- Email: [email protected]
Removal of Dissolved Aromatics from Water
0263±8762/98/$10.00+0.00 Institution of Chemical Engineers Trans IChemE, Vol. 76, Part A, July 1998
REMOVAL OF DISSOLVED AROMATICS FROM WATER Comparison of a High Intensity Contactor with a Packed Column B. WALDIE (FELLOW) and W. K. HARRIS Offshore Processing Research Group, Department of Mechanical & Chemical Engineering, Heriot-Watt University, Edinburgh, UK
A
n experimental comparison of the performances of a new high intensity contactor and a conventional packed column in stripping dissolved aromatics from water is described. The contactor gives signi® cantly lower values of HO L . These, combined with the much higher liquid ¯ uxes, give volumetric mass transfer coef® cients, KL a, several hundred times higher than those for the packed column and substantially higher than those in the literature for bubble columns. Performance is particularly good on saline water due to suppression of bubble coalescence. That characteristic would be bene® cial in the removal of aromatics from produced water in oil and gas production. The contactor utilizes enhanced acceleration ® elds achieved by water injection and so should be insensitive to the marine motion and tilt experienced on offshore ¯ oating production systems, as well as being compact. At constant water rate, KL a 3 G n with n ranging from 0.50 to 0.92. Some preliminary results on the effects of liquid feed rate and initial bubble size are included, the latter con® rming the importance of bubble size on performance. Pressure drop and consequent energy requirements are also considered. Keywords: process intensi® cation; bubble size and coalescence; aromatics; marine environment; offshore oil and gas; mass transfer; produced water
INTRODUCTION
The present study commenced in 1993 with funding from a group of operating companies and EPSRC as part of a Marine Technology Directorate programme on Treatment of Water Offshore. The main objective was to determine whether dissolved aromatics could be removed in a novel gas/liquid contactor and, if so, what reduction in equipment size might be achieved relative to a packed column on the same aromatics removal duty. Space and weight saving is obviously important offshore, especially in retro® tting an existing platform. The expected insensitivity to marine motion and tilt of the new contactor relative to a packed column would be a further advantage on ¯ oating production systems which are increasingly used for new ® eld developments. The packed column was of conventional design with the liquid dispersed, as for example in the stripping of contaminated water on-shore1 8 and deoxygenation of injection water offshore. As discussed later, other equipment types which could be considered for the present duty include packed bubble columns where the gas rather than the liquid is the dispersed phase. The overall process concept, based either on the new contactor or a packed column, is to use air as stripping gas and feed the ef¯ uent air to the main gas turbines or other combustion engine on the platform to dispose of the aromatics by combustion. If this were not possible then stripping with natural gas or condensation of aromatics from the ef¯ uent air might be considered.
The research described here was aimed at assessing the potential of a new high intensity contactor1 ,2 for providing a more compact means of removing dissolved aromatic compounds from produced water. Produced water can be the largest tonnage output stream in oil and gas production. This is certainly the case as a ® eld ages, the ratio of water/ oil ¯ ow rates reaching perhaps ten or more in the later years of ® eld life. In offshore production the produced water is most often discharged to sea and its composition, therefore, must be environmentally acceptable. Reinjection into the reservoir is being increasingly used as a more environmentally, and sometimes technically preferred alternative, but discharge to sea is likely to remain signi® cant for the foreseeable future. Most produced water contains dissolved aromatics such as benzene and toluene which, being soluble in the sea, are potentially more harmful to marine species than the insoluble dispersed oil. Current produced water treatment equipment on platforms, e.g. hydrocyclones, is designed only to remove dispersed or particulate oil. For the North Sea and several other offshore areas, current legislation and analytical monitoring techniques for oil content of produced water also deal only with that dispersed oil. With the possibility of more rigorous legislation on the composition of water discharged to sea, practical means of removing other components such as the dissolved aromatics need to be considered. In a review by Davies and Hansen3 of potential processes for removing various dissolved components from produced water, air stripping in a packed column was noted as likely to be the most economic method for removing aromatics.
AROMATICS AND PRODUCED WATER Aromatics of interest here are the lower molecular weight ones, principally benzene, toluene, ethyl benzene and 562
REMOVAL OF DISSOLVED AROMATICS FROM WATER
563
Table 1. Solubility12 of lower molecular weight aromatics in water at atmospheric pressure and 25°C, mg l ±1.
Benzene Toluene Ethyl Benzene o-Xylene m-Xylene
Distilled Water
Sea Water (3.45% wt)
1700 535 161 171 146
379 111 130 106
xylene, the so called BTEX group. Other aromatics present in produced water include naphthalenes. Solubilities of some lower molecular weight aromatics in water are shown in Table 1. It should be noted that the solubilities depend not only on temperature but also on the salinity of the water which can vary considerably from ® eld to ® eld and with time over the life of a ® eld. However, the levels are all well above 40 mg l ± 1 , the maximum level for dispersed oil permitted in UK and Norwegian regulations. It is partly due to this solubility that the aromatics are potentially more harmful to marine species being more readily incorporated into sea water and accessible to marine life. Industry based surveys of concentrations of dissolved aromatics in produced waters from oil and gas production platforms in the Norwegian North Sea3 and the Gulf of Mexico4 yielded data that is summarized in Table 2. Concentrations are well below saturation levels. Generally, the produced water from gas production contains higher concentrations. Also, values much higher than those in Table 2, up to several hundred mg l ± 1 , were reported5 for some gas platforms in the Dutch sector of the North Sea. The main source there was identi® ed as water condensate from the glycol regeneration system, aromatics being dissolved in the glycol during gas drying. Re-routing of that aromatics rich water stream or separate treatment was expected5 to eliminate that source. This is an example of how the detailed processing scheme on a platform can in¯ uence the aromatics content of the produced water. Other variables which contribute to the width of the ranges of concentrations in Table 2 include the type of oil, e.g. paraf® nic, asphaltenic or gas condensate and the salinity of the produced water. Whilst concentrations of aromatics tend to be higher from gas production platforms, the ¯ ow rates of produced water there are much lower, perhaps less than 1% of those encountered on some oil production platforms. On oil
Figure 1. High intensity contactor.
production platforms produced water ¯ ows of 1,000 m3 hr± 1 (< 150,000 barrels/day) and even higher may need to be treated. A packed column for that ¯ ow would be quite large, about 4 m or so in diameter. Finding space for such a column on an existing platform could be problematic. On a ¯ oating production ship there are additional problems in providing mechanical support and compensating for the effects of marine tilt and motion. HIGH INTENSITY CONTACTOR The high intensity gas/liquid contactor as described by Waldie1 ,2 utilizes high liquid shearing forces to produce small gas bubbles and hence high interfacial area coupled with centrifugal forces to enhance the mass transfer coef® cient and promote subsequent phase separation. These bene® ts are achieved by injecting liquid tangentially into a tube with a permeable wall through which the gas ¯ ows (Figure 1). The liquid forms a spinning layer on the inside of the tube wall. Gas passes through this layer as
Table 2. Concentrations of aromatics in produced waters from oil and gas ® elds, mg l ±1. Oil Field Area Norwegian North Sea Ref 3 Gulf of Mexico Ref 5
Components
{BTX { {Naphthalenes {Benzene { {Toluene { {C2 Benzenes
Trans IChemE, Vol 76, Part A, July 1998
Typical or mean
Gas Field Range
8
0±20
1.5
0±4
Typical or mean 25
Range 02
> 50
1.5
0±5
1.3
0.002±8.7
5.8
0.68±12.2
1.1
0.06±4.9
5.2
1.0±19.8
0.22
0.006±6.0
0.7
0.05±3.7
564
WALDIE and HARRIS
® ne bubbles and, due to the centrifugal forces, forms a central core. Gas leaves from the top and liquid from the base. With appropriate internal design, these two exit streams are virtually free of the other phase, e.g. the outgoing liquid is visually free of dispersed gas. Increase in mass transfer coef® cient resulting from an enhanced acceleration ® eld is thus achieved without the complication of mechanical rotation of the contactor, the latter approach being used in Higee type devices6 . Gas sparged cyclones have been described previously by Miller7 and others but there the objective was removal of particulates rather than dissolved materials, i.e. no diffusion based mass transfer. Enhanced mass transfer was the objective in a liquid cyclone with gas injection used by Beenackers and von Swaaij8 as a reactor. The geometry of that cyclone was different to that used here and phase separation was incomplete, the outgoing gas entraining about 20% of the liquid feed. The potential effect of the enhanced acceleration ® eld caused by tangential liquid injection can be seen from correlations for liquid side mass transfer coef® cients for small bubbles in agitated vessels and bubble columns. Taking for example the correlation by Calderbank9 kL =
2DL 1 db
0.31 Sc2
2/3
( rL 2
rG )mL g r 2L
1/3
(1)
kL increases approximately as (acceleration)1 / 3 if, as shown previously1 0 , the ® rst term on the right-hand side is relatively small. Here the liquid can be injected at velocities suf® cient to give radial accelerations, v 2 r, of the order 100±1000 g, at least in the inlet region of the device. There is thus the potential for the liquid side coef® cient, kL , to be increased up to some 3±10 times. As discussed later, the gas
stripping of dissolved aromatics is most likely liquid ® lm controlled and hence a signi® cant ampli® cation of the overall coef® cient KL would be expected. Equation (1) applies to bubbles smaller than about 1 mm which behave as rigid spheres. Preliminary measurements here on a shorter version of the contactor1 0 indicate bubble sizes well below that, about 300 mm and lower for the Sauter mean diameter. In equation (1) the mass transfer coef® cient is approximately independent of bubble size due to the relatively small magnitude of the ® rst right hand term. Reduction in bubble size does give a signi® cant additional contribution through the associated increased interfacial area per unit volume of gas. In previous deoxygenation studies2 , the reduction in bubble size in salt water as compared to fresh water resulting from suppression of bubble coalescence in salt water gave substantial improvements in mass transfer rate.
