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Operations & Maintenance
Distillation Reboiler Startup Can Pose Challenges Future designs can benefit from the lessons that were learned via a well-planned investigation of this startup problem
A Clean solvent
FIGURE 1. The extractive distillation tower discussed in this article employs two reboilers
Feed 93% A 7% B
108°C 109°C
Thermosiphon reboiler
Kettle hs, REB reboiler
#3 ∆Pout 140°C 1
Steam
5 3
∆Pin
H
2
#1 109°C #2 140°C
4
Contaminated solvent 2% A 1% B 97% solvent
Etienne Rubbers, Kirsten Green and Terry Fowler, Sasol Technology (Pty.) Ltd. Henry Z. Kister and Walter J. Stupin, Fluor
D
uring the initial startup of a new extractive distillation tower* at a Sasol plant in Secunda, South Africa, the chimney tray feeding liquid to the tower’s once-through kettle reboiler unexpectedly overflowed. The plant quickly implemented a successful fix, but the cause of the overflow remained obscure. Further joint troubleshooting by a Sasol/Fluor team cleared up the mystery, and provided lessons for how to avoid similar problems in other extractive distillation units. The investigation utilized pressure drop measurements, neutron backscatter, surface temperature surveys, and hydraulic calculations to establish the force balance that led to the overflow. Among other things, the measurements offered strong evidence of boiling maldistribution in the reboiler. The team came up with a theory that is consistent with all the measurements and force balance; it subdivides the reboiler into a stagnant region, an intense boiling region, and a kettle region. This regional maldistribution pattern is believed to be unique to mixtures such as those encountered in ex*The tower was not designed by the authors’ employers.
TABLE 1. DESIGN STREAM COMPOSITIONS Stream #1 Stream #2 Liquid to Liquid Flashed Flashed kettle ex kettle Temperature,°C 109 124 137 140 A, wt, % 30 3 2 2 B, wt, % 14 2 1 1 C, % wt 56 95 97 97 Vaporization, mol% — 69.0 71.7 72.3 Mass flow, kg/h 134,000 72,000
tractive distillation, with the boiling liquid consisting of a large concentration of very high boilers and a small concentration of volatile components. The mulitifeed and multidraw kettle arrangement, and possible foam formation, may also play a role.
Process description The process feed to the tower (Figure 1) contains: 93% Component A, an organic with an atmospheric boiling point of 80°C, which is recovered as a high-purity overhead product from the top of the tower; and 7% Component B, an organic that boils at 84°C under atmospheric pressure. B has a higher affinity to the solvent and leaves at the tower bottom. The solvent boils at above 250˚C at atmospheric pressure. In the unusual boilup supply system for the tower (Figure 1), liquid at 108˚C from the bottom trays is collected on an upper chimney tray, from where it gravity-flows to a oncethrough thermosiphon side-reboiler. The liquid from the outlet of the thermosiphon is collected on a lower chim-
Stream #3 Vapor ex kettle 140 62 29 9 — 62,000
ney tray at 109˚C, and gravity-flows to the kettle reboiler. Vapor generated in the kettle reboiler at 140˚C is returned to the tower above the lower chimney tray. Unboiled liquid from the kettle reboiler, also at 140˚C, overflows the reboiler weir into the kettle draw compartment, from where it gravity-flows into the tower bottomproduct sump. Table 1 shows the rapid change in component concentration that takes place near the bottom of the tower. The tower bottoms contain 2% by weight of Component A, 1% of Component B, and 97% solvent (Component C). One stage up, at the bottom chimney tray, the liquid composition changes to 30 wt.% A, 14 wt.% B, and only 56 wt.% solvent. This difference is the reason for the steep temperature change near the tower bottom. Such steep composition and temperature changes are typical of extractive distillations, using a high-boiling solvent to separate relatively low-boiling organics. (Continues on p. 56)
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The startup problem Upon startup, the lower chimney tray operated at the design temperature of 109˚C. The kettle reboiler vapor was at 145˚C, somewhat hotter than the design 140˚C. However, the bottom sump temperature was 109˚C, much colder than the design temperature of 140˚C. This low temperature meant a much larger concentration of A (around 30 wt.%) in the sump, compared to the design (around 2 wt.