2.67 4.00 6.13 9.06
148.9 158.8 168.7 178.5 188.5 198.3 209.8
0.80 1.33 2.13 3.33 5.07 7.46 10.93
148.9 158.7 168.6 178.3 188.1 197.7 212.1
0.67 1.07 1.73 2.67 4.00 5.86 8.66
150.4 159.6 169.3 179.1 190.0 198.8 209.5 217.4
0.53 0.93 1.47 2.27 3.47 5.07 7.20 9.06
Table 4. Total Pressure Measurements for the THEED-H2O System (σT ) 0.2 °C, σP ) 0.07 kPa, σx ) 0.0005) xTHEED ) 0.1001
xTHEED ) 0.2919
xTHEED ) 0.4857
xTHEED ) 0.6963
xTHEED ) 0.8931
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
24.6 40.6 59.7 79.8 99.8
2.67 6.93 17.99 43.19 92.77
23.6 43.3 49.4 62.7 79.8 99.9
1.47 4.80 6.53 12.66 26.12 61.05
24.5 38.9 59.1 79.4 99.8 119.8
0.67 1.73 5.20 13.86 31.46 65.58
24.7 42.5 59.3 79.6 99.9 119.7
0.27 0.80 2.13 5.86 13.60 28.92
60.1 78.5 89.8 119.7 139.8
0.67 1.47 2.40 8.00 15.59
Table 5. Total Pressure Measurements for the DEA-THEED System (σT ) 0.2 °C, σP ) 0.07 kPa, σx ) 0.0005) xTHEED ) 0.0993
xTHEED ) 0.3056
xTHEED ) 0.6111
xTHEED ) 0.8055
xTHEED ) 0.9165
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
109.8 140.2 150.1 168.3 189.8 209.1
0.13 0.67 1.13 2.67 6.73 14.06
119.7 139.8 159.8 179.9 199.8 214.8
0.20 0.47 1.33 3.47 7.60 12.80
129.7 158.9 179.7 199.8 209.6
0.13 0.73 1.87 4.20 6.00
149.4 159.6 179.7 189.1 199.7 209.6
0.20 0.33 0.93 1.33 2.00 2.93
169.5 179.9 189 199.7 213.8
0.20 0.40 0.53 0.93 1.47
Table 6. Total Pressure Measurements for the BHEP-THEED System (σT ) 0.2 °C, σP ) 0.07 kPa, σx ) 0.0005) xTHEED ) 0.2037
xTHEED ) 0.4074
xTHEED ) 0.6045
xTHEED ) 0.7932
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
T, °C
P, kPa
148.7 159.1 179.7 189.8 199.9 209.8
0.33 0.53 1.33 2.00 2.93 4.13
149.2 159.2 179.6 189.3 198.3 214.7
0.20 0.40 0.93 1.47 2.00 3.60
159.6 169.4 179.5 189.4 199.9 209.5
0.20 0.40 0.67 0.93 1.40 2.00
175.0 184.9 189.9 199.9 209.7 219.6
0.20 0.40 0.53 0.73 1.07 1.47
Table 7. NRTL Bindary Interaction Parameters for Binary Systems Containing DEA, BHEP, THEED, and/or H2O NRTL parameters (ASPEN estimates) aij
σa
BHEP-DEA DEA-BHEP
-1.2522 1.9612
0.0428 0.8727
H2O-DEA DEA-H2O
-4.6697 12.4298
1.0562 2.7532
1177.12 -3685.53
BHEP-H2O H2O-BHEP
5.5313 -2.3338
1.2662 0.5288
THEED-H2O H2O-THEED
-0.1549 -1.7745
THEED-DEA DEA-THEED THEED-BHEP BHEP-THEED
binary system
σb
SSQ
RRMS
0.71×103
4.87
378.20 975.76
4.50×103
10.28
-1400.14 480.85
399.49 180.37
4.65×103
12.5
0.0015 0.1361
-854.58 1812.43
8.67 433.15
2.09×103
9.35
-2.5678 1.9890
1.5969 1.7020
-1213.54 -957.10
607.24 782.79
2.86×103
3.45
13.7973 -8.3323
5.4597 3.2896
-6080.17 585.20
2182.60 1749.42
6.08×103
17.44
For binary systems containing hexadecane, the mutual solubility data were used to estimate the vaporliquid equilibria. Azeotropes were observed for the BHEP-hexadecane and DEA-hexadecane systems. Measurements conducted by Abedinzadegan Abdi9 with a vapor circulation VLE cell using these mixtures gave results which corresponded closely to the NRTL predictions using the solubility information.10,11 Solubility tests for the THEED-hexadecane and water-hexadecane mixtures showed negligible mutual
bij 33.255 136.480
1.132 16.040