EXPERIMENTAL An outline ¯ ow scheme for the experimental facility built for this study is shown in Figure 2. The high intensity contactor or the packed column were connected in turn to the feed, sample and analytical system for comparison. Feed solutions of aromatics in water were made up batchwise in a 10,000 l stainless steel tank. Complete solution was ensured by pumped recirculation for about 2 hrs. For safety reasons, toluene and ethyl benzene were used in preference to benzene. Aromatic concentrations in the feed ranged from 29 to 194 mg l ± 1 in different runs. These are higher than the levels expected on oil platforms but are possible on some gas platforms. Using these higher concentrations makes for more reliable analysis.
Figure 2. Outline ¯ owscheme of experimental facility.
Trans IChemE, Vol 76, Part A, July 1998
REMOVAL OF DISSOLVED AROMATICS FROM WATER
Water was either fresh from the mains or a solution of 3.7% wt commercial granular salt (NaCl) in mains water. Maximum water feed rate was 1.5 kg s ± 1 (5.4 m3 hr ± 1 ), the limit of the centrifugal feed pump. Stripping air was metered from the laboratory compressed air system. Gas/ liquid (G/L) volumetric ratios up to 35 were used, depending on the contacting device. The packed column, an existing one incorporated into the new facility, was 220 mm in diameter with 2.15 m depth of 16 mm polypropylene Pall rings. This was operated countercurrently with liquid distributed evenly through a multipoint drip type distributor. The high intensity contactor (Figure 1) utilized a permeable ceramic tube, 32 mm internal diameter by 250 mm long for gas injection. Water was fed in tangentially through two feed points at one end. Gas essentially free of water drops leaves from a central top outlet and water free of gas from the base. Sample streams of feed and product water and of discharged air were passed through on-line analysers. Aromatics in water were measured by ¯ uorometry (Turner Designs Field Fluorometer Model No 10-AU-005). A free falling sample stream in a windowless cell is subjected to a narrow band excitation light beam and the intensity of the resulting emission from the aromatic molecule is measured. By using appropriate optical ® lters, the excitation and emission signals can be optimized for the compounds of interest. Here an excitation wave length of 254 nm and emission wave length of 289 nm were used. The aromatics, either toluene or ethyl benzene, were present singly in the feed and product solutions. Considerable effort was given to developing calibration procedures and overcoming problems in the original instrument. Critical factors in ensuring reliable calibration and on-line data include precise control of sample temperature and internal cell temperature. Replacement of the original acrylic light feedback pipe with a quartz one eliminated drift caused by uv degradation of the acrylic. Aromatics concentrations in the outgoing air were measured with a photo ionization instrument (ELE Instruments Portable Solvent Detector No 300-100). Cross checks on the output of this instrument were made on a chromatograph (Perkin Elmer 8500 series with ¯ ame ionization detector) using batchwise sample bags. The photo ionization instrument proved to be very sensitive to sample conditions. Under carefully controlled conditions the toluene balance from simultaneous data from the water and air analysers was consistent within 6 10%. For routine runs, the ¯ uorometer data for concentrations of aromatic in the water feed and product, the streams of most interest, were used to calculate mass transfer performance. Water, feed and product temperatures which could differ by up to 2°C were measured and averaged. These temperatures were steady over a run but the absolute level depended on weather conditions as the feed tank was located outside. RESULTS AND DISCUSSION Analysis Measured concentrations of aromatic in feed and product water, together with water and air ¯ ow rates, were used to calculate the number of overall liquid transfer units, NO L , achieved in a given run. As the incoming air was free of Trans IChemE, Vol 76, Part A, July 1998
565
aromatics, NO L is given by1 1 : x2 12 x1
1 1 1 S S NOL = 1 12 S where the stripping factor S is given by H Gm S= P Lm
loge
(2)
(3)
The removal of toluene and ethyl benzene has been assumed to be liquid ® lm diffusion controlled as the calculated Henry’ s Law constants, H, are suf® ciently high, 36.2 MN m ± 2 mole fraction and 46.5 MN m ± 2 mole fraction respectively in distilled water at 25°C. Thus in equation (4) for the overall volumetric mass transfer coef® cient based on the liquid side, 1 1 1 = (4) 1 KL a kL a H kG a the last term is relatively small and is ignored. This simpli® cation is adopted in the design of columns1 8 for stripping of these materials from water. It also appears valid for conventional bubble column according to the somewhat limited literature data on kG and kG a for such columns, e.g. that reviewed by Deckwer1 9 including the experimental data of Mashelkar and Sharma2 0 . The validity for the present contactor where enhanced values of kL a are expected is discussed later. Calculated values of HO L depend on the values taken for H which varies with temperature and salinity. The two values quoted before were derived from published data on solubility1 2 and vapour pressure1 6 of the two aromatics. Solubilities measured by Sutton and Calder1 2 show a signi® cant reduction in solubility in sea water as compared to distilled water. The resulting increases in calculated H values are over 40% and this effect has been taken into account here. The temperature dependencies of solubility and hence of H were based on experimental solubilities at 25° and 0°C reported by Polak and Lu1 3 . These data were for distilled water. In the absence of measured data, the same dependency was assumed for salt water. Calculated values of HO L are most sensitive to uncertainties in H at lower values of G/L. One basis for comparing the mass transfer performances of the high intensity contactor and the packed column is the equivalent height of an overall liquid transfer unit, HO L , where HOL =
Height of mass transfer zone NOL
(5)
This is a somewhat limited and approximate comparator compared to the volumetric mass transfer coef® cient considered later but is applied in view of the widespread use of the height of a transfer unit as a measure of the mass transfer performance of conventional packed columns. As in the previous study of deoxygenation2 , the height of the mass transfer zone for the high intensity contactor is taken as the tube length, 0.25 m. For the column it is obviously the packed height. Comparison of HO L values is of practical interest as regards height requirements, an important factor in installing equipment on offshore platforms
566
WALDIE and HARRIS
especially in retro-® tting. The present data on HO L for the contactor is however limited as regards extrapolation to separation processes requiring more transfer units than the relatively simple stripping of volatile aromatics. Further work is needed to establish how performance varies with tube length, as discussed later. Strictly, equation (5) applies to a pure countercurrent ¯ ow situation whereas ¯ ow in the new contactor is more complex involving cross and counter¯ ow. However, even in a packed column there are conditions at high L/G where some recirculation of gas is likely and yet the height of a transfer unit concept is still applied for simplicity. That simple concept is used for convenience here together with the volumetric mass transfer coef® cient, KL a, where Lm KL a = (6) rm HOL KL a is a more comprehensive comparator than HO L as it takes into account both the height requirement and the liquid ¯ ux capabilities of the two devices. The in¯ uence of the ¯ uid ¯ ow model assumed for the contactor on calculated KL a values is considered later. In these mass transfer calculations, no deductions were made for potential contributions from end effects in the packed column. The packing performance data may therefore be somewhat overstated. End effects seem less likely in the high intensity contactor. Comparison of HO L Values An overall comparison of HO L results for the high intensity contactor and the packed column is presented in Figure 3. The contactor gives lower values of HO L at all gas/liquid ratios, the improvement being particularly marked when treating salt water. For example on salt water at G/L of 3.0, the ® ne bubble version of the contactor gives HO L values some 7±10 times better than the packed column.
Figure 3. Comparison of HOL values for high intensity contactor and packed column in removal of toluene from water.
The relative effect of water salinity on the two devices is interesting. Changing from fresh to salt water on the packed column introduces a foaming problem which promotes ¯ ooding. This prevented operation of the column with 75 l/min of salt water at G/L ratios greater than 2.0. HO L was consequently very high at 2 m. The new contactor, however, performs much better on salt water than fresh water. This con® rms the trends reported previously2 for deoxygenation of water in the contactor and in a packed column. The improvement with salt water in the contactor was attributed to suppression of bubble coalescence by dissolved salt leading to smaller mean bubble sizes and greater interfacial area. More recent1 0 bubble size distribution data for the outlet of a shorter version of the contactor helped quantify that effect. The presence of an electrolyte such as salt creates concentration and surface tension gradients which promote the stability of bubbles in liquids1 4 . As produced water is inevitably saline, often to levels much higher than the near sea water salinity used here, then this effect is particularly bene® cial. In contrast, the same concentration and surface tension mechanisms promote foam persistence in the packed column and hence reduce the maximum liquid ¯ ux. Measured Sauter mean diameters at the exit of the shorter contactor1 0 were 285 mm and 150 mm respectively for fresh and salt water. Whilst these are unlikely to represent the absolute values of mean bubble size in the present version, they con® rm the strong in¯ uence of salinity. A reduction from 285 to 150 mm approximately doubled1 0 the value of KL a. The signi® cance of bubble size in the high intensity contactor has been further demonstrated here by changing the initial bubble size range. This was done by using a permeable tube with generally coarser pores, the nominal mean pore size being 90 mm as compared to 10 mm. Quantitative data on the relative bubble sizes is not yet available but the effect on HO L in fresh water is quite marked (Figure 3). It will be noted in Figure 3 that for the ® ne bubbles in a given water two plots of HO L data are shown. These represent respectively the best, i.e. lowest, and worst HO L values obtained in different series of runs. In a given series of runs the gas rate was progressively changed at a constant liquid rate, usually 1.5 kg s2 1 . Resultant HO L values for such a series then decrease systematically with increasing G as shown (Figure 3). More detailed data for the coarse bubbles showing maximum and minimum trend lines and a mean line are given in Figure 4. The extent of the offset between different series of runs is too great to be attributed to temperature differences which could affect HO L through the temperature dependence of diffusivity. Following extensive work on the ¯ uorometry analysis technique, step changes in say analyser sensitivity from series to series appear unlikely. The possibility of a change in the mean bubble size from series to series is to be investigated once techniques for measuring bubble sizes within the actual contactor have been developed. For the present, even the worst HO L data are signi® cantly better than those for the packed column. Fewer runs were done with ethyl benzene the purpose being mainly to establish that performance of the contactor was not speci® c to toluene. HO L values decrease systematically with increase in G at constant L but the absolute Trans IChemE, Vol 76, Part A, July 1998
REMOVAL OF DISSOLVED AROMATICS FROM WATER
Figure 4. HOL for high intensity contactor with coarse bubbles in separation of toluene from fresh water.