%). Because of the discrepancy, the recovery of A in the product stream was much poorer than the design recovery. What’s more, flooding was apparent in the solvent stripper downstream of the extractive distillation tower, due to the significant variation in its feed composition. Operators first tried to improve performance by manipulating some of the tower operating variables. Raising the steam rate as far as possible only had a small effect, raising the sump temperature from 109 to 112˚C. Reducing the feed flowrate had a much greater effect, raising the sump temperature as high as 130˚C, but at the penalty of lower production rates. The situation was not tolerable, and changing operating variables did not produce a satisfactory solution. The higher reboiler temperature suggests that the flowrate through the kettle reboiler was lower than as designed. The only plausible explanation for the low sump temperature is that the liquid from the lower chimney tray was bypassing the kettle reboiler and reaching the bottom sump. Such bypassing can be due either to liquid leaking from the lower chimney tray, or to liquid overflowing into the chimneys. The observation that the sump temperature rose at lower feed rates argued against tray leakage being the root cause, thus supporting liquid overflow into the chimneys. There was another possible route for liquid to bypass the lower chimney tray. Vapor-liquid effluent from the thermosiphon side reboiler enters the vapor space above the lower chimney tray (Figure 1). Though the bottom of this nozzle is 300 mm above the hat of the chimneys, it was possible that liquid trajectories from the side reboiler effluent were finding their way into a chimney that had its opening oriented towards the effluent nozzle. 56
FIGURE 2. Temperature measurements made during troubleshooting of the kettle reboiler constituted a key input for the analysis of its performance problem
Based on the overflow theory, the tower was shut down after one day in service, and the chimney heights on the lower chimney tray were extended from 200 mm to 900 mm. Furthermore, the chimney into which the side reboiler could have been blowing liquid was blanked. Upon return of the system to service, the sump temperature reached the design 140˚C and stayed there during all periods of normal operation. The reboiler temperature declined from 145˚C to the design 140˚C. The chimney tray modifications were successful in preventing any further liquid bypassing from the lower chimney tray into the bottom sump.
Taking the pulse Although the problem thus went away, understanding its root cause posed an interesting challenge. Furthermore, such understanding would be valuable, both for dealing with similar systems and for application during any future debottlenecking of the same Sasol distillation unit. So, the following checks and measurements were performed: Pressures: The team measured pressures at the vapor space above the kettle reboiler and in the vapor above the lower chimney tray in the tower. The measured pressure drop at startup was 4–8 kPa, higher than the design 2.5kPa maximum, even though the flowrates were lower. The outlet-line pressure drop was recalculated, using three alternative procedures: a Fluor method, a commercial simulation, and a simple hand method based on Perry’s Handbook [2]. All three methods gave numbers in the 4–6-kPa range, which were slightly lower than, yet well in line with, the field measurements Liquid level: Neutron backscatter scans were performed to determine
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the liquid level on the lower chimney tray. The level was found at 300 mm. This finding confirmed that, with the initial 200-mm chimneys, the liquid level would have been above the top of the chimneys and overflow would have occurred. It also confirmed that increasing the chimney height to 900 mm would have indeed eliminated the overflow. Thus, this test fully confirmed the startup team’s analysis and solution. The neutron backscatter scans also argue against the liquid trajectories from the side reboiler causing the bypassing. With the trajectories fully eliminated during the startup fix, the chimney tray liquid level still exceeded the original chimney height of 200 mm. Entrainment: Entrainment of liquid in the vapor leaving kettle reboilers has been known to occur when the vapor space above the tube bundle is small [1]. To explore this possibility, the team assessed the design of the existing kettle in light of good-design criteria listed in the literature [1]. The ratio of kettle diameter to overflow baffle height in the existing kettle is 1.4, which is near the recommended minimum (1.3 to 1.6). The ratio of tube bundle to kettle diameter is 0.63, which slightly exceeds the recommended maximum 0.6. Foaming: The team conducted some limited tests to check for foaming. The tray-pressure drop across the column did not provide any evidence of foaming on the trays. So, if foaming took place, it would be confined to the kettle, where the solvent concentration was high and where vigorous boiling takes place. A sample of reboiler liquid was shaken, and was observed to produce about 25 mm of froth above the liquid, which took 1–2 min. to disappear. Because of the high fluid tempera-