solubilities. Since the very low mutual solubility of the aforementioned systems has virtually no effect on the calculations of activity coefficients and the overall equilibrium properties of the systems under study, no further investigations were carried out. The NRTL binary parameters for these systems were estimated by the UNIFAC group contribution technique and were used in the process simulation reported in part 2. Multicomponent Mixtures. The vapor-liquid equilibrium properties of multicomponent mixtures can be
3100 Ind. Eng. Chem. Res., Vol. 38, No. 8, 1999
Figure 2. Phase (T-xy) diagram for the DEA-H2O system at 200 mmHg; comparison of data from various sources. Table 8. Vapor-Liquid Equilibrium Data for the DEA-H2O System Using the Dynamic Cell (P ) 26.7 kPa (abs); σT ) 0.2 °C, σx ) σy ) 0.0005)
Figure 3. Equilibrium (T-xy) diagram for the BHEP-DEA system at various pressures. Table 9. Fitted Values of the NRTL Binary Interaction Parameters for the Binary Systems Containing DEA, BHEP, THEED, and/or H2O; the a12 and a21 Parameters Were Estimated When b12 ) b21 ) 0 parameters
DEA mole fraction liquid phase (x) 0.0000 0.4524 0.6292 0.7766 0.9080 0.9771 1.0000
vapor phase (y) 0.0000 0.0005 0.0016 0.0062 0.0286 0.1141 1.0000
ASPEN estimates
temperature, °C 66.4 86.7 98.6 117.3 142.9 172.8 224.4
predicted using NRTL binary parameters.22 To confirm the accuracy of the method for partially degraded DEA solutions in the presence of hexadecane, three apparati were used. The static cell and dynamic still were used to measure vapor pressures and vapor-liquid compositions, respectively. A flask-total reflux condenser arrangement was employed to verify the dew point temperatures of multicomponent vapors entering the condenser. The compositions were determined by means of a gas chromatograph. The experimental results, which are presented in Table 12, indicate very good agreement between the experimental results and the equilibrium predictions based on binary parameters in the multicomponent NRTL equation. In all cases, the absolute discrepancies between the experimental data and the NRTL predictions are less than 6%. Considering the errors involved in the preparation of samples for the static cell and also the analyses of multicomponent samples by gas chromatography ((0.0005 and (0.005 in mole fractions, respectively), the observed deviations are within a satisfactory range. Regression Analysis of Vapor-Liquid Equilibrium Data. The NRTL model has three adjustable parameters, i.e., τ12, τ21, and R (see Appendix B). The nonrandomness parameter, R, in the NRTL model is normally set to a fixed value (e.g., 0.3 by default in
NRTL
aij
σa
Aij ) aijRTa
RRMS
BHEP-DEA DEA-BHEP
-0.5976 0.9842
0.0313 0.0004
-1481.50 2439.93
8.13
H2O-DEA DEA-H2O
-1.3295 2.01796
0.0737 0.1863
-3295.60 5002.67
11.82
BHEP-H2O H2O-BHEP
1.7313 -1.0464
0.2317 0.0738
4292.02 -2594.12
16.16
THEED-H2O H2O-THEED
-2.4991 2.9790
0.0140 0.2810
-6195.46 7385.19
35.74
THEED-DEA DEA-THEED
0.2021 -0.3171
0.0004 0.0209
501.02 -786.13
15.56
THEED-BHEP BHEP-THEED
-0.7904 0.9238
0.8172 1.3813
-1959.46 2290.18
19.36
binary system
a
The values were calculated at T ) 298.15 K.