values are generally lower than for toluene at the same liquid rate (Figure 5). As the diffusivity of ethyl benzene in water is lower than that of toluene, HO L values were expected to be somewhat greater for ethyl benzene. The few runs done at liquid rates lower than the standard 1.5 kg s2 1 gave higher values of HO L . This is in line with theoretical expectations from equation (1) of increased mass transfer rate in an enhanced acceleration ® eld and an expected decrease in bubble size. More data are needed to establish a quantitative relationship. The systematic decrease in contactor HO L with increasing G is similar to the trend found previously2 for deoxygenation. There HO L varied as G 2 n with values of 0.56 and 0.67 for the exponent n for salt and fresh water respectively. Here there is more variation in n from run series to run series for a given water, the values ranging overall from 0.50 to 0.92. As for deoxygenation, the dependence of HO L on G
567
is more pronounced with fresh water. This systematic decrease in HO L with G for the contactor contrasts with the behaviour of a packed column where HO L decreases initially but eventually reaches a near plateau value. The 0.83 kg s± 1 data for fresh water (Figure 3) in the packed column is an example. Most published data on height of a transfer unit for packings apply to this plateau region, G/L values for many packed column applications typically exceeding 10. The present comparison of the contactor with a packed column applies strictly to one particular type and size of packing and the generality of this comparison may be queried. First it should be noted that the present measured HO L values for the packed column are about 40% better than those predicted by applying Onda’ s correlations1 5 . Those correlated experimental data to within 6 20%, although Pall rings were not available for those experiments. One factor which will have led to some overstatement of the present packing performance is the omission of a correction for column end effects. Experimental determination of that correction would have required a series of runs with different depths of packing. Other evidence that the present packed column HO L data do not understate the performance of a typical random packed column comes from later work on the same toluene/water system using a 1 metre diameter column with a nominally more advanced metal packing. Comparison of Volumetric Mass Transfer Coef® cients, KL a A comparison of the performance of the high intensity contactor and packed column in terms of mass transfer rate per unit of contacting volume is given by the ratio of the KL a values (Table 3). KL a values were derived from HO L data and equation (6). These KL data therefore are based on the countercurrent plug ¯ ow approximation for water and gas ¯ ow within the contactor and the packed column. For the contactor where, as discussed previously, there were apparent changes in performance from run series to run series the worst and a mean of best and worst HO L Table 3. Comparison of mass transfer performance of high intensity contactor and packed column on volumetric basis. KL a Contactor KL a Packed Column Fresh Water Column feed rate* l/min CONTACTOR Fine bubbles
Trans IChemE, Vol 76, Part A, July 1998
75
50
75
G/L
Mean Worst
2.0 2.0
Mean Worst Mean Worst
3.3 3.3 5.0 5.0
231 154 240 170
175 118 208 148
3.3 5.0
98 110
75 95
Coarse Bubbles Mean Mean
Figure 5. HOL for high intensity contactor in separation of ethylbenzene from fresh water.
50
Salt Water
344 285
* Contactor feed rate 90 l/min (1.5 kg s ±1) throughout.
668 589 711 654
568
WALDIE and HARRIS
values for a given G/L value have been used. Thus even with the best results excluded, the high intensity contactor can achieve mass transfer rates per unit volume some several hundred times greater than a packed column. That is due to the lower equivalent HO L but more signi® cantly to the much higher liquid ¯ uxes per unit cross sectional area. The initial results at different liquid rates (Figure 5) suggest that at rates above 1.5 kg s ± 1 still higher KL a values would be achievable, the increased ¯ ux effect being supplemented by reduction in HO L . The ¯ ow patterns of gas and liquid within the contactor are clearly more complex than for countercurrent plug ¯ ow. There will be some initial cross ¯ ow of gas through the liquid, some co-current ¯ ow as bubbles begin to travel with the liquid and some countercurrent ¯ ow as gas and liquid leave from opposite ends of the tube. Effective mean mass transfer driving force will therefore be less than that for countercurrent plug ¯ ow. Thus true KL a values for the contactor will be higher than those reported here. Substantial further work is needed to quantify the actual ¯ ow patterns and develop a more exact model for the mass transfer processes. The potential extent of the dependence of KL a on the ¯ ow model assumed can be seen by treating the contactor as a perfect mixer. This model has been used in the analysis of some more conventional bubble columns1 9 ,2 1 ,2 2 . Taking for example a contactor run in salt water with ® ne bubbles and water and air rates of 90 and 450 l/min respectively, KL a calculated for the plug ¯ ow model is 10.1 s ± 1 . For the same feed and product compositions, KL a for perfect mixing is either 13.9 or 28.7 s ± 1 depending on the mean mass transfer driving force DC in equation (7) for a perfectly mixed contactor. KL a =
1 Ci 2 Co Åt DC
(7)
In modelling studies of oxygenation in a perforated plate bubble column2 2 and a packed bubble column2 1 , the compositions of the gas, air or oxygen, were essentially unchanged through the columns. DC was therefore taken as C * 2 Co , C * being the liquid concentration in equilibrium with air or oxygen respectively. Here the toluene partial pressure in the air increases through the contactor from zero to the outlet value. Taking C * as that for equilibrium with the inlet gas, i.e. zero, gives the lower value of KL a for the perfectly mixed case, 13.9 s ± 1 . Taking C * as the mean of the equilibrium values for inlet and outlet gas gives the higher KL a value, 28.7 s ± 1 . The latter mean concentration appears more reasonable but yields a KL a value which is 2.84 times higher than the value from the countercurrent plug ¯ ow model. Thus the countercurrent plug ¯ ow values of KL a for the contactor used in Table 2 are, if anything, conservative. From equation (5) and the experimental variation of HO L with G (Figures 3, 4 and 5). KL a 3 G n
(8)
n ranges from 0.50 to 0.92. The range of n values was narrower in the previous study of deoxygenation in the contactor2 . For classical bubble columns where gas is bubbled through a relatively slow axially ¯ owing liquid, the equivalent exponent is typically1 7 about 0.7 with some values as high as 0.82.
