- Email: [email protected]
Design of a new compact scrubber for deodorisation
Available online at www.sciencedirect.com
Chemical Engineering and Processing 47 (2008) 1793–1798
Design of a new compact scrubber for deodorisation A. Couvert a,∗ , C. Sanchez b , A. Laplanche a , C. Renner c , J.-P. Levasseur d a
ENSCR, UMR CNRS 6226 - CIP, 263 Avenue du G´en´eral Leclerc, 35700 Rennes Cedex, France Burgeap, Pˆole air et Sant´e publique, Parc de la Bastide Blanche, Bˆat. 6, 13127 Vitrolles, France c Anjou Recherche – Veolia Environnement, Chemin de la digue, BP 76, 78603 Maisons Laffitte, France d Direction Technique, Veolia Eau, 1 place Montgolfier, 94417 Saint-Maurice Cedex, France b
Received 28 June 2007; received in revised form 31 August 2007; accepted 1 October 2007 Available online 12 October 2007
Abstract A new compact scrubber was designed to remove odorous compounds from WWTPs (named Aquilair PlusTM , for Compact Reticulated Contactor). Hydrodynamics and mass transfer parameters (P, a, kL a and kG a) were quantified versus superficial gas and liquid velocities (USG and USL ). This study showed that all these parameters increase with USG and USL , except for kG a which is almost independent of USL . Thanks to the lab-study, a half-size industrial pilot plant was designed. First results are encouraging. Finally, an economic study showed that Aquilair PlusTM allows a gain of 15–40% compared to traditional packed towers. © 2007 Elsevier B.V. All rights reserved. Keywords: Absorption; Scrubber; Process intensification; Hydrodynamics; Mass transfer; Design
1. Introduction Public concern over odours from wastewater treatment plants (WWTPs) is increasing because of the development around existing, or in construction works. Species identified as responsible for odours in WWTPs are mainly reduced sulphur and nitrogen compounds. Their removal is most of time realised by chemical scrubbing in a two or three stage packed tower (PT) system. This approach has been showing good results for many years [1,2], but has also revealed high investment and working costs (high residence time implying the construction of high and wide towers, important reactive consumption mainly due to competitive CO2 absorption). In the last decades, structured PT [3,4] have gained increased attention because of their better performances in terms of gas–liquid contact and distribution, coupled with their higher bed void fraction resulting in lower pressure drops [5]. Filled with random or structured packing, most of the studies on PT found in literature focus on counter-current flow applications; nevertheless, when absorbed species react in the liquid phase according to a fast irreversible reaction, and especially when high fluid velocities are implemented, the cocurrent flow configuration is preferable. Moreover, co-current ∗
Corresponding author. Tel.: +33 2 23 23 80 48; fax: +33 2 23 23 81 20. E-mail address: [email protected] (A. Couvert).
0255-2701/$ – see front matter © 2007 Elsevier B.V. All rights reserved. doi:10.1016/j.cep.2007.10.006
working avoids flooding, allowing higher flexibility [6]. Then, some researchers paid attention on static mixers (SM) because of their excellent gas–liquid contact due to a high turbulence rate inside [7,8]. Different types of static mixers were designed (Sulzer, Lightnin, Ross, Kenics, etc.) providing good results in terms of mixing (their first application), but also in terms of mass transfer, which is interesting for scrubbing. Unfortunately, these gas–liquid contactors generate too high-pressure drop to be industrially viable. Consequently, the objective of this study was to design a compact apparatus, able to generate high interfacial areas and good mass transfer coefficients (high efficiency), without driving to a high-energy consumption. In order to conclude on the best operating conditions and design parameters for scale-up, pressure drop, interfacial area and volumetric mass transfer coefficients were studied in a laboratory-scale pilot unit. This step permitted to design a half-sized industrial pilot plant. 2. Experimental The lab-reactor (named Aquilair PlusTM ) consists of a transparent PVC pipe of 0.32 m length and 0.025 m inner diameter (volume of 1.53 × 10−4 m3 ), through which flow patterns as well as the type of dispersion can be observed (Fig. 1). The contactor is a new wire mesh packing structure which geometric characteristics are given in Table 1. Meshes are 12.5 mm side, made
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Fig. 1. Scheme of the experimental set-up.