tures, it was not possible to simulate the operating conditions. So foaming could have played a role, even though no stable foam was observed. Oldershaw tests [1] during the piloting of the column indicated some wall-supported foaming, but the foam was unstable. Temperatures: Laser-guided infrared pyrometers measured the surface temperatures around the kettle reboiler shell. Several sets of these measurements were made. While the temperature varied from one set of readings to another due to changes in steam pressure, the trends observed did not vary and were completely repeatable. Figure 2 shows the results of the most comprehensive temperature survey. The kettle reboiler has four liquid inlets and four vapor outlets, a configuration that is often used to provide good fluid distribution in horizontal shell-side reboilers. The average surface temperatures at the two vapor outlet nozzles furthest away from the overflow baffle were low (122 and 128°C). The surface temperature at the next vapor-outlet nozzle was 141˚C, and at the vapor nozzle closest to the overflow baffle, the temperature was 144˚C. The surface temperatures at the reboiler liquid pool displayed the opposite trend. Temperatures were high (176 and 186°C) near the flange, and were progressively lower along the final half of the reboiler length. The surface temperatures at the liquid pool were 160–168°C in the third quadrant away from the overflow baffle, 157–160°C in the second quadrant away from it, and 154°C at the quadrant just before the overflow baffle. Heat transfer: The heat transfer coefficient was calculated to be 839 kW/(m2)(˚C), which is lower than the design 1,026 kW/(m2)(˚C).
Assessing reboiler operation The measurements and determinations made as outlined above enabled the team to assess the operation within the reboiler. Pressure balance: The pressure balance between the elevation at Point 1 at the vapor space above the lower chimney tray and that at Point 2 at the reboiler floor (see Figure 1) is as follows (with all pressures expressed
in millimeters of liquid head): H = ⌬Pin + hs,REB + ⌬Pout
(1)
where H is the liquid driving head (top of chimney tray liquid level to reboiler floor), ⌬Pin is the pressure drop in the inlet piping, hs,REB is the static head in the reboiler and ⌬Pout is the pressure drop in outlet piping. The neutron scan measurements gave H as 15.5 kPa at normal operating conditions. The ⌬Pin was calculated at about 1 kPa. The ⌬Pout at normal operating conditions was calculated at 5–7 kPa; this range was slightly higher than the startup ⌬Pout, which was 4–6 kPa and agreed with the plant data at its original startup. With the assumption of no aeration in the reboiler liquid, and allowing for the head over the weir as recommended [3], hs,REB was calculated at about 8 kPa. The right side of Equation (1) thus gives 14-16 kPa, which is slightly on the low side (but well within calculation accuracy) of the 15.5 kPa head on the left side of Equation (1). Basic to this force balance is the assumption that the static head inside the reboiler, hs,REB, is equal to the head of actual liquid from the top of the overflow baffle (plus the small head over the weir) to the reboiler floor. This assumption makes hs,REB the head of non-aerated liquid, which may be questionable for a pool of boiling liquid. This assumption is made in some key literature references on the subject [3, 4], and is often the conservative assumption made for design. More-comprehensive models [5, 6] replace the non-aeration assumption by a model that takes into account the generated vapor volume, as well as the tube-bundle two-phase pressure drop. Application of these models to kettle reboilers is further complicated by the recirculation of liquid inside the kettle and the existence of fairly clear liquid zones between the bundle and the sides of a kettle reboiler [5]. In the Sasol reboiler, there is an additional justification for the non-aeration assumption. As is discussed below, there is evidence that the nozzle closest to the tubesheet enters a stagnant liquid region. In that case, this inlet will see the full hydrostatic head between the overflow (Point 3) and the