The high intensity contactor has been compared experimentally with a conventional packed column in which the liquid is the dispersed phase. Other potential competing equipment includes various forms of bubble column where, as in the high intensity contactor, the gas is the dispersed phase. Comparison of present KL a values with those for bubble columns is therefore of interest. The highest values of KL a for bubble columns found in the literature are those reported by Wang and Fan2 1 for a packed bubble column and by Voigt and SchuÈgerl 2 2 for a multi stage perforated plate bubble column. Both groups studied oxygenation. Wang and Fan2 1 con® rmed that KL a values for an unpacked column could be increased signi® cantly by installing motionless mixer elements of corrugated metal sheet. KL a values for the packed column reached 0.23 s ± 1 or 0.16 s± 1 depending on whether the perfectly mixed or plug ¯ ow model was applied to the measured data. That column was operated co-currently in upward ¯ ow. The KL a values are thus lower by a factor of 50 or more than present values. Also, if co-current operation were used on the present duty, much higher G/L ratios would be needed for a given separation duty than in the high intensity contactor to maintain a positive driving force throughout the column. This latter problem would not arise with the perforated plate bubble column of Voigt and SchuÈgerl2 2 which was operated countercurrently. For salt water, KL a values based on the perfect mixer model were mainly below 0.3 s± 1 except for one value of 1.66 s ± 1 which was attributed to foam formation. Film Control Present values of HO L and KL a are based on the assumption of liquid ® lm control, reasonable for a conventional packed column or bubble column at the level of Henry’ s Law coef® cient involved. The anticipated improvement in kL in the new contactor could, if suf® ciently great, invalidate that assumption (equation (4)). To check this possibility, estimates of the gas side mass transfer coef® cient inside a bubble have been made for comparison with the overall coef® cient KL . Transfer by molecular diffusion only is assumed in view of the relatively small bubbles. Equations (8) and (9) are equivalent to those developed by Calderbank9 for the internal ® lm coef® cient for a ¯ uid drop where the concentration of diffusing species at the interface is constant. H kG = 2 H kG <
DG pt
1 p 2 DG 3 R
0.5
(9) (10)
Equation (9) applies to short contact times where the fractional approach to equilibrium is <0.4 and (10) to longer times where there is a closer approach to equilibrium. Applying equation (10) and taking R as 75 mm from the Sauter mean diameter and the speci® c area, a, from that diameter and an assumed fractional gas hold up of 0.05, gives a value of H kG a of about 700 s ± 1 . Equation (9) with t of 0.026 s, an upper limit for the mean gas residence time, gives about 40 s ± 1 . The residence time of the gas when present as individual bubbles will be less than this Trans IChemE, Vol 76, Part A, July 1998
REMOVAL OF DISSOLVED AROMATICS FROM WATER giving a higher value of H kG a. In practice, the interface concentration will change with time but these values of H kG a even if very approximate are suf® ciently greater than the measured overall KL a values to suggest that the process is still liquid ® lm controlled. Equipment Size and Scale-up The KL a ratios in Table 3 show that the internal contacting volume of a high intensity contactor could be over two orders of magnitude smaller than that of a packed column on the same duty. The ratio of overall equipment volumes would not be as large but still be signi® cant. Comparison with literature data on KL a for bubble columns, including packed bubble columns, suggests that the new contactor would also be signi® cantly more compact than those. Scale-up for larger water ¯ ow rates could involve increasing the tube diameter and/or using several tubes in parallel. The latter would be achieved most compactly by mounting the tubes in a common shell as is done for de-oiling hydrocyclones. The water ¯ ux in the present tube is higher than in some commercial de-oiling hydrocyclones and so an acceptably compact arrangement should be possible. Research is needed to establish how tube diameter and tube length affect performance. With the present tube length, 0.25 m, somewhat over 1 transfer unit has been achieved. Mass transfer per unit length will presumably decrease with tube length due to frictional losses and eventual decay of the swirling water ¯ ow pattern. The device as currently con® gured is therefore most suited to relatively simple separations such as stripping of aromatics or oxygen from water which require few transfer units. Energy Consumption As in deoxygenation2 , the substantial reduction in contacting volume inevitably incurs an increase in energy consumption over that for a packed column. The increase may however be less pronounced for produced water as that could be available under pressure from the primary separators. At water and air ¯ ow rates of 90 l/min and 300 l/min, i.e. G/L = 3.33, pressure drops were 5.5 ´ 105 N/m2 and 0.86 ´ 105 N/m2 respectively. Theoretical minimum energy requirements per m3 water are therefore 0.153 Kw-hr and 0.079 Kw-hr respectively giving a total of 0.232 kW-hr/m3 . Pumping and gas compression losses need to be added. These data apply to one particular con® guration which was not optimized. There are trade-offs between contactor volume and energy consumption involving for example the choice of G/L ratio and the permeability of the tube and hence bubble size. Energy consumption could be reduced if produced water were taken from a pressurized primary separator and further reduced if stripping were done by hydrocarbon gas from the primary separators. Comparing the high intensity contactor with the bubble columns giving the highest reported KL a values shows that the contactor gives signi® cantly higher values of KL a for a given energy input per unit volume of water treated. For example, the perforated plate bubble column of Voigt and Trans IChemE, Vol 76, Part A, July 1998
569
SchuÈgerl2 2 even when giving enhanced KL a values due apparently to foaming required some 2.8 times more energy per unit volume of feed water per unit KL a. Without foaming, the ratio would be over 10. CONCLUSIONS The high intensity contactor has the potential for signi® cant reductions in equipment volume and weight as compared to a packed column in the removal of dissolved aromatics from water by gas stripping. The expected insensitivity of the contactor to marine motion and tilt would be a further advantage on a ¯ oating production system. Results on stripping of aromatics con® rm several aspects of contactor performance characteristics found in a previous study of deoxygenation. The volumetric mass transfer coef® cient, KL a, can be several hundred times higher than that for a conventional packed column. The high intensity contactor also compares favourably in that respect with perforated plate and packed bubble columns in the literature. Performance is enhanced in saline water due to suppression of bubble coalescence. Experiments with two levels of initial bubble size con® rm the importance of bubble size on mass transfer performance. The volumetric mass transfer coef® cient, KL a, varies as G n with n here ranging from 0.50 to 0.92, a wider range than that for deoxygenation. Reasons for this variation and other fundamental aspects including bubble size distribution, gas hold up, capacity limits and scale-up require further study. Speci® c energy consumption is higher than for a packed column but could be reduced in produced water treatment if use can be made of already pressurized liquid feed and perhaps gas from the primary separators. NOMENCLATURE a db DG DL G Gm g H HOL kG kL KL L Lm n NOL P r R S t Åt Sc x1 , x 2
speci® c surface area, m ±1 bubble diameter, m gas phase diffusivity, m2 s ±1 liquid phase diffusivity, m2 s ±1 gas ¯ ow rate, m3 s ±1 gas molar ¯ ux, kmol m ±2 s ±1 gravitational acceleration, m2 s ±1 Henry’ s Law coef® cient, Nm ±2 mole fraction ±1 height of overall liquid transfer unit, m gas side mass transfer coef® cient, mole fraction m3 N ±1 s ±1 liquid side mass transfer coef® cient, m s ±1 overall liquid mass transfer coef® cient, m s ±1 liquid ¯ ow rate, m3 s ±1 liquid molar ¯ ux, kmol m ±2 s ±1 exponent number of overall liquid transfer units total absolute pressure, N m ±2 radial distance, m bubble radius, m stripping factor time, s mean liquid residence time, s Schmidt number mole fractions of aromatic in product and feed
Greek rL , rG rm v
letters liquid and gas densities, kg m ±3 molar density of liquid,kmol m ±3 rotational speed, s ±1
570
WALDIE and HARRIS REFERENCES
1. Waldie, B., 1995, Novel high intensity gas/liquid contactor, Proc 1st Int Conf on Science, Engineering and Technology of Intensive Processing (Akay, G. and Azzopardi, B. J. (eds)) University of Nottingham, 5. 2. Waldie, B., 1996, A new high intensity contactor for de-oxygenation of water, Trans IChemE, 74(A2): 183±189. 3. Hansen, B. R. and Davies, S. R. H., 1994, Review of potential technologies for the removal of dissolved components from produced water, Trans IChemE, 72(A2): 176±8. 4. Stephenson, M. T., 1992, Components of produced water: a compilation of industry studies, J Pet Tech, 44: 548±550 and 602±603. 5. Callaghan, D. and Baumgartner, W., 1990, Characterization of residual hydrocarbons in produced water discharged from gas production platforms, SPE Paper 2088. 6. Balasundaram, V., Porter, J. E. and Ramshaw, C., 1989, Process intensi® cation: a rotary seawater deaerator, Proc Symp on Separation of Gases (Royal Soc of Chemistry), 306. 7. Miller, J. D. and Van Camp, M. C., 1982, Mining Eng, 1: 57. 8. Beenackers, A. A. C. M. and van Swaaij, W. P. M., 1976, Mass transfer and chemical reaction in a gas-liquid cyclone reactor, Proc 4th Int 6th Europ Symp on Chemical Reaction Engineering, (Dechema Frankfurt), 260. 9. Calderbank, P. H., 1967, Mass transfer, Chapt 6, in Mixing Theory and Practice, Vol II, Uhl, V. W. and Gray, J. B. (Academic Press). 10. Waldie, B. and Harris, W. K., 1997, Stripping performance of a new high intensity contactor, Proc Distillation and Absorption 97’ IChemE Symp Series No 142, Vol 1: 371±378 Ed Darton, R. 11. Treybal, R. E., 1980, Mass Transfer Operations, 3rd ed (McGraw Hill, New York). 12. Sutton, C. and Calder, J. A., 1975, Solubility of alkylbenzenes in distilled water and seawater at 25°C, J Chem and Eng Data, 20: 320. 13. Polak, J. and Lu, B. C. Y., 1973, Mutual solubilities of hydrocarbons and water at 0 and 25°C, Can J Chem Eng, 51: 4013.
14. Marrrucci, G., 1969, A theory of coalescence, Chem Eng Sci, 24: 975. 15. Onda, K., Takeuchi, H. and Okumoto, Y., 1968, Mass transfer coef® cients between gas and liquid phases in packed columns, J Chem Eng Jpn, 1: 56. 16. Perry, R. H. and Chilton, C. H., 1979, Chemical Engineers Handbook, 5th Ed (McGraw Hill, New York). 17. Akita, K. and Yoshida, F., 1973, Ind Eng Chem Proc Dev, 12: 76. 18. Striggle, R. F., 1987, Random packings and packed towers, (Gulf Publishing Co, Houston). 19. Deckwer, W. D., 1992, Bubble Column Reactors, (J. Wiley, Chichester). 20. Mashelkar, R. A. and Sharma, M. M., 1970, Mass transfer in bubble and packed bubble columns, Trans IChemE, 48: T162. 21. Wang, K. B. and Fan, L. T., 1978, Mass transfer in packed columns packed with motionless mixers, Chem Eng Sci 33: 945. 22. Voigt, J. and SchuÈgerl, K., 1979, Absorption of oxygen in countercurrent multistage bubble columns ± I, Chem Eng Sci, 34: 1221.
ACKNOWLEDGEMENTS This project was supported by EPSRC, DTI and eight operating companies as part of the MTD managed programme on Treatment of Water Offshore, III. Craig Bell and Ron Millar are thanked for their contributions to the experimental runs and rig construction.
ADDRESS Correspondence concerning this paper should be addressed to Professor B. Waldie, Department of Mechanical and Chemical Engineering, HeriotWatt University, Edinburgh EH14 4AS, UK. The manuscript was received 27 March 1998 and accepted for publication after revision 15 May 1998.