of stainless steel. Then, the structure is bent (to avoid too much compactness, and so, high pressure drop) and rolled up so that it can be inserted in an empty tube of 0.025 m diameter. The packing geometric specific area, which is equal to 176 m2 m−3 of reactor, takes the surface offered by the packing itself and by the column wall into account; its packing factor calculation is higher than the ones of conventional packing, which supposes that a greater interfacial area will be offered to gas–liquid contact; the void fraction is also very high (97.5%) in order to minimize pressure drop. The volume gas flow rate QG , controlled with a rotameter, ranged from 10 to 50 m3 h−1 (superficial gas velocity USG from 5.6 to 28 m s−1 ). The liquid was injected into the reactor at liquid flow rates QL ranging between 0.0275 and 0.0975 m3 h−1 (superficial liquid velocity USL between 0.016 and 0.055 m s−1 ). The fluids circulated in vertical descendant (VD) co-current flow inside the pipe. pH and temperature were continuously measured with specific probes. The pressure drop P within the reactor was measured with two manometers: the first one was filled with water (for the low pressure drop values) and the second one was filled with mercury (for the high-pressure drop values). They were connected just before and after Aquilair PlusTM and the pressure drop was measured for the different liquid and gas flow rates mentioned previously. The interfacial area a was determined by transferring CO2 (CG,CO2 = CG,A = 4000 ppmv) into an aqueous sodium hydroxide solution (CL,OH − = 1N). Details are given in the Annex. The CO2 gas concentrations (inlet and outlet) were measured with an IR analyser type Beryl 100 from COSMA and in the liquid phase by the measurement of the alkalinity and total alkalinity (AFNOR NF T 90-036, July 1977). The liquid-side volumetric mass transfer coefficient kL a was measured thanks to the implementation of a physical transfer
of gaseous 2-butanone in an aqueous phase (water, pH equal to 7). 2-Butanone, chosen because of its small solubility (Henry’s constant = 6.7 × 10−2 atm L mol−1 at 25 ◦ C), was injected as a liquid in the gas flow with a syringe distributor, vaporised in it in order to reach a concentration of about 50 mg m−3 in the gas phase (see Eq. (5) in Appendix B). 2-Butanone concentrations in the gas phase (inlet and outlet) were indirectly measured after trapping it in a 2,4-dinitrophenylhydrazine (2,4-DNPH) solution. Indeed, the trapping yields for 2-butanone exceed 99%. This method is based on the reaction between 2-butanone and 2,4-DNPH, giving an hydrazone which is then analysed by HPLC (C18, mobile phase composed of 80% methanol and 20% water, UV detection for a wavelength of 365 nm). The gas-side volumetric mass transfer coefficient kG a was obtained by implementing an irreversible and instantaneous reaction between NH3 and HCl (see Eq. (6) in Appendix B). The condition that must be verified to implement this method is kG EkL /H. This was confirmed once kG a was measured (EkL /H varied from 500 to 1080, kG values range between 0.05 and 0.28 m s−1 ). Concerning the half-size pilot plant, it was of 0.3 m diameter and 0.6 m length, with slightly different internal packing characteristics (void fraction, geometric surface) from the labpilot ones. Liquid and gas were also distributed by the top of the contactor, in co-current flow. Gas velocities ranged between 5 and 20 m s−1 for a chosen liquid flow rate of 8 m3 h−1 , in order to reproduce the lab operating conditions. P was measured with pressure sensors type Cerabar S PMC71 located before and after the contactor, and the gas velocity was obtained thanks to a Pitot probe from Endress Hauser. The H2 S gas concentration was measured with Odalog specific probes (App-Tek type I), which range from 0 to 200 ppm.