reboiler floor (Point 2) in Figure 1. Based on the Equation (1) force balance, the chimney tray overflowed because the pressure drop in the reboiler outlet line was high and because the full static head of the non-aerated reboiler liquid acted to raise the driving head on the chimney tray. Entrainment from the kettle could have also contributed. This entrainment would further increase the pressure drop in the reboiler vapor outlet line, and account for the outlet line pressure-drop measurement being slightly higher than calculated. Calculations show that about 10-30% entrainment in the vapor would account for the slight pressure drop difference. However, this difference can also be explained by simple inaccuracies in measurement and calculation. Although the pressure balance explains the chimney tray overflow, it is unable to explain the observations from our temperature measurements. These are discussed below. Vapor disengagement between tubesheet and first vapor outlet nozzle: The tubes are 6,100 mm long (this measurement does not appear in Figure 2). Of this length, about 1,500 mm (or 25%) lies between the tubesheet and the beginning of the first vapor outlet nozzle. So if the heat exchanger were to vaporize the liquid uniformly, practically all the vapor exiting at the first vapor outlet nozzle should initiate in the exchanger section between the tubesheet and the beginning of that nozzle. The kettle overflow baffle is 1,150 mm high. A Francis weir formula calculation [2] shows that a 72,000-kg/h liquid overflow over the outlet weir will incur an additional liquid head of 50 mm. So, the liquid level is at least 1,200 mm above the kettle floor (without allowance for hydraulic gradients or for aeration of the liquid, which will tend to make the liquid height even taller). The first 236 mm of tubes are in the 1,000-mm-dia portion of the kettle, with no vapor disengagement space. Over the next 1,090 mm of the tubes, the kettle diameter expands from 1,000 to 1,628 mm, at a 30-deg angle. Geometry dictates that the first 350 mm of that expansion will be fully submerged by the 1,200 mm of liquid. So the total tube length at which there
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is no vapor disengagement whatsoever is about 590 mm. In this length, 40-50% of the vapor reaching the first vapor nozzle should be generated. Further, in light of the expected vapor generation, vapor velocities should be locally high. And, there will not be much disengagement of liquid in the next 500 mm or so of tube length, as this section has a restricted vaporpassage area due to the smaller diameter. (However, as we have concluded below, the actual vapor generation in this section is very low.) Stagnant liquid region (Figure 3): For the liquid in the tubesheet region, the surface temperatures measured were high. This is also the region that has little disengagement space above. Liquid temperatures in the range of 168 to 186°C imply concentration of Components A and B of less than 0.5% by weight; in other words, very low. Under the assumption that these temperatures apply to the mixture close to the tubesheet throughout the tube bundle, the material in this region thus consists almost entirely of the non-volatile solvent, Component C. It then follows that in this region, very little boiling and heat transfer take place. This absence will occur if there is only a small flux of material flowing around the tubes in this area. The small flow generates only a small quantity of vapor, and the liquid around the tubes will get hotter and hotter as its low-boiling components flash off. All that will be left is the heavy liquid, which tends to sit there. This behavior turns most of the first quadrant into a stagnant zone or dead pocket of hot, heavy liquid, mostly Component C. It is worth noting that temperatures of 168 to 186°C were surface measurements at the shell, some distance away from the hot tubes. The stagnant layer between the tubes and shell is a good insulator, and it is possible that near the tubes the temperatures will be higher, probably approaching that of the condensing steam. Since not much liquid vaporizes in the stagnant zone, the froth density in this region is high, much closer to that of the liquid than to that of a vapor/liquid mix, and much higher than that of the froth or aerated liquid in the other 58
FIGURE 3. Boiling behavior varied significantly across regions of the kettle reboiler