Trans IChemE, Vol 76, Part A, July 1998
REMOVAL OF DISSOLVED AROMATICS FROM WATER Comparison of a High Intensity Contactor with a Packed Column B. WALDIE (FELLOW) and W. K. HARRIS Offshore Processing Research Group, Department of Mechanical & Chemical Engineering, Heriot-Watt University, Edinburgh, UK
A
n experimental comparison of the performances of a new high intensity contactor and a conventional packed column in stripping dissolved aromatics from water is described. The contactor gives signi® cantly lower values of HO L . These, combined with the much higher liquid ¯ uxes, give volumetric mass transfer coef® cients, KL a, several hundred times higher than those for the packed column and substantially higher than those in the literature for bubble columns. Performance is particularly good on saline water due to suppression of bubble coalescence. That characteristic would be bene® cial in the removal of aromatics from produced water in oil and gas production. The contactor utilizes enhanced acceleration ® elds achieved by water injection and so should be insensitive to the marine motion and tilt experienced on offshore ¯ oating production systems, as well as being compact. At constant water rate, KL a 3 G n with n ranging from 0.50 to 0.92. Some preliminary results on the effects of liquid feed rate and initial bubble size are included, the latter con® rming the importance of bubble size on performance. Pressure drop and consequent energy requirements are also considered. Keywords: process intensi® cation; bubble size and coalescence; aromatics; marine environment; offshore oil and gas; mass transfer; produced water
INTRODUCTION
The present study commenced in 1993 with funding from a group of operating companies and EPSRC as part of a Marine Technology Directorate programme on Treatment of Water Offshore. The main objective was to determine whether dissolved aromatics could be removed in a novel gas/liquid contactor and, if so, what reduction in equipment size might be achieved relative to a packed column on the same aromatics removal duty. Space and weight saving is obviously important offshore, especially in retro® tting an existing platform. The expected insensitivity to marine motion and tilt of the new contactor relative to a packed column would be a further advantage on ¯ oating production systems which are increasingly used for new ® eld developments. The packed column was of conventional design with the liquid dispersed, as for example in the stripping of contaminated water on-shore1 8 and deoxygenation of injection water offshore. As discussed later, other equipment types which could be considered for the present duty include packed bubble columns where the gas rather than the liquid is the dispersed phase. The overall process concept, based either on the new contactor or a packed column, is to use air as stripping gas and feed the ef¯ uent air to the main gas turbines or other combustion engine on the platform to dispose of the aromatics by combustion. If this were not possible then stripping with natural gas or condensation of aromatics from the ef¯ uent air might be considered.
The research described here was aimed at assessing the potential of a new high intensity contactor1 ,2 for providing a more compact means of removing dissolved aromatic compounds from produced water. Produced water can be the largest tonnage output stream in oil and gas production. This is certainly the case as a ® eld ages, the ratio of water/ oil ¯ ow rates reaching perhaps ten or more in the later years of ® eld life. In offshore production the produced water is most often discharged to sea and its composition, therefore, must be environmentally acceptable. Reinjection into the reservoir is being increasingly used as a more environmentally, and sometimes technically preferred alternative, but discharge to sea is likely to remain signi® cant for the foreseeable future. Most produced water contains dissolved aromatics such as benzene and toluene which, being soluble in the sea, are potentially more harmful to marine species than the insoluble dispersed oil. Current produced water treatment equipment on platforms, e.g. hydrocyclones, is designed only to remove dispersed or particulate oil. For the North Sea and several other offshore areas, current legislation and analytical monitoring techniques for oil content of produced water also deal only with that dispersed oil. With the possibility of more rigorous legislation on the composition of water discharged to sea, practical means of removing other components such as the dissolved aromatics need to be considered. In a review by Davies and Hansen3 of potential processes for removing various dissolved components from produced water, air stripping in a packed column was noted as likely to be the most economic method for removing aromatics.
AROMATICS AND PRODUCED WATER Aromatics of interest here are the lower molecular weight ones, principally benzene, toluene, ethyl benzene and 562
REMOVAL OF DISSOLVED AROMATICS FROM WATER
563
Table 1. Solubility12 of lower molecular weight aromatics in water at atmospheric pressure and 25°C, mg l ±1.
Benzene Toluene Ethyl Benzene o-Xylene m-Xylene
Distilled Water
Sea Water (3.45% wt)
1700 535 161 171 146
379 111 130 106
xylene, the so called BTEX group. Other aromatics present in produced water include naphthalenes. Solubilities of some lower molecular weight aromatics in water are shown in Table 1. It should be noted that the solubilities depend not only on temperature but also on the salinity of the water which can vary considerably from ® eld to ® eld and with time over the life of a ® eld. However, the levels are all well above 40 mg l ± 1 , the maximum level for dispersed oil permitted in UK and Norwegian regulations. It is partly due to this solubility that the aromatics are potentially more harmful to marine species being more readily incorporated into sea water and accessible to marine life. Industry based surveys of concentrations of dissolved aromatics in produced waters from oil and gas production platforms in the Norwegian North Sea3 and the Gulf of Mexico4 yielded data that is summarized in Table 2. Concentrations are well below saturation levels. Generally, the produced water from gas production contains higher concentrations. Also, values much higher than those in Table 2, up to several hundred mg l ± 1 , were reported5 for some gas platforms in the Dutch sector of the North Sea. The main source there was identi® ed as water condensate from the glycol regeneration system, aromatics being dissolved in the glycol during gas drying. Re-routing of that aromatics rich water stream or separate treatment was expected5 to eliminate that source. This is an example of how the detailed processing scheme on a platform can in¯ uence the aromatics content of the produced water. Other variables which contribute to the width of the ranges of concentrations in Table 2 include the type of oil, e.g. paraf® nic, asphaltenic or gas condensate and the salinity of the produced water. Whilst concentrations of aromatics tend to be higher from gas production platforms, the ¯ ow rates of produced water there are much lower, perhaps less than 1% of those encountered on some oil production platforms. On oil
Figure 1. High intensity contactor.
production platforms produced water ¯ ows of 1,000 m3 hr± 1 (< 150,000 barrels/day) and even higher may need to be treated. A packed column for that ¯ ow would be quite large, about 4 m or so in diameter. Finding space for such a column on an existing platform could be problematic. On a ¯ oating production ship there are additional problems in providing mechanical support and compensating for the effects of marine tilt and motion. HIGH INTENSITY CONTACTOR The high intensity gas/liquid contactor as described by Waldie1 ,2 utilizes high liquid shearing forces to produce small gas bubbles and hence high interfacial area coupled with centrifugal forces to enhance the mass transfer coef® cient and promote subsequent phase separation. These bene® ts are achieved by injecting liquid tangentially into a tube with a permeable wall through which the gas ¯ ows (Figure 1). The liquid forms a spinning layer on the inside of the tube wall. Gas passes through this layer as
Table 2. Concentrations of aromatics in produced waters from oil and gas ® elds, mg l ±1. Oil Field Area Norwegian North Sea Ref 3 Gulf of Mexico Ref 5
Components
{BTX { {Naphthalenes {Benzene { {Toluene { {C2 Benzenes
Trans IChemE, Vol 76, Part A, July 1998
Typical or mean
Gas Field Range
8
0±20
1.5
0±4
Typical or mean 25
Range 02
> 50
1.5
0±5
1.3
0.002±8.7
5.8
0.68±12.2
1.1
0.06±4.9
5.2
1.0±19.8
0.22
0.006±6.0
0.7
0.05±3.7
564
WALDIE and HARRIS
® ne bubbles and, due to the centrifugal forces, forms a central core. Gas leaves from the top and liquid from the base. With appropriate internal design, these two exit streams are virtually free of the other phase, e.g. the outgoing liquid is visually free of dispersed gas. Increase in mass transfer coef® cient resulting from an enhanced acceleration ® eld is thus achieved without the complication of mechanical rotation of the contactor, the latter approach being used in Higee type devices6 . Gas sparged cyclones have been described previously by Miller7 and others but there the objective was removal of particulates rather than dissolved materials, i.e. no diffusion based mass transfer. Enhanced mass transfer was the objective in a liquid cyclone with gas injection used by Beenackers and von Swaaij8 as a reactor. The geometry of that cyclone was different to that used here and phase separation was incomplete, the outgoing gas entraining about 20% of the liquid feed. The potential effect of the enhanced acceleration ® eld caused by tangential liquid injection can be seen from correlations for liquid side mass transfer coef® cients for small bubbles in agitated vessels and bubble columns. Taking for example the correlation by Calderbank9 kL =
2DL 1 db
0.31 Sc2
2/3
( rL 2
rG )mL g r 2L
1/3
(1)
kL increases approximately as (acceleration)1 / 3 if, as shown previously1 0 , the ® rst term on the right-hand side is relatively small. Here the liquid can be injected at velocities suf® cient to give radial accelerations, v 2 r, of the order 100±1000 g, at least in the inlet region of the device. There is thus the potential for the liquid side coef® cient, kL , to be increased up to some 3±10 times. As discussed later, the gas
stripping of dissolved aromatics is most likely liquid ® lm controlled and hence a signi® cant ampli® cation of the overall coef® cient KL would be expected. Equation (1) applies to bubbles smaller than about 1 mm which behave as rigid spheres. Preliminary measurements here on a shorter version of the contactor1 0 indicate bubble sizes well below that, about 300 mm and lower for the Sauter mean diameter. In equation (1) the mass transfer coef® cient is approximately independent of bubble size due to the relatively small magnitude of the ® rst right hand term. Reduction in bubble size does give a signi® cant additional contribution through the associated increased interfacial area per unit volume of gas. In previous deoxygenation studies2 , the reduction in bubble size in salt water as compared to fresh water resulting from suppression of bubble coalescence in salt water gave substantial improvements in mass transfer rate.