ASPEN). Therefore, the number of adjustable parameters for the NRTL model reduces to two. The binary parameters are temperature-dependent, and this effect can be represented by the following equations:
τ12 )
g12 - g22 A12(T) b12 ) ) a12 + RT RT T
g21 - g11 A21(T) b21 ) ) a21 + τ21 ) RT RT T
(2)
The above equations relate the ASPEN parameters (a12, a21, b12, and b21) to the corresponding NRTL parameters (A12 and A21) published in vapor-liquid equilibrium collections such as Gmehling tables.12 The parameters Aij can be estimated from either aij or bij values using the data regression system (DRS) of the ASPEN PLUS process simulator (Aij ) bijR when aij )
Ind. Eng. Chem. Res., Vol. 38, No. 8, 1999 3101 Table 10. Fitted Values of the NRTL Binary Interaction Parameters for the Binary Systems Containing DEA, BHEP, THEED, and/or H2O; the b12 and b21 Parameters Were Estimated When a12 ) a21 ) 0 parameters ASPEN Estimates
NRTL
bij
σb
Aij ) bijRa
RRMS
BHEP-DEA DEA-BHEP
-252.2877 416.7987
12.4486 0.31243
-2098.17 3465.65
8.32
H2O-DEA DEA-H2O
-426.0847 575.0024
71.0446 155.4069
-3542.84 4781.07
46.92
BHEP-H2O H2O-BHEP
492.2916 -318.1957
41.1848 15.8283
4093.36 -2645.75
12.72
THEED-H2O H2O-THEED
-721.8277 3778.9980
3.9848 147.1007
-6001.90 31421.90
35.06
THEED-DEA DEA-THEED
-662.0205 993.9380
-5504.60 8264.46
17.74
THEED-BHEP BHEP-THEED
-491.0504 665.4100
-4083.02 5532.81
18.65
binary system
a
1.557189 54.50492 128.6425 282.6795
R ) 8.314 J/mol‚K.
aji ) O or Aij ) aijRT when bij ) bji ) O). The fitted values of the NRTL binary interaction parameters for the systems under study, calculated from individual aij and bij parameters, are shown in Tables 9 and 10, respectively. The parameters Aij were also estimated using both aij and bij. Because of the larger number of adjustable parameters and the limited number of data points, the regression resulted in larger standard deviations in the parameter estimates as shown by Table 7. Experience
has shown that reducing the number of adjustable parameters in the NRTL model gives rise to smaller values of these standard deviations. However, larger values of relative residual root mean squares of error (RRMS) may result (see Tables 9 and 10). The analysis of residuals is therefore the best way to verify whether a fit is satisfactory.23 RRMS values for the individual regressions are therefore also given in Table 7. In the maximum-likelihood method used in the data regression system (DRS) of the ASPEN process simulator, the “true” value of each measured variable is also found in the course of parameter estimation. The difference between these true values and the corresponding experimentally measured values are the residuals. The standard deviation of parameters for the THEED-DEA and THEED-BHEP systems are quite large (see Table 7). Reducing the number of parameters, from four to two, results in smaller values for standard deviations. However, it does not lead necessarily to a better fit with the experimental data, as evidenced by the larger values of RRMS (see Tables 9 and 10). Table 11 gives the measured data, estimates of the true values corresponding to the measurements and deviations of the measured values from the model predictions for the THEEDDEA system. The standard deviations of estimated parameters for these systems are the largest for the binary systems studied in this work. When the model is appropriate and there are no systematic errors, such plots show random distribution of residuals with zero means. The standard deviations of residuals are also smaller than those of the measured variables (see Table 11). A model which has more adjustable parameters and
Table 11. Measured Variables and Estimates of Their Values for the THEED-DEA System Using the ASPEN Parameters Shown in Table 7 temperature, °C
pressure, kPa
liquid composition [x]
exp. no.
meas.
calc.
dev.
meas.
calc.
dev.
meas.
calc.
dev.