quadrants of the kettle. This means that there is less flow resistance, due to the lower static head of the froth in the central and baffle regions. Consequently more incoming liquid will channel toward the central region. Our calculations indicate only a few hundred millimeters of head difference are enough to account for the essential absence of flow in one feed nozzle while the adjacent nozzle feeds the bulk of the liquid to the reboiler. As less feed flows towards the tubesheet end, the liquid in the tubesheet region will stagnate more and vaporize less. The end result is a self-accelerating mechanism that promotes the dead zone in the tubesheet region and a high-liquid-flux central region. The high-L/V, intense boiling, central region (Figure 3): This region has a higher-than-design liquid feedrate (and a high liquid/vapor, or L/V, ratio), to compensate for the reduced flow to the stagnant region. The high rate results in lower liquid temperatures, but with intense boiling and higher than design vapor generation. A pressure balance across the reboiler (Figure 1) equates the pressure difference between the common liquid inlet line (Point 4) and the common reboiler vapor space (Point 3), no matter which region one travels through. The pressure difference is given by Equation (2), again with all pressures expressed as millimeters of liquid head: P4 - P3 = hs + Pn + PREB
(2)
where hs is the static head in a given reboiler region, Pn is the pressure drop in the liquid line leading to the region, and PREB is the reboiler pressure drop in that region. In the stagnant region discussed above, hs approaches the full head of the non-aerated liquid, and the P terms are quite small due to the low flow. Con-
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versely, in the intense boiling region, the boiling aerates the liquid, which lowers the liquid head. This induces high flowrates of liquid into the region until the pressure-drop terms balance the difference between the aerated-liquid head in this region and the non-aerated-liquid head in the stagnant region. The high liquid flowrate induced into the vigorous-boiling central region means that the fraction of the liquid vaporized in this region is lower, so less of the low-boiling components are removed from the liquid. This lights-rich liquid could possibly persist all the way to the two vapor nozzles closest to the tube sheet. The lower temperatures observed near the vapor nozzles closer to the tube sheet are the result of the high liquid flux in the tube field and the resulting lower temperatures through the regions providing the vapor to these nozzles. The vapor nozzle closest to the tubesheet is the coldest, because that is where the intense boiling starts and is more vigorous, and there is a smaller fraction vaporized. Also, some of the liquid is likely to be projected into the vapor outlet nozzles in the form of entrainment. The action in this region has some similarities to the action in horizontal thermosiphon reboilers, in that the quantity of vapor generated results in a low-density froth, which promotes liquid flow to this region, in contrast to the behavior with the more dense liquid in the more stagnant zones. One might wonder what caused the measured wall temperatures of the liquid in this region to remain hot. The key is that our survey measures the wall temperatures, not the bulk temperatures. The intense boiling takes place at the bulk of the bundle, and may not reach the wall. Equation (2) shows that along the wall in this region, there is horizon-
tal flow of hot liquid from the higherstatic-head stagnant region towards the overflow baffle. This flow is somewhat tempered by the boiling. On the other hand, the vapor temperatures were measured right above the bulk of the bundle, where the boiling is most intense. The kettle boiling region (Figure 3): There is a net flow of liquid from the central region to the overflow baffle, because this is the only manner in which the solvent can permanently exit the kettle reboiler system. As one moves closer to the overflow baffle, there are two effects: this net flow increases, and the composition of the liquid can be expected to contain less of low-boiling material. As a result, less vigorous boiling and a higher aerated-liquid density can be expected in this region. This higher aerated-liquid density will provide more resistance to liquid feed from the inlet nozzle in this region and, in turn, reduce the inlet feed to the kettle boiling region. This is the one region of the reboiler that actually operates as it should, like a kettle. Not much entrainment takes place here, and there is much disengagement space in this region, both of which promote the true kettle action. As the reboiler-feed liquid flow through this region is lower than the design and the percent vaporization is close to or higher than design, the temperature of vapor flowing to the last nozzles is close to or above design. If liquid entrainment from the intense boiling region is significant, it will generate a higher pressure loss in the reboiler outlet lines above the intense-boiling region. Since the pressure difference between Points 5 and 3 in Figure 1 is constant no matter which path is taken, this higher pressure loss in the nozzles above the intense boiling zone will induce a higher vapor flow into the outlet nozzles closer to the outlet baffle.