EXPERIMENTAL An outline ¯ ow scheme for the experimental facility built for this study is shown in Figure 2. The high intensity contactor or the packed column were connected in turn to the feed, sample and analytical system for comparison. Feed solutions of aromatics in water were made up batchwise in a 10,000 l stainless steel tank. Complete solution was ensured by pumped recirculation for about 2 hrs. For safety reasons, toluene and ethyl benzene were used in preference to benzene. Aromatic concentrations in the feed ranged from 29 to 194 mg l ± 1 in different runs. These are higher than the levels expected on oil platforms but are possible on some gas platforms. Using these higher concentrations makes for more reliable analysis.
Figure 2. Outline ¯ owscheme of experimental facility.
Trans IChemE, Vol 76, Part A, July 1998
REMOVAL OF DISSOLVED AROMATICS FROM WATER
Water was either fresh from the mains or a solution of 3.7% wt commercial granular salt (NaCl) in mains water. Maximum water feed rate was 1.5 kg s ± 1 (5.4 m3 hr ± 1 ), the limit of the centrifugal feed pump. Stripping air was metered from the laboratory compressed air system. Gas/ liquid (G/L) volumetric ratios up to 35 were used, depending on the contacting device. The packed column, an existing one incorporated into the new facility, was 220 mm in diameter with 2.15 m depth of 16 mm polypropylene Pall rings. This was operated countercurrently with liquid distributed evenly through a multipoint drip type distributor. The high intensity contactor (Figure 1) utilized a permeable ceramic tube, 32 mm internal diameter by 250 mm long for gas injection. Water was fed in tangentially through two feed points at one end. Gas essentially free of water drops leaves from a central top outlet and water free of gas from the base. Sample streams of feed and product water and of discharged air were passed through on-line analysers. Aromatics in water were measured by ¯ uorometry (Turner Designs Field Fluorometer Model No 10-AU-005). A free falling sample stream in a windowless cell is subjected to a narrow band excitation light beam and the intensity of the resulting emission from the aromatic molecule is measured. By using appropriate optical ® lters, the excitation and emission signals can be optimized for the compounds of interest. Here an excitation wave length of 254 nm and emission wave length of 289 nm were used. The aromatics, either toluene or ethyl benzene, were present singly in the feed and product solutions. Considerable effort was given to developing calibration procedures and overcoming problems in the original instrument. Critical factors in ensuring reliable calibration and on-line data include precise control of sample temperature and internal cell temperature. Replacement of the original acrylic light feedback pipe with a quartz one eliminated drift caused by uv degradation of the acrylic. Aromatics concentrations in the outgoing air were measured with a photo ionization instrument (ELE Instruments Portable Solvent Detector No 300-100). Cross checks on the output of this instrument were made on a chromatograph (Perkin Elmer 8500 series with ¯ ame ionization detector) using batchwise sample bags. The photo ionization instrument proved to be very sensitive to sample conditions. Under carefully controlled conditions the toluene balance from simultaneous data from the water and air analysers was consistent within 6 10%. For routine runs, the ¯ uorometer data for concentrations of aromatic in the water feed and product, the streams of most interest, were used to calculate mass transfer performance. Water, feed and product temperatures which could differ by up to 2°C were measured and averaged. These temperatures were steady over a run but the absolute level depended on weather conditions as the feed tank was located outside. RESULTS AND DISCUSSION Analysis Measured concentrations of aromatic in feed and product water, together with water and air ¯ ow rates, were used to calculate the number of overall liquid transfer units, NO L , achieved in a given run. As the incoming air was free of Trans IChemE, Vol 76, Part A, July 1998
565
aromatics, NO L is given by1 1 : x2 12 x1
1 1 1 S S NOL = 1 12 S where the stripping factor S is given by H Gm S= P Lm
loge
(2)
(3)
The removal of toluene and ethyl benzene has been assumed to be liquid ® lm diffusion controlled as the calculated Henry’ s Law constants, H, are suf® ciently high, 36.2 MN m ± 2 mole fraction and 46.5 MN m ± 2 mole fraction respectively in distilled water at 25°C. Thus in equation (4) for the overall volumetric mass transfer coef® cient based on the liquid side, 1 1 1 = (4) 1 KL a kL a H kG a the last term is relatively small and is ignored. This simpli® cation is adopted in the design of columns1 8 for stripping of these materials from water. It also appears valid for conventional bubble column according to the somewhat limited literature data on kG and kG a for such columns, e.g. that reviewed by Deckwer1 9 including the experimental data of Mashelkar and Sharma2 0 . The validity for the present contactor where enhanced values of kL a are expected is discussed later. Calculated values of HO L depend on the values taken for H which varies with temperature and salinity. The two values quoted before were derived from published data on solubility1 2 and vapour pressure1 6 of the two aromatics. Solubilities measured by Sutton and Calder1 2 show a signi® cant reduction in solubility in sea water as compared to distilled water. The resulting increases in calculated H values are over 40% and this effect has been taken into account here. The temperature dependencies of solubility and hence of H were based on experimental solubilities at 25° and 0°C reported by Polak and Lu1 3 . These data were for distilled water. In the absence of measured data, the same dependency was assumed for salt water. Calculated values of HO L are most sensitive to uncertainties in H at lower values of G/L. One basis for comparing the mass transfer performances of the high intensity contactor and the packed column is the equivalent height of an overall liquid transfer unit, HO L , where HOL =
Height of mass transfer zone NOL
(5)
This is a somewhat limited and approximate comparator compared to the volumetric mass transfer coef® cient considered later but is applied in view of the widespread use of the height of a transfer unit as a measure of the mass transfer performance of conventional packed columns. As in the previous study of deoxygenation2 , the height of the mass transfer zone for the high intensity contactor is taken as the tube length, 0.25 m. For the column it is obviously the packed height. Comparison of HO L values is of practical interest as regards height requirements, an important factor in installing equipment on offshore platforms
566
WALDIE and HARRIS
especially in retro-® tting. The present data on HO L for the contactor is however limited as regards extrapolation to separation processes requiring more transfer units than the relatively simple stripping of volatile aromatics. Further work is needed to establish how performance varies with tube length, as discussed later. Strictly, equation (5) applies to a pure countercurrent ¯ ow situation whereas ¯ ow in the new contactor is more complex involving cross and counter¯ ow. However, even in a packed column there are conditions at high L/G where some recirculation of gas is likely and yet the height of a transfer unit concept is still applied for simplicity. That simple concept is used for convenience here together with the volumetric mass transfer coef® cient, KL a, where Lm KL a = (6) rm HOL KL a is a more comprehensive comparator than HO L as it takes into account both the height requirement and the liquid ¯ ux capabilities of the two devices. The in¯ uence of the ¯ uid ¯ ow model assumed for the contactor on calculated KL a values is considered later. In these mass transfer calculations, no deductions were made for potential contributions from end effects in the packed column. The packing performance data may therefore be somewhat overstated. End effects seem less likely in the high intensity contactor. Comparison of HO L Values An overall comparison of HO L results for the high intensity contactor and the packed column is presented in Figure 3. The contactor gives lower values of HO L at all gas/liquid ratios, the improvement being particularly marked when treating salt water. For example on salt water at G/L of 3.0, the ® ne bubble version of the contactor gives HO L values some 7±10 times better than the packed column.
Figure 3. Comparison of HOL values for high intensity contactor and packed column in removal of toluene from water.
The relative effect of water salinity on the two devices is interesting. Changing from fresh to salt water on the packed column introduces a foaming problem which promotes ¯ ooding. This prevented operation of the column with 75 l/min of salt water at G/L ratios greater than 2.0. HO L was consequently very high at 2 m. The new contactor, however, performs much better on salt water than fresh water. This con® rms the trends reported previously2 for deoxygenation of water in the contactor and in a packed column. The improvement with salt water in the contactor was attributed to suppression of bubble coalescence by dissolved salt leading to smaller mean bubble sizes and greater interfacial area. More recent1 0 bubble size distribution data for the outlet of a shorter version of the contactor helped quantify that effect. The presence of an electrolyte such as salt creates concentration and surface tension gradients which promote the stability of bubbles in liquids1 4 . As produced water is inevitably saline, often to levels much higher than the near sea water salinity used here, then this effect is particularly bene® cial. In contrast, the same concentration and surface tension mechanisms promote foam persistence in the packed column and hence reduce the maximum liquid ¯ ux. Measured Sauter mean diameters at the exit of the shorter contactor1 0 were 285 mm and 150 mm respectively for fresh and salt water. Whilst these are unlikely to represent the absolute values of mean bubble size in the present version, they con® rm the strong in¯ uence of salinity. A reduction from 285 to 150 mm approximately doubled1 0 the value of KL a. The signi® cance of bubble size in the high intensity contactor has been further demonstrated here by changing the initial bubble size range. This was done by using a permeable tube with generally coarser pores, the nominal mean pore size being 90 mm as compared to 10 mm. Quantitative data on the relative bubble sizes is not yet available but the effect on HO L in fresh water is quite marked (Figure 3). It will be noted in Figure 3 that for the ® ne bubbles in a given water two plots of HO L data are shown. These represent respectively the best, i.e. lowest, and worst HO L values obtained in different series of runs. In a given series of runs the gas rate was progressively changed at a constant liquid rate, usually 1.5 kg s2 1 . Resultant HO L values for such a series then decrease systematically with increasing G as shown (Figure 3). More detailed data for the coarse bubbles showing maximum and minimum trend lines and a mean line are given in Figure 4. The extent of the offset between different series of runs is too great to be attributed to temperature differences which could affect HO L through the temperature dependence of diffusivity. Following extensive work on the ¯ uorometry analysis technique, step changes in say analyser sensitivity from series to series appear unlikely. The possibility of a change in the mean bubble size from series to series is to be investigated once techniques for measuring bubble sizes within the actual contactor have been developed. For the present, even the worst HO L data are signi® cantly better than those for the packed column. Fewer runs were done with ethyl benzene the purpose being mainly to establish that performance of the contactor was not speci® c to toluene. HO L values decrease systematically with increase in G at constant L but the absolute Trans IChemE, Vol 76, Part A, July 1998
REMOVAL OF DISSOLVED AROMATICS FROM WATER
Figure 4. HOL for high intensity contactor with coarse bubbles in separation of toluene from fresh water.