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28
109.8 140.2 150.1 168.3 189.8 209.1 119.7 139.8 159.8 179.9 199.8 214.0 129.7 158.9 179.7 199.8 209.6 149.4 159.6 179.7 189.1 199.7 209.6 169.5 179.9 189.0 199.7 213.8
109.81 140.22 150.19 168.40 190.04 209.27 119.73 139.76 159.62 180.17 199.83 213.98 129.69 159.91 179.65 199.72 209.39 149.38 159.56 179.67 188.96 199.39 209.30 169.47 179.88 188.91 199.73 213.66
0.01 0.02 0.09 0.01 0.24 0.17 0.03 -0.04 -0.18 0.27 0.01 -0.02 -0.01 -0.01 -0.05 -0.08 -0.21 -0.02 -0.04 -0.03 -0.14 -0.31 -0.30 -0.03 -0.02 -0.09 0.03 -0.14
0.13 0.67 1.13 2.67 6.73 14.06 0.20 0.47 1.33 3.47 7.60 12.80 0.13 0.73 1.87 4.20 6.00 0.20 0.33 0.93 1.33 2.00 2.93 0.20 0.40 0.53 0.93 1.47
0.11 0.66 1.10 2.65 6.72 14.06 0.16 0.49 1.36 3.43 7.58 12.82 0.16 0.74 1.88 4.21 6.01 0.23 0.38 0.94 1.38 2.07 2.98 0.26 0.41 0.60 0.91 1.51
0.02 -0.01 -0.04 -0.02 -0.02 -0.01 -0.04 0.03 0.02 -0.03 -0.01 0.03 0.03 0.00 0.01 0.01 0.02 0.03 0.05 0.00 0.05 0.07 0.05 0.06 0.01 0.07 -0.02 0.04
0.0993 0.0993 0.0993 0.0993 0.0993 0.0993 0.3056 0.3056 0.3056 0.3056 0.3056 0.3056 0.6111 0.6111 0.6111 0.6111 0.6111 0.8055 0.8055 0.8055 0.8055 0.8055 0.8055 0.9165 0.9165 0.9165 0.9165 0.9165
0.09930 0.09929 0.09929 0.09930 0.09932 0.09934 0.30558 0.30560 0.30563 0.30554 0.30565 0.30570 0.61110 0.61110 0.61112 0.61112 0.61113 0.80551 0.80551 0.80548 0.80553 0.80560 0.80559 0.91650 0.91649 0.91650 0.91646 0.91649
4.8 × 10-6 -4.1 × 10-6 -4.4 × 10-6 3.4 × 10-6 1.5 × 10-5 3.7 × 10-5 -2.0 × 10-5 -1.1 × 10-6 2.7 × 10-5 -6.2 × 10-5 5.3 × 10-5 1.0 × 10-4 4.1 × 10-6 -1.2 × 10-6 2.1 × 10-5 2.1 × 10-5 1.5 × 10-4 1.0 × 10-5 1.4 × 10-5 -1.6 × 10-6 3.0 × 10-5 1.0 × 10-4 8.6 × 10-5 -1.1 × 10-6 -9.9 × 10-6 -3.1 × 10-6 -3.6 × 10-5 8.8 × 10-6
avg. avg. Abs. max. std. dev. of res’d. (σE) max./σEa a
Maximum standardized residua.
-0.03 0.09 -0.31 0.13 2.38
0.01 0.03 0.07 0.03 0.30
1.8 × 10-6 3.0 × 10-5 1.5 × 10-4 4.5 × 10-5 3.3
3102 Ind. Eng. Chem. Res., Vol. 38, No. 8, 1999 Table 12. Comparison of Experimental Results and Multicomponent NRTL Predictions for DEA, BHEP, THEED, H2O, and/or Hexadecane a. Static Cell Results: xDEA ) 0.4681, xBHEP ) 0.0307, xTHEED ) 0.0484, xH2O ) 0.4528 bubble point temp., °C meas.
calc.
pressure, kPa
error, °C
101.1 123.3
104.5 116.4
26.67 40.00
3.4 -6.9
b. Dynamic Still Results: xDEA ) 0.2028, xBHEP ) 0.0142, xTHEED ) 0.0195, xH2O ) 0.7635; yDEA ) 0.00, yBHEP ) 0.00, yTHEED ) 0.00, yH2O ) 1.00 equilibrium temp., °C meas.
calc.
pressure, kPa
error, °C
75.0 86.1
73.1 89.3
26.67 40.00
-1.9 3.2
c. Flask and Condenser Set-Up Results (exp. no. 1): yDEA ) 0.212, yBHEP ) 0.01, yHexadecane ) 0.778 dew point temp., °C meas.
calc.