Table 1 Geometric characteristics of Aquilair PlusTM
3. Results and discussion
Height (m) Internal diameter (m) Void fraction (%) Geometric surface ag (m2 m−3 ) Packing factor (m2 m−3 ) Volume (m3 )
0.32 0.025 97.5 176 6667 1.53 × 10−4
3.1. Lab-plant study Fig. 2 shows the pressure drop per reactor length unit versus superficial gas and liquid velocities. As expected, P/L increases with both operating parameters. Moreover, it can be
A. Couvert et al. / Chemical Engineering and Processing 47 (2008) 1793–1798
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Fig. 2. Evolution of pressure drop vs. superficial gas and liquid velocities in the VD configuration.
Fig. 3. Evolution of the interfacial area vs. the superficial gas and liquid velocities in the VD configuration of Aquilair PlusTM .
noticed that P/L in our reactor is three times lower than the one in the Lightnin SM, and far lower than those generated by SMV4 Sulzer SM, which can reach values up to 50,000 Pa m−1 for USG = 15 m s−1 [10]. This seems logical if one compares void fractions in the different gas–liquid contactors (78% for the Sulzer SM, 87% for the Lightnin SM, 97.5% for Aquilair PlusTM ). The more the void fraction, the less the resistance opposed to the fluid flow. Otherwise, the interfacial area a measured ranged from 600 to 2800 m2 m−3 (Fig. 3). These values are twice higher than the ones measured in the Lightnin SM and the empty tube, and far higher than the ones obtained in packed towers (around 100–300 m2 m−3 ). Moreover, the interfacial area measured in towers filled with structured packing and operating in counter-current flows can reach values ranging from 250 up to 750 m2 m−3 [11,12]. The increasing of this parameter with the superficial liquid velocity was foreseeable since when USL increases, the number of liquid drops or trickles in contact with the gas phase rises. Several authors observed the same phenomenon in different gas–liquid contactors: Benadda [13] and Lara-Marquez [14] in packed towers, Metha et al. [15] in spraying towers. In the same time, when USG increases, turbulence is more important, so that liquid trickles are broken into small drops, driving to a higher gas–liquid contact area.
Concerning mass transfer parameters, the VD configuration led to promising values of kL a coefficients, up to 0.20 s−1 (Fig. 4). As a comparison, kL a measured in the empty tube does not overcome 0.05 s−1 for USG < 15 m s−1 and USL < 0.055 m s−1 , kL a found in PT varies between 2.5 × 10−3 and 0.035 s−1 [16], whereas the values found in structured PT are expected to be higher depending on the geometric characteristics of the packing elements. As an example, Raynal et al. [17] found kL a values ranging from 0.15 up to 0.25 s−1 with their packing made of smooth stainless sheets, but at very high L /G
Fig. 4. Evolution of the liquid-side volumetric mass transfer coefficient vs. the superficial gas and liquid velocities in the VD configuration of Aquilair PlusTM .
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Fig. 5. Evolution of the gas-side volumetric mass transfer coefficient vs. the superficial gas and liquid velocities in the VD configuration of Aquilair PlusTM .
ratios (ours vary from 0.5 to 5). Otherwise, kL a values measured in Statiflo and Lightnin SM are lower than the ones obtained in Aquilair PlusTM as soon as USG ≥ 10 m s−1 [18]. Like the other parameters, kL a is an increasing function of the superficial gas and liquid velocities. Indeed, on the one hand, a increases with USG and USL , and on the other hand, the higher these two parameters, the higher turbulence in the contactor, which lets suppose higher values of kL . Finally, kG a was found to vary between 40 and 530 s−1 (Fig. 5). These values are similar to those found in other classical reactors like packed towers or venturis [19] and show that the mass transfer resistance is located in the liquid phase. kG a increases with USG , but not with USL . In 1968, Jhaveri et al. [20] have noticed the same phenomenon. 3.2. Extrapolation parameters Based on the study of hydrodynamic and mass transfer parameters, but also of efficiency performances revealed by the lab-contactor [21], a half-sized industrial pilot plant was designed in order to treat a polluted gas flow rate of 4000 m3 h−1 (Fig. 6). Here, it is important to specify that to treat the same gas flow rate in a packed tower, the volume of the contactor would be far higher; indeed, a pilot packed tower was implemented in the same area, in order to treat the same gas flow rate, and its volume was 25 times higher than the Aquilair PlusTM one. The first step was to choose a superficial gas velocity according to an economically acceptable pressure drop. This pressure drop was reached for USG = 15 m s−1 in the lab plant. At this superficial gas velocity, mass transfer parameters were satisfactory, and the sulphur compounds well removed according to the objectives defined. Taking these considerations into account, and basing the extrapolation on a contact time Tc of 0.04 s, a larger and longer reactor (0.3 m diameter, 0.6 m length) was implemented in the WWTP. Its void fraction was slightly higher (98.2% instead of 97.5%), and its geometric specific area lower (ag = 120 m2 m−3 instead of 176 m2 m−3 ) than the ones of the lab-plant; moreover, wall effects are lower in the industrial plant because of the higher diameter compared to the packing (45% in the lab-plant against 7% in the pilot plant), which lets suppose that the pressure drop would be lower in the industrial plant than in the lab-plant if all the other geometric and operating characteristics were similar.