values are generally lower than for toluene at the same liquid rate (Figure 5). As the diffusivity of ethyl benzene in water is lower than that of toluene, HO L values were expected to be somewhat greater for ethyl benzene. The few runs done at liquid rates lower than the standard 1.5 kg s2 1 gave higher values of HO L . This is in line with theoretical expectations from equation (1) of increased mass transfer rate in an enhanced acceleration ® eld and an expected decrease in bubble size. More data are needed to establish a quantitative relationship. The systematic decrease in contactor HO L with increasing G is similar to the trend found previously2 for deoxygenation. There HO L varied as G 2 n with values of 0.56 and 0.67 for the exponent n for salt and fresh water respectively. Here there is more variation in n from run series to run series for a given water, the values ranging overall from 0.50 to 0.92. As for deoxygenation, the dependence of HO L on G
567
is more pronounced with fresh water. This systematic decrease in HO L with G for the contactor contrasts with the behaviour of a packed column where HO L decreases initially but eventually reaches a near plateau value. The 0.83 kg s± 1 data for fresh water (Figure 3) in the packed column is an example. Most published data on height of a transfer unit for packings apply to this plateau region, G/L values for many packed column applications typically exceeding 10. The present comparison of the contactor with a packed column applies strictly to one particular type and size of packing and the generality of this comparison may be queried. First it should be noted that the present measured HO L values for the packed column are about 40% better than those predicted by applying Onda’ s correlations1 5 . Those correlated experimental data to within 6 20%, although Pall rings were not available for those experiments. One factor which will have led to some overstatement of the present packing performance is the omission of a correction for column end effects. Experimental determination of that correction would have required a series of runs with different depths of packing. Other evidence that the present packed column HO L data do not understate the performance of a typical random packed column comes from later work on the same toluene/water system using a 1 metre diameter column with a nominally more advanced metal packing. Comparison of Volumetric Mass Transfer Coef® cients, KL a A comparison of the performance of the high intensity contactor and packed column in terms of mass transfer rate per unit of contacting volume is given by the ratio of the KL a values (Table 3). KL a values were derived from HO L data and equation (6). These KL data therefore are based on the countercurrent plug ¯ ow approximation for water and gas ¯ ow within the contactor and the packed column. For the contactor where, as discussed previously, there were apparent changes in performance from run series to run series the worst and a mean of best and worst HO L Table 3. Comparison of mass transfer performance of high intensity contactor and packed column on volumetric basis. KL a Contactor KL a Packed Column Fresh Water Column feed rate* l/min CONTACTOR Fine bubbles
Trans IChemE, Vol 76, Part A, July 1998
75
50
75
G/L
Mean Worst
2.0 2.0
Mean Worst Mean Worst
3.3 3.3 5.0 5.0
231 154 240 170
175 118 208 148
3.3 5.0
98 110
75 95
Coarse Bubbles Mean Mean
Figure 5. HOL for high intensity contactor in separation of ethylbenzene from fresh water.
50
Salt Water
344 285
* Contactor feed rate 90 l/min (1.5 kg s ±1) throughout.
668 589 711 654
568
WALDIE and HARRIS
values for a given G/L value have been used. Thus even with the best results excluded, the high intensity contactor can achieve mass transfer rates per unit volume some several hundred times greater than a packed column. That is due to the lower equivalent HO L but more signi® cantly to the much higher liquid ¯ uxes per unit cross sectional area. The initial results at different liquid rates (Figure 5) suggest that at rates above 1.5 kg s ± 1 still higher KL a values would be achievable, the increased ¯ ux effect being supplemented by reduction in HO L . The ¯ ow patterns of gas and liquid within the contactor are clearly more complex than for countercurrent plug ¯ ow. There will be some initial cross ¯ ow of gas through the liquid, some co-current ¯ ow as bubbles begin to travel with the liquid and some countercurrent ¯ ow as gas and liquid leave from opposite ends of the tube. Effective mean mass transfer driving force will therefore be less than that for countercurrent plug ¯ ow. Thus true KL a values for the contactor will be higher than those reported here. Substantial further work is needed to quantify the actual ¯ ow patterns and develop a more exact model for the mass transfer processes. The potential extent of the dependence of KL a on the ¯ ow model assumed can be seen by treating the contactor as a perfect mixer. This model has been used in the analysis of some more conventional bubble columns1 9 ,2 1 ,2 2 . Taking for example a contactor run in salt water with ® ne bubbles and water and air rates of 90 and 450 l/min respectively, KL a calculated for the plug ¯ ow model is 10.1 s ± 1 . For the same feed and product compositions, KL a for perfect mixing is either 13.9 or 28.7 s ± 1 depending on the mean mass transfer driving force DC in equation (7) for a perfectly mixed contactor. KL a =
1 Ci 2 Co Åt DC
(7)
In modelling studies of oxygenation in a perforated plate bubble column2 2 and a packed bubble column2 1 , the compositions of the gas, air or oxygen, were essentially unchanged through the columns. DC was therefore taken as C * 2 Co , C * being the liquid concentration in equilibrium with air or oxygen respectively. Here the toluene partial pressure in the air increases through the contactor from zero to the outlet value. Taking C * as that for equilibrium with the inlet gas, i.e. zero, gives the lower value of KL a for the perfectly mixed case, 13.9 s ± 1 . Taking C * as the mean of the equilibrium values for inlet and outlet gas gives the higher KL a value, 28.7 s ± 1 . The latter mean concentration appears more reasonable but yields a KL a value which is 2.84 times higher than the value from the countercurrent plug ¯ ow model. Thus the countercurrent plug ¯ ow values of KL a for the contactor used in Table 2 are, if anything, conservative. From equation (5) and the experimental variation of HO L with G (Figures 3, 4 and 5). KL a 3 G n
(8)
n ranges from 0.50 to 0.92. The range of n values was narrower in the previous study of deoxygenation in the contactor2 . For classical bubble columns where gas is bubbled through a relatively slow axially ¯ owing liquid, the equivalent exponent is typically1 7 about 0.7 with some values as high as 0.82.