pressure, kPa
error, °C
132.9
125.5
0.67
7.4
d. Flask and Condenser Set-Up Results (exp. no. 2): yDEA ) 0.2479, yBHEP ) 0.0111, yHexadecane ) 0.7410 dew point temp., °C
A ) NRTL model parameter b ) temperature coefficient in temperature-dependent equations for adjustable parameters in activity coefficient models (ASPEN parameters) f ) fugacity g ) binary interaction energy in the NRTL model G ) NRTL model parameter P ) total pressure RRMS ) relative residual root mean square R ) gas constant SSQ ) sum of squares of errors T ) temperature v ) molar volume V ) volume x ) mole fraction in liquid phase y ) mole fraction in vapor phase Greek Letters R ) nonrandomness parameter in the NRTL model φ ) fugacity coefficient γ ) activity coefficient σ ) standard deviation of parameter estimate τ ) NRTL binary interaction energy parameter Superscripts L ) liquid phase V ) vapor phase Subscripts
meas.
calc.
pressure, kPa
error, °C
154.0
152.6
2.67
1.4
gives a good fit to the experimental data may be safely used for predictions within the range of observed values. Therefore, the RRMS values were used as the criterion to select the model and verify the quality of the fit. The analyses of residuals for the other binary systems show similar results and therefore prove that NRTL is a suitable model and that regression of the experimental data is satisfactory. Conclusions The vapor-liquid equilibrium results obtained in this study reveal that the separation of DEA degradation products (i.e., BHEP and THEED) from aqueous DEA is feasible by distillation. Because of its high boiling point, THEED can be separated efficiently in a singlestage flash process. However, BHEP has a volatility which is close to that of DEA and therefore requires more stages for its separation. The presence of azeotropes in mixtures consisting of hexadecane, DEA, BHEP, and water results in some hexadecane leaving with the overhead product in the proposed process. However, because of the low mutual solubility of hexadecane and DEA at near ambient temperatures, hexadecane can be readily separated by gravity and recycled. Acknowledgment The financial support provided by the Natural Sciences and Engineering Research Council of Canada and the National Iranian Oil Co. for this work is gratefully acknowledged. Nomenclature a ) constant in temperature-dependent equations for adjustable parameters of activity coefficient models (ASPEN parameter)
i,j ) species i,j 1,2 ) components in binary systems P ) pressure T ) temperature x ) liquid phase y ) vapor phase Acronyms of Chemical Compounds BHEP ) bis(hydroxyethyl)piperazine DEA ) diethanolamine DGA ) diglycolamine DIPA ) di-2-propanolamine HEOD ) hydroxyethyloxazolidone MDEA ) methyldiethanolamine MEA ) monoethanolamine THEED ) tris(hydroxyethyl)ethylenediamine
Appendix A Materials. BHEP with a purity exceeding 99% was supplied by Aldrich Chemical Co. (Milwaukee, WI) and was used without further purification. THEED was specially prepared with a purity exceeding 95% by Abedinzadegan Abdi.9 Its primary impurities were BHEP and DEA. DEA and hexadecane were both supplied by Aldrich with purities exceeding 99.5%. Experimental Procedure and Apparatus. Two different techniques were used to perform the experiments: a static cell arrangement to measure the total pressure of various mixtures and a dynamic still to determine the vapor-liquid equilibria of binary and multicomponent systems. The solutions were prepared gravimetrically using an electronic balance with at least three significant figures after the decimal point. The compositions for the DEAH2O and BHEP-H2O systems were verified using a refractometer (model ABBE-32, Milton Roy, NY) carefully calibrated with standard solutions. The accuracy of the measurements was within (0.0001 of the refractive index and the compositions in mole fractions were determined with a precision better than (0.0005. The
Ind. Eng. Chem. Res., Vol. 38, No. 8, 1999 3103
Figure 4. Static cell for measuring the total system pressures.