Mysteries explained This overall assessment of the reboiler operation led us to the development of a model. Described in greater detail in Reference [8], the model quantifies the theory presented above, and it explains all of the observations. The model confirms a maldistribu-
tion situation: a single zone can handle substantially more than the design liquid flowrate and produce much more than the design amount of vapor, based on a much improved temperature driving force. Further, the model confirms that this high vapor production comes about because of relatively small gradients in head within the tube field, brought about by regions of low froth density due to high vapor production. The relatively inactive zones explain the lower-than-design overall heat transfer. The regional maldistribution pattern reported here may be universal among kettle reboilers. But in most cases, it will not result in significant performance issues. Instead, this maldistribution is more significant in reboilers for mixtures such as those used in extractive distillation, where the boiling liquid consists of a large concentration of very high boilers and a small concentration of volatile components. The mulitifeed, multidraw arrangement plays a major role. A single feed, single draw kettle would have experienced a different maldistribution pattern, possibly even more severe. A stagnant region or dead zone is formed in the tubesheet quadrant by a self-accelerating mechanism. With little disengagement space, not much boiling takes place. The low-boiling components tend to preferentially vaporize out of the liquid, so the liquid becomes rich in the non-volatile component. The nonvolatile liquid, with little aeration, develops a high hydrostatic head, which resists movement of fresh liquid into this region and promotes stagnation. This region could constitute a significant performance issue, with lower than expected heat transfer rates, or lead to fouling in heat sensitive materials The high hydrostatic head in the stagnant zone, and the enhanced amount of fresh feed diverted into the central region, initiate a region of high L/V ratio, with vigorous boiling in the bulk of the bundle. The boiling is probably accompanied by entrainment into the overhead vapor line, low fractional vaporization, and a relatively high concentration of lights. Finally, the region near the overflow baffle receives close to the design feed flow and has ample disengagement
space, and is the only region behaving like a kettle. This pattern explains all the observations made: the high pressure drop, the measured temperatures, the lower heat transfer coefficient (caused by the stagnant region), and even the likely entrainment. The regional maldistribution pattern reported here gives complete support to the assumption that the reboiler liquid head in Equation (1), hs,REB, is the static head of non-aerated liquid in the reboiler. This head is set by the static head of the liquid in the stagnant region. This explanation closes the loop on the reason for the chimney tray overflow. The high reboiler static head, plus the high reboiler outlet line pressure drop led to the high liquid heads in the chimney tray. Entrainment from the intense boiling central zone could also have been a contributor. The low temperatures at the vapor outlet nozzles right above the intenseboiling region were caused by the high L/V ratios in this region, possibly assisted by entrainment. Above the kettle region, the L/V ratios were much lower, and so was entrainment, leading to the higher than expected temperatures. The high temperatures measured near the walls of the exchanger were caused by horizontal flow of hot liquid from the stagnant zone along the wall towards the boiling zones. The wall-supported foaming observed on the Oldershaw tests, and the slight foaming observed in the foaming tests, could also be a contributing factor that could aggravate the intense boiling effect. The malfunction reported here, namely, excess pressure drop in a kettle reboiler circuit when compared to the available head, was identified [7] as the most common malfunction experienced in reboilers. The employment of good design practices, coupled with compiling a valid reboiler pressure balance, can readily circumvent such malfunctions. In contrast, the maldistribution pattern reported here has not previously been reported. It is unique to mixtures such as those encountered in extractive distillation, where the boiling liquid consists of a large concen-
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Authors
Operations & Maintenance tration of high boilers and a small concentration of volatile components. This maldistribution affects both kettle heat transfer and pressure drop, and it, too, needs to be accounted for in the kettle design.