Fig. 6. Photo of the half-size industrial plant.
First, the pressure drop was measured and compared to the one predicted by the correlation stemmed from the lab experimental data. Fig. 7 shows the variations of the pressure drop with USG for QL = 8000 L h−1 (USL = 0.03 m s−1 ). It can be observed that the experimental values are slightly lower than the ones measured in the lab-plant (see correlation of Sanchez et al. [22]). This result can be related to the higher void fraction and the lower wall effects in the pilot plant. No measurement of the mass transfer parameters (a or kL a) was done on this pilot plant, but observing the gas and liquid flows inside the contactor, it could be suspected that the dispersed regime, favourable to good mass transfer performances, was not completely reached. Then, the performances of the pilot plant in terms of sulphur hydrogen removal were characterised. As expected, the efficiency was not as good as the one found in the lab-plant (76% instead of 95% for L /G ≈ 3). This could be attributed to the mass transfer performances of the industrial plant, lower than the ones measured in
Fig. 7. Evolution of the pressure drop in the half-size industrial plant vs. the superficial velocity (QL = 8000 L h−1 ).
A. Couvert et al. / Chemical Engineering and Processing 47 (2008) 1793–1798
the lab-plant, certainly because of its higher void fraction. This could be verified by realising one test on a configuration more similar to the lab one (void fraction of 97.7%, ag = 160 m2 m−3 ). Indeed, the H2 S removal for USG = 12 m s−1 was found equal to the one measured in the lab-plant (about 90%). Unfortunately, the pressure drop was higher than the one obtained in the lab-plant. Some improvements are being realising. 3.3. Economical evaluation A study consisting in replacing a classical system in a WWTP (two or three packed towers designed to treat a gas flow rate of 100,000 m3 h−1 ) by five Aquilair PlusTM units was led by the technical department of Veolia Water. The difference between both systems comes from the fact that the Aquilair PlusTM units are located near each source of odorous compounds (four to treat sulphur compounds of the pre-treatment, biological treatment, sludge treatment and relieving post; one more to treat ammonia of the sludge treatment post), whereas the two or three packed towers are situated in a building dedicated to deodorisation. Therefore, in the classical system, the odorous gas has to be collected and carried to these premises, which implicates more ventilators and ventilation pipes. This is why the economical evaluation took the costs related to civil engineering, instrumentation, reactive and electricity consumption into account, but also ventilation expenses. This estimation led us to conclude that the gas–liquid contactor developed in this study drives to savings in terms of investment and working (15–40%), without consuming more energy, and decreasing ground hold in the same time. 4. Conclusion Hydrodynamic and mass transfer parameters were quantified in a new compact scrubber, Aquilair PlusTM . This gas–liquid contactor proved to be economically interesting in term of pressure drop compared with conventional structured packing reactors, but also in terms of performances (high-interfacial areas and liquid-side volumetric mass transfer coefficients). Moreover, Aquilair PlusTM presents the plus to be compact, therefore to limit ground hold, and costs related to it. Thanks to the lab study, a half-size industrial plant was designed. First results concerning the pressure drop it generates are promising. Efficiency experiments will confirm the potentiality of Aquilair PlusTM as a scrubber for gas containing sulphur compound. Finally, a comparison between a classical system and the Aquilair PlusTM solution proved that our reactor is economically more advantageous than two or three conventional packed towers. Appendix A. Nomenclature
A a C D
solute (CO2 , 2-butanone or NH3 ) volumetric interfacial area (m2 m−3 ) concentration (mol m−3 ) diffusivity coefficient (m2 s−1 )
E G H k kL , kG L L n N Q S Us v V
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enhancement factor mass gas flow rate (kg s−1 ) Henry’s constant (Pa m3 mol−1 ) kinetic constant ((mol m−3 )1−n s−1 ) film mass transfer coefficient (m s−1 ) contactor length (m) mass liquid flow rate (kg s−1 ) reaction global order molar mass flux transferred (mol s−1 ) volume flow rate (m3 s−1 ) section area of Aquilair PlusTM (m2 ) superficial velocity (m s−1 ) reaction rate (mol m−3 s−1 ) reactor volume (m3 )
Symbols P pressure drop (Pa) μ dynamic viscosity (Pl) Index G, L i, o
gas, liquid inlet, outlet
Appendix B. Annex B.1. Interfacial area The interfacial area a was determined by transferring CO2 into an aqueous sodium hydroxide solution. It was deduced from Eq. (3) considering the reaction between CO2 and NaOH as rapid (Ha > 5, see Eq. (1)), irreversible (v = kOH− CL,A CL,OH− with kOH− = 6.4 m3 mol−1 s−1 ) and of pseudo-first order (CL,OH− CL,A , so that CL,OH− ≈constant): Ha2 =
DA kOH− CL,OH−
(1)
kL2
Then, the molar mass flux of CO2 transferred can be written as the following equation: ∗ −C N = kL aVE (CL,A L,A )
(2)
In such conditions, CO2 concentrations in the liquid are considered equal to zero (all the CO2 absorbed is consumed by the rapid reaction), and Eq. (2) gives the following equation: a=
V
N
∗ ∗ ∗ DA kOH− CL,OH− (CL,A,i /CL,A,o )/ ln((CL,A,i
∗ − CL,A,o ))
(3) A: solute (CO2 here), N: molar mass flux of CO2 trans∗ : CO ferred (mol s−1 ), V: contactor volume (m3 ), CL,A 2 liquid concentration at the gas–liquid interface (mol m−3 ), kOH− = kinetic constant of the reaction between CO2 and NaOH (1 N), DA = 1.62 × 10−9 m2 s−1 : CO2 diffusivity in sodium hydroxide calculated from the CO2 diffusivity in water DA/water μwater = DA/solution μsolution [9]
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In these conditions, we assume that gas-side resistance is negligible (HA kG kL with HA = 2630 Pa m3 mol−1 Henry’s constant of CO2 ). Then interfacial concentration can be calcu∗ ) and molar mass transfer lated with Henry’s law (PA = HA CL,A flux from the variation of the CO2 concentration in the gas phase and in the liquid phase (Eq. (4)). N = QL (CL,A,o − CL,A,i ) = QG (CG,A,i − CG,A,o )
(4)
B.2. Liquid-side volumetric mass transfer coefficient The liquid-side volumetric mass transfer coefficient kL a was measured thanks to the implementation of a physical transfer of gaseous 2-butanone in an aqueous phase. In this case, the mass transfer resistance is located in the liquid film, and the 2-butanone liquid concentration at the inlet is zero; then kL a is calculated from the following equation: kL a =
N ∗ E.V.(CL,A,i
∗ − (CL,A,o
∗ ∗ − CL,A,o )/ ln(CL,A,i /(CL,A,o − CL,A,o ))
(5) A: solute (2-butanone here); E: enhancement factor = 1 when there is no reaction in the liquid phase; N: molar mass flux of 2-butanone transferred (mol s−1 ); CL,A,o : 2-butanone liquid con∗ centration in the liquid phase at the outlet (mol m−3 ); CL,A,i , ∗ CL,A,o : 2-butanone liquid concentration at the gas–liquid interface (inlet and outlet) (mol m−3 ) B.3. Gas-side volumetric mass transfer coefficient The gas-side volumetric mass transfer coefficient kG a was obtained by implementing an irreversible and instantaneous reaction between NH3 and HCl. kG a =
QG CG,A,i ln SL CG,A,o
(6)
A: solute (NH3 here): S: section area of Aquilair PlusTM L: Aquilair PlusTM length (m); CG,A,i , CG,A,o : ammonia gas concentration, inlet and outlet (mol m−3 ).
(m2 );
References [1] L.H. Hentz, B.M. Balchunas, Chemical and physical processes associated with mass transfer in odor control scrubbers, in: Proceedings of the Odors and VOC Emissions 2000, Cincinnati, United States, April 16–19, 2000, pp. 668–685. [2] L. Chen, J. Huang, C.L. Yang, Absorption of H2 S in NaOCl caustic aqueous solution, Environ. Prog. 20 (3) (2001) 175–181.
[3] J.R. Fair, A.F. Seibert, M. Behrens, P.P. Saraber, Z. Olujic, Structured packing performance-experimental evaluation of two predictive models, Ind. Eng. Chem. Res. 39 (2000) 1788–1796. [4] J.P. Ballaguet, C. Barrere-Tricca, C. Streicher, Improvements to tail gas treatment process, Petrol. Technol. Quat. (Summer) (2003) 09–115. [5] E. Brunazzi, A. Paglianti, S. Pintus, A capacitance probe and a new model to identify and predict the capacity of columns equipped with structured packings, Ind. Eng. Chem. Res. 40 (2001) 1205–1212. [6] P.Y. Le Strat, M. Cot, J.-P. Ballaguet, J.-L. Ambrosino, C. Streicher, J.-P. Cousin, New redox process successful in high-pressure gas streams, Oil Gas J. 11 (2001) 46–54. [7] F.A. Streiff, J.A. Rogers, Dont overlook static-mixer reactors, Chem. Eng. (June) (1994) 76–81. [8] O.N. Cavatorta, U. Bohm, A.M. Chiappori, Fluid-dynamic and masstransfer behavior of static mixers and regular packings, AIChE J. 45 (5) (1999) 938–948. [9] R. Pohorecki, W. Moniuk, Kinetics of reaction between carbon dioxide and hydroxyl ions in aqueous electrolyte solutions, Chem. Eng. Sci. 43 (1988) 1677–1684. [10] A. Couvert, M.-F. P´eculier, A. Laplanche, Pressure drop and mass transfer study in static mixers with gas continuous phase, CJChE 80 (4) (2002) 727–733. [11] P.A. Nawrocki, Z.P. Xu, K.T. Chuang, Mass transfer in structured corrugated packing, CJChE 69 (1991) 1336–1343. [12] M. Dragan, A. Friedl, M. Harasek, S. Dragan, I. Simiseanu, Measuring the effective mass transfer area of a Mellapack 750Y structured packing, in: Proceedings of the 14th International Congress of Chemical and Process Engineering, CHISA, August 27–31, 2000. [13] B. Benadda, Contribution a` l’´etude du transfert de mati`ere dans une colonne a` garnissage. Influence de la pression, PhD Thesis, University Claude Bernard of Lyon, France, 1994. [14] A. Lara-Marquez, T. Menguy, G. Wild, D´etermination des param`etres de transfert de mati`ere gaz–liquide dans un r´eacteur a` lit fixe co-courant ascendant de gaz et de liquide, Rec. Prog. Gen. Proc. 15 (16) (1991) 135–140. [15] K.C. Metha, M.M. Sharma, Mass transfer in spray columns, Brit. Chem. Eng. 15 (1970), 1440-1444 and 1556-1558. [16] J.-C. Charpentier, Mass transfer rates in gas liquid absorbers and reactors, in: T.B. Drew, G.R. Cokelet, J.W. Hoopes, T. Vermeulen (Eds.), Advanced Chemical Engineering, 11, Academic Press, New York, 1981, pp. 1133–1134. [17] L. Raynal, J.-P. Ballaguet, C. Barrere-Tricca, Determination of mass transfer characteristics of co-current two-phase flow within structured packing, Chem. Eng. Sci. 59 (2004) 5395–5402. [18] A. Couvert, C. Sanchez, I. Charron, A. Laplanche, C. Renner, Static mixers with a gas continuous phase, Chem. Eng. Sci. 61 (11) (2006) 3429–3434. [19] M. Roustan, Transferts gaz–liquide dans les proc´ed´es de traitement des eaux et des effluents gazeux, Tech. Doc. Lavoisier, Paris, 2003. [20] A.S. Jhaveri, M.M. Sharma, Effective interfacial area in a packed column, Chem. Eng. Sci. 23 (7) (1968) 669–676. [21] C. Sanchez, A. Couvert, A. Laplanche, C. Renner, New compact scrubber for odour removal in wastewater treatment plants, Wat. Sci. Technol. 54 (9) (2006) 45–52. [22] C. Sanchez, A. Couvert, A. Laplanche, C. Renner, Hydrodynamic and mass transfer in a new co-current two-phase flow gas–liquid contactor, Chem. Eng. J. 131 (1–3) (2007) 49–58.