gas chromatograph technique used for the analysis of DEA-H2O mixtures proved to be less precise as the best standard deviation obtained for this technique was below (0.005 in mole fractions. The compositions of the mixtures were measured once before and after each experiment, and the absolute differences between the two measurements were less than 0.05%. For the THEED-H2O and other binary mixtures containing THEED, BHEP, and DEA, the initial charges to the static cell were prepared gravimetrially and reported without any further verification by other analytical techniques. The test samples were degassed according to the method proposed by Van Ness and Abbott.24 The system consisted of a rectifying column (10-mm i.d.) packed over a height of 200 mm with glass beads, 2-3-mm diameter. The mixtures of interest were distilled under total reflux and low pressures for at least 2 h to ensure complete removal of all highly volatile components. The pressure sensor was calibrated by linearization between zero (high vacuum) and atmospheric pressures measured with a barometer (model PRINCO, Fortin Type, Precision Thermometer & Instrument Co., Philadelphia, PA) with an accuracy of (1 Pa. Static Cell. In this method the solution is charged into a closed vessel which is placed in a thermostat. The solution is stirred until equilibrium is attained between the liquid and vapor phases. The total pressure exerted by the feed mixture of known composition is measured with the pressure sensor. The total pressure measurement apparatus is shown in Figure 4. Since there is no vapor mixing by either refluxing or vapor recycling, equilibrium is established by component transfer across the large liquid volume/small vapor volume interface. The liquid was mixed and the interface was agitated by a Teflon-coated stirrer bar. The equilibrium cells, which were either 25- or 50-mL flat-ended flasks, were sealed with glass ball-and-socket joints. In typical experiments, a cell was charged with liquid, almost to its total capacity, to ensure a small vapor space and therefore a liquid composition very close to that of the original charge. By appropriate mixing of the liquid phase, equilibrium could be reached in typically less than 15 min.
Figure 5. Dynamic equilibrium still for determining the equilibrium compositions.
The static method is simple in design, but degassing of the liquid, particularly high-boiling compounds, is very important.9 Dynamic Still. Equilibrium stills are suitable for systems with heat-sensitive components and have the advantage that they avoid recirculation of the condensed vapor; they yield reasonably precise results in short times (∼15 min).15,25 The still used in this work differed from those described previously because the vacuum jacket was extended to include the droplet separator (see Figure 5) to minimize thermal losses. The liquid and vapor sampling lines were also changed such that free flow of viscous liquid mixtures could be achieved. Because of large differences between the boiling points in some systems (e.g., DEA-H2O), a large amount of liquid was collected before equilibrium could be reached. Therefore, a 100-mL liquid collector was provided. Rotating glass junctions were used to enable easy sampling from liquid and vapor streams. The starting liquid (binary or multicomponent) was introduced into the reservoir. The float mechanism immediately establishes a constant level in the vaporizing tube where the preheated liquid is brought to ebullition with the aid of an electric heating tape. The liquid in the heating tube is propelled by the rising vapor. A mixture of liquid and vapor enters the separator and drains to a droplet separator. The vapor is separated from the liquid in the droplet separator, and after having been totally condensed in a cooler, it drains to a condensate collector. The liquid, which is discharged by a siphon, is cooled and collected. Appendix B: NRTL Equation and Related Parameters
[( [(
) ( ) (
ln γ1 ) x22 τ21
G21 x1 + x2G21
2
ln γ2 ) x12 τ12
G12 x2 + x1G12
2
where
+
+
τ12G12
)] )]
(x2 + x1G12)2 τ21G21
(x1 + x2G21)2
3104 Ind. Eng. Chem. Res., Vol. 38, No. 8, 1999
τ12 )
g12 - g22 RT
G12 ) exp{-R12τ12}
τ21 )
g21 - g11 RT
G21 ) exp{-R21τ21}
The NRTL parameters are related to the ASPEN “a” and “b” parameters by eq 4. Literature Cited (1) Kohl, A. L.; Riesenfeld, F. C. Gas Purification, 4th ed.; Gulf Publishing Co.: Houston, TX, 1985. (2) Kennard, M. L.; Meisen, A. Mechanisms and Kinetics of Diethanolamine Degradation. Ind. Eng. Chem. Fundam. 1985, 24 (2), 129-40. (3) Hsu, C. S.; Kim, C. J. Diethanolamine (DEA) Degradation under Gas Treating Conditions. Ind. Eng. Chem. Prod. Res. Dev. 1985, 24, 630-35. (4) Stewart, E. J. Systematic Technical Approach to Reducing Amine Plant Solvent Losses. Proceedings of the 41st Annual Laurance Reid Gas Conditioning Conference; Oklahoma University: Norman, OK, 1991; pp 60-95. (5) Dawodu, O. F.; Meisen, A.; Beasley, T. Reclamation of Spent Amine Solutions. AIChE Spring National Meeting, March 28April 1, 1993, Houston, TX. (6) Bedell S. A. U.S. Patent, 4,814,051, 1989 and 4,808,284, 1989. (7) Chakma, A. Studies on DEA and MDEA Degradation. Ph.D. Thesis, University of British Columbia, British Columbia, 1987. (8) Meisen, A.; Abedinzadegan, A.; Abry, R. G.; Millard, M. G. Degraded Amine Solutions; Nature, Problems and Distillative Reclamation. Proceedings of the 45th Annual Laurance Reid Gas Conditioning Conference; Oklahoma University: Norman, OK, 1996; pp 168-189. (9) Abedinzadegan Abdi, M. Purification of Partially Degraded Diethanolamine Solutions. Ph.D. Dissertation, University of British Columbia, Vancouver, British Columbia, Canada, 1997. (10) Abedinzadegan Abdi, M.; Meisen, A. Vapor Pressure Measurements of Bis(hydroxyethyl)piperazine and Tris(hydroxyethyl)ethylenediamine. J. Chem. Eng. Data 1998, 43, 133-137. (11) Abedinzadegan Abdi, M.; Meisen, A. Mutual Solubility of Hexadecane with Diethanolamine and with Bis(hydroxyethyl)piperazine. J. Chem. Eng. Data 1998, 43, 138-142.