The high intensity contactor has been compared experimentally with a conventional packed column in which the liquid is the dispersed phase. Other potential competing equipment includes various forms of bubble column where, as in the high intensity contactor, the gas is the dispersed phase. Comparison of present KL a values with those for bubble columns is therefore of interest. The highest values of KL a for bubble columns found in the literature are those reported by Wang and Fan2 1 for a packed bubble column and by Voigt and SchuÈgerl 2 2 for a multi stage perforated plate bubble column. Both groups studied oxygenation. Wang and Fan2 1 con® rmed that KL a values for an unpacked column could be increased signi® cantly by installing motionless mixer elements of corrugated metal sheet. KL a values for the packed column reached 0.23 s ± 1 or 0.16 s± 1 depending on whether the perfectly mixed or plug ¯ ow model was applied to the measured data. That column was operated co-currently in upward ¯ ow. The KL a values are thus lower by a factor of 50 or more than present values. Also, if co-current operation were used on the present duty, much higher G/L ratios would be needed for a given separation duty than in the high intensity contactor to maintain a positive driving force throughout the column. This latter problem would not arise with the perforated plate bubble column of Voigt and SchuÈgerl2 2 which was operated countercurrently. For salt water, KL a values based on the perfect mixer model were mainly below 0.3 s± 1 except for one value of 1.66 s ± 1 which was attributed to foam formation. Film Control Present values of HO L and KL a are based on the assumption of liquid ® lm control, reasonable for a conventional packed column or bubble column at the level of Henry’ s Law coef® cient involved. The anticipated improvement in kL in the new contactor could, if suf® ciently great, invalidate that assumption (equation (4)). To check this possibility, estimates of the gas side mass transfer coef® cient inside a bubble have been made for comparison with the overall coef® cient KL . Transfer by molecular diffusion only is assumed in view of the relatively small bubbles. Equations (8) and (9) are equivalent to those developed by Calderbank9 for the internal ® lm coef® cient for a ¯ uid drop where the concentration of diffusing species at the interface is constant. H kG = 2 H kG <
DG pt
1 p 2 DG 3 R
0.5
(9) (10)
Equation (9) applies to short contact times where the fractional approach to equilibrium is <0.4 and (10) to longer times where there is a closer approach to equilibrium. Applying equation (10) and taking R as 75 mm from the Sauter mean diameter and the speci® c area, a, from that diameter and an assumed fractional gas hold up of 0.05, gives a value of H kG a of about 700 s ± 1 . Equation (9) with t of 0.026 s, an upper limit for the mean gas residence time, gives about 40 s ± 1 . The residence time of the gas when present as individual bubbles will be less than this Trans IChemE, Vol 76, Part A, July 1998
REMOVAL OF DISSOLVED AROMATICS FROM WATER giving a higher value of H kG a. In practice, the interface concentration will change with time but these values of H kG a even if very approximate are suf® ciently greater than the measured overall KL a values to suggest that the process is still liquid ® lm controlled. Equipment Size and Scale-up The KL a ratios in Table 3 show that the internal contacting volume of a high intensity contactor could be over two orders of magnitude smaller than that of a packed column on the same duty. The ratio of overall equipment volumes would not be as large but still be signi® cant. Comparison with literature data on KL a for bubble columns, including packed bubble columns, suggests that the new contactor would also be signi® cantly more compact than those. Scale-up for larger water ¯ ow rates could involve increasing the tube diameter and/or using several tubes in parallel. The latter would be achieved most compactly by mounting the tubes in a common shell as is done for de-oiling hydrocyclones. The water ¯ ux in the present tube is higher than in some commercial de-oiling hydrocyclones and so an acceptably compact arrangement should be possible. Research is needed to establish how tube diameter and tube length affect performance. With the present tube length, 0.25 m, somewhat over 1 transfer unit has been achieved. Mass transfer per unit length will presumably decrease with tube length due to frictional losses and eventual decay of the swirling water ¯ ow pattern. The device as currently con® gured is therefore most suited to relatively simple separations such as stripping of aromatics or oxygen from water which require few transfer units. Energy Consumption As in deoxygenation2 , the substantial reduction in contacting volume inevitably incurs an increase in energy consumption over that for a packed column. The increase may however be less pronounced for produced water as that could be available under pressure from the primary separators. At water and air ¯ ow rates of 90 l/min and 300 l/min, i.e. G/L = 3.33, pressure drops were 5.5 ´ 105 N/m2 and 0.86 ´ 105 N/m2 respectively. Theoretical minimum energy requirements per m3 water are therefore 0.153 Kw-hr and 0.079 Kw-hr respectively giving a total of 0.232 kW-hr/m3 . Pumping and gas compression losses need to be added. These data apply to one particular con® guration which was not optimized. There are trade-offs between contactor volume and energy consumption involving for example the choice of G/L ratio and the permeability of the tube and hence bubble size. Energy consumption could be reduced if produced water were taken from a pressurized primary separator and further reduced if stripping were done by hydrocarbon gas from the primary separators. Comparing the high intensity contactor with the bubble columns giving the highest reported KL a values shows that the contactor gives signi® cantly higher values of KL a for a given energy input per unit volume of water treated. For example, the perforated plate bubble column of Voigt and Trans IChemE, Vol 76, Part A, July 1998
569
SchuÈgerl2 2 even when giving enhanced KL a values due apparently to foaming required some 2.8 times more energy per unit volume of feed water per unit KL a. Without foaming, the ratio would be over 10. CONCLUSIONS The high intensity contactor has the potential for signi® cant reductions in equipment volume and weight as compared to a packed column in the removal of dissolved aromatics from water by gas stripping. The expected insensitivity of the contactor to marine motion and tilt would be a further advantage on a ¯ oating production system. Results on stripping of aromatics con® rm several aspects of contactor performance characteristics found in a previous study of deoxygenation. The volumetric mass transfer coef® cient, KL a, can be several hundred times higher than that for a conventional packed column. The high intensity contactor also compares favourably in that respect with perforated plate and packed bubble columns in the literature. Performance is enhanced in saline water due to suppression of bubble coalescence. Experiments with two levels of initial bubble size con® rm the importance of bubble size on mass transfer performance. The volumetric mass transfer coef® cient, KL a, varies as G n with n here ranging from 0.50 to 0.92, a wider range than that for deoxygenation. Reasons for this variation and other fundamental aspects including bubble size distribution, gas hold up, capacity limits and scale-up require further study. Speci® c energy consumption is higher than for a packed column but could be reduced in produced water treatment if use can be made of already pressurized liquid feed and perhaps gas from the primary separators. NOMENCLATURE a db DG DL G Gm g H HOL kG kL KL L Lm n NOL P r R S t Åt Sc x1 , x 2
speci® c surface area, m ±1 bubble diameter, m gas phase diffusivity, m2 s ±1 liquid phase diffusivity, m2 s ±1 gas ¯ ow rate, m3 s ±1 gas molar ¯ ux, kmol m ±2 s ±1 gravitational acceleration, m2 s ±1 Henry’ s Law coef® cient, Nm ±2 mole fraction ±1 height of overall liquid transfer unit, m gas side mass transfer coef® cient, mole fraction m3 N ±1 s ±1 liquid side mass transfer coef® cient, m s ±1 overall liquid mass transfer coef® cient, m s ±1 liquid ¯ ow rate, m3 s ±1 liquid molar ¯ ux, kmol m ±2 s ±1 exponent number of overall liquid transfer units total absolute pressure, N m ±2 radial distance, m bubble radius, m stripping factor time, s mean liquid residence time, s Schmidt number mole fractions of aromatic in product and feed
Greek rL , rG rm v
letters liquid and gas densities, kg m ±3 molar density of liquid,kmol m ±3 rotational speed, s ±1
570
WALDIE and HARRIS REFERENCES
1. Waldie, B., 1995, Novel high intensity gas/liquid contactor, Proc 1st Int Conf on Science, Engineering and Technology of Intensive Processing (Akay, G. and Azzopardi, B. J. (eds)) University of Nottingham, 5. 2. Waldie, B., 1996, A new high intensity contactor for de-oxygenation of water, Trans IChemE, 74(A2): 183±189. 3. Hansen, B. R. and Davies, S. R. H., 1994, Review of potential technologies for the removal of dissolved components from produced water, Trans IChemE, 72(A2): 176±8. 4. Stephenson, M. T., 1992, Components of produced water: a compilation of industry studies, J Pet Tech, 44: 548±550 and 602±603. 5. Callaghan, D. and Baumgartner, W., 1990, Characterization of residual hydrocarbons in produced water discharged from gas production platforms, SPE Paper 2088. 6. Balasundaram, V., Porter, J. E. and Ramshaw, C., 1989, Process intensi® cation: a rotary seawater deaerator, Proc Symp on Separation of Gases (Royal Soc of Chemistry), 306. 7. Miller, J. D. and Van Camp, M. C., 1982, Mining Eng, 1: 57. 8. Beenackers, A. A. C. M. and van Swaaij, W. P. M., 1976, Mass transfer and chemical reaction in a gas-liquid cyclone reactor, Proc 4th Int 6th Europ Symp on Chemical Reaction Engineering, (Dechema Frankfurt), 260. 9. Calderbank, P. H., 1967, Mass transfer, Chapt 6, in Mixing Theory and Practice, Vol II, Uhl, V. W. and Gray, J. B. (Academic Press). 10. Waldie, B. and Harris, W. K., 1997, Stripping performance of a new high intensity contactor, Proc Distillation and Absorption 97’ IChemE Symp Series No 142, Vol 1: 371±378 Ed Darton, R. 11. Treybal, R. E., 1980, Mass Transfer Operations, 3rd ed (McGraw Hill, New York). 12. Sutton, C. and Calder, J. A., 1975, Solubility of alkylbenzenes in distilled water and seawater at 25°C, J Chem and Eng Data, 20: 320. 13. Polak, J. and Lu, B. C. Y., 1973, Mutual solubilities of hydrocarbons and water at 0 and 25°C, Can J Chem Eng, 51: 4013.
14. Marrrucci, G., 1969, A theory of coalescence, Chem Eng Sci, 24: 975. 15. Onda, K., Takeuchi, H. and Okumoto, Y., 1968, Mass transfer coef® cients between gas and liquid phases in packed columns, J Chem Eng Jpn, 1: 56. 16. Perry, R. H. and Chilton, C. H., 1979, Chemical Engineers Handbook, 5th Ed (McGraw Hill, New York). 17. Akita, K. and Yoshida, F., 1973, Ind Eng Chem Proc Dev, 12: 76. 18. Striggle, R. F., 1987, Random packings and packed towers, (Gulf Publishing Co, Houston). 19. Deckwer, W. D., 1992, Bubble Column Reactors, (J. Wiley, Chichester). 20. Mashelkar, R. A. and Sharma, M. M., 1970, Mass transfer in bubble and packed bubble columns, Trans IChemE, 48: T162. 21. Wang, K. B. and Fan, L. T., 1978, Mass transfer in packed columns packed with motionless mixers, Chem Eng Sci 33: 945. 22. Voigt, J. and SchuÈgerl, K., 1979, Absorption of oxygen in countercurrent multistage bubble columns ± I, Chem Eng Sci, 34: 1221.
ACKNOWLEDGEMENTS This project was supported by EPSRC, DTI and eight operating companies as part of the MTD managed programme on Treatment of Water Offshore, III. Craig Bell and Ron Millar are thanked for their contributions to the experimental runs and rig construction.
ADDRESS Correspondence concerning this paper should be addressed to Professor B. Waldie, Department of Mechanical and Chemical Engineering, HeriotWatt University, Edinburgh EH14 4AS, UK. The manuscript was received 27 March 1998 and accepted for publication after revision 15 May 1998.
Trans IChemE, Vol 76, Part A, July 1998