Chemical Engineering and Processing 47 (2008) 1793–1798
Design of a new compact scrubber for deodorisation A. Couvert a,∗ , C. Sanchez b , A. Laplanche a , C. Renner c , J.-P. Levasseur d a
ENSCR, UMR CNRS 6226 - CIP, 263 Avenue du G´en´eral Leclerc, 35700 Rennes Cedex, France Burgeap, Pˆole air et Sant´e publique, Parc de la Bastide Blanche, Bˆat. 6, 13127 Vitrolles, France c Anjou Recherche – Veolia Environnement, Chemin de la digue, BP 76, 78603 Maisons Laffitte, France d Direction Technique, Veolia Eau, 1 place Montgolfier, 94417 Saint-Maurice Cedex, France b
Received 28 June 2007; received in revised form 31 August 2007; accepted 1 October 2007 Available online 12 October 2007
Abstract A new compact scrubber was designed to remove odorous compounds from WWTPs (named Aquilair PlusTM , for Compact Reticulated Contactor). Hydrodynamics and mass transfer parameters (P, a, kL a and kG a) were quantified versus superficial gas and liquid velocities (USG and USL ). This study showed that all these parameters increase with USG and USL , except for kG a which is almost independent of USL . Thanks to the lab-study, a half-size industrial pilot plant was designed. First results are encouraging. Finally, an economic study showed that Aquilair PlusTM allows a gain of 15–40% compared to traditional packed towers. © 2007 Elsevier B.V. All rights reserved. Keywords: Absorption; Scrubber; Process intensification; Hydrodynamics; Mass transfer; Design
1. Introduction Public concern over odours from wastewater treatment plants (WWTPs) is increasing because of the development around existing, or in construction works. Species identified as responsible for odours in WWTPs are mainly reduced sulphur and nitrogen compounds. Their removal is most of time realised by chemical scrubbing in a two or three stage packed tower (PT) system. This approach has been showing good results for many years [1,2], but has also revealed high investment and working costs (high residence time implying the construction of high and wide towers, important reactive consumption mainly due to competitive CO2 absorption). In the last decades, structured PT [3,4] have gained increased attention because of their better performances in terms of gas–liquid contact and distribution, coupled with their higher bed void fraction resulting in lower pressure drops [5]. Filled with random or structured packing, most of the studies on PT found in literature focus on counter-current flow applications; nevertheless, when absorbed species react in the liquid phase according to a fast irreversible reaction, and especially when high fluid velocities are implemented, the cocurrent flow configuration is preferable. Moreover, co-current ∗
Corresponding author. Tel.: +33 2 23 23 80 48; fax: +33 2 23 23 81 20. E-mail address: [email protected] (A. Couvert).
0255-2701/$ – see front matter © 2007 Elsevier B.V. All rights reserved. doi:10.1016/j.cep.2007.10.006
working avoids flooding, allowing higher flexibility [6]. Then, some researchers paid attention on static mixers (SM) because of their excellent gas–liquid contact due to a high turbulence rate