(12) Gmehling, J.; Onken, U.; Arlt, W. Vapor-Liquid Equilibrium Data Collection. DECHEMA Chemistry Data Series; DECHEMA: Frankfurt, Germany, 1980; Vol. 1. (13) Hirata, M.; Ohe, S.; Nagahama, K. Computer Aided Data Book of Vapor-Liquid Equilibria; Elsevier: Amsterdam, 1975. (14) Wichterle, I.; Linek, J.; Hala, E. Vapor-Liquid Equilibrium Data Bibliography; Elsevier: Amsterdam, 1973. (15) Ha´la, E.; Pick, J.; Fried, V.; Vilim, O. Vapor-Liquid Equilibrium, 2nd ed.; Pergamon: Oxford, 1967; Part 3. (16) Britt, H. I.; Luecke, R. H. The Estimation of Parameters in Nonlinear, Implicit Models. Technometrics 1973, 15 (2), 25970. (17) Chu, J. C.; Wang, S. L.; Levy, S. L.; Paul, R. Vapor-Liquid Equilibrium Data; Edwards Inc.: Ann Arbor, MI, 1956. (18) Chang, H. T.; Posey, M.; Rochelle, G. T. Thermodynamics of Alkanolamine-Water Solutions from Freezing Point Measurements. Ind. Eng. Chem. Res. 1993, 32, 2324-2335. (19) Austgen, D. M.; Rochelle, G. T.; Chen, C. C. Model of Vapor-Liquid Equilibria for Aqueous Acid Gas-Alkanolamine Systems. 2. Representation of H2S and CO2 Solubility in Aqueous Mixtures of MDEA with MEA or DEA. Ind. Eng. Chem. Res. 1991, 30, 543-555. (20) Austgen, D. M.; Rochelle, G. T.; Peng, X.; Chen, C. C. Model of Vapor-Liquid Equilibria for Aqueous Acid Gas-Alkanolamine Systems Using the Electrolyte-NRTL Equation. Ind. Eng. Chem. Res. 1989, 28, 1060-1073. (21) Anon. Gas Conditioning Fact Book; The Dow Chemical Co.: Midland, MI, 1962. (22) Prausnitz, J. M.; Lichtenhaler, R. N.; deAzevedo, E. G. Molecular Thermodynamics of Fluid Phase Equilibria, 2nd ed.; Prentice Hall: Englewood Cliffs, NJ, 1986. (23) Prausnitz, J. M.; Anderson, T. F.; Grens, E. A.; Eckert, C. A.; Hsieh, R.; O’Connell, J. P. Computer Calculations For Multicomponent Vapor-Liquid and Liquid-Liquid Equilibria; Prentice Hall: Englewood Cliffs, NJ, 1980. (24) Van Ness, H. C.; Abbott, M. M. A Procedure for Rapid Degassing of Liquids. Ind. Eng. Chem. Fundam. 1978, 17 (1), 6667. (25) Dawe, R. A.; Newsham, D. M. T.; Bee Ng, S. Vapor-Liquid Equilibria in Mixtures of Water, n-Propanol, and n-Butanol. J. Chem. Eng. Data. 1973, 18 (1), 44-49.
Received for review July 13, 1998 Revised manuscript received April 27, 1999 Accepted April 27, 1999 IE980454T