Lessons for design Kettle reboiler designs should: 1. Strictly adhere to the reboiler pressure-balance equation. Four-fifths of kettle reboiler malfunctions are caused by excess pressure drops in reboiler circuits [7]. 2. Be based upon the clear liquid head in the reboiler, not on aerated-liquid head. 3. Take account of froth density gradients within the reboiler, and the resulting maldistribution. 4. Trade off pressure loss through inlet nozzles (which provide liquid distribution to the reboiler) against the head and column height required. 5. Consider a horizontal thermosiphon
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reboiler as an alternative for mixtures containing a large concentration of components boiling at high temperatures together with components that are volatile. ■ Edited by Nicholas P. Chopey
References 1. Kister, H. Z., “Distillation Operation,” McGraw-Hill, New York, 1990. 2. Perry, J. H., “Chemical Engineers' Handbook,” 7th Ed., McGraw-Hill, New York, 1998. 3. Lieberman, N. P., “Process Design for Reliable Operation,” 2nd Ed., Gulf Publishing, Houston, 1988. 4. Palen, J. W. and Small, W. M., Kettle and Internal Reboilers, Hydrocarb. Proc., 43 No. 11, p. 199, 1964. 5. Fair, J. R. and Klip, A., Thermal Design of Horizontal Reboilers,” Chem. Eng. Prog., p. 86, March 1983. 6. Ishihara, K., others, Critical Review of Correlations for Predicting Two-Phase Flow Pressure Drop Across Tube Banks, Heat Transfer Engineering, 1, No. 3, p. 23, January–March 1980. 7. Kister, H. Z., What Caused Tower Malfunctions in the Last 50 Years?, Trans. IChemE, Vol. 81, Part A, p. 5, January 2003. 8. Rubbers, E., Green, K., Fowler, T., Kister, H. Z., and Stupin, W. J. Once-Thru Reboiler Startup Can be Exciting, in “Distillation 2003: on the Path to High Capacity, Efficient Splits,” Proceedings of Topical Conference, AIChE Spring Meeting, New Orleans, La., March 31–April 3, 2003.
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Etienne Rubbers is a corporate finance consultant for Sasol (Rosebank, Johannesburg, South Africa; Phone: +27 11 441 3472; e-mail:
[email protected]), involved in merger and acquisition activities. Previously, he was a process engineer in the concept development group, developing a number of the firm’s projects from concept to basic engineering, construction and commissioning. Despite his career change, he still has keen interest in process engineering. He holds a B.Sc. (chemical engineering) from Wits University (Johannesburg) and a master’s degree in industrial management from Belgium’s University of Leuven, and is a CFA charterholder. Kirsten Green is a senior process engineer for Sasol Technology (Phone: +27 11 344 0082; e-mail: kirsten.
[email protected]). She started her career at Sasol’s coal preparation plant, gaining experience in materials handling. She then joined the concept development group, where she has been involved in feasibility studies for the company’s chemical projects. She spent the past year on rotational training with MWKL in London, doing basic engineering in petroleum refining. She holds a B.Ing. from the University of Stellenbosch and an M.Sc. from University of Cape Town, both in chemical engineering. Terry-Ann Fowler is a senior process engineer for Sasol (Phone: +27 11 344 0092; email:
[email protected]). She forms part of a team involved in conceptual development and design of a number of the firm’s chemical and refinery projects. She has gained experience on projects from pilot plant studies, through feasibility and conceptual development, basic engineering, and on to construction and commissioning. Her B.Sc. and M.Sc., both being in chemical engineering, are from the University of Cape Town, and her Ph.D., also in chemical engineering, is from Wits University. Henry Z. Kister, a Fluor Corp. senior fellow and director of fractionation technology (Phone: 1-949-349-4679; email:
[email protected]), has over 25 years experience in design, troubleshooting, revamping, field consulting, control and startup of fractionation processes and equipment. Previously, he was Brown & Root’s staff consultant on fractionation, and worked for ICI Australia and Fractionation Research Inc. (FRI). The author of textbooks “Distillation Design” and “Distillation Operation,” as well as 70 published technical articles, he has taught the IChemE-sponsored “Practical Distillation Technology” course 250 times. A recipient of Chemical Engineering’s 2002 Award for Personal Achievement in Chemical Engineering, Kister holds B.E. and M.E. degrees from the University of NSW in Australia. He is a Fellow of IChemE and a member of AIChE, and serves on the FRI Technical Advisory and Design Practices Committees. Walter J. Stupin is an executive director for Fluor (Phone: 1-949-349-5209; email: walter.stupin @fluor.com). He manages and executes process engineering for petroleum refining and chemical plants. His over 40 years of process engineering experience ranges from vice president of technology at C F Braun Inc. to research at Fractionation Research Inc. (FRI). The distillation field has been an area of key interest, in which he has published over 20 technical papers. Dr. Stupin holds B.S., M.S. and Ph.D. degrees, all in chemical engineering, from the University of Southern California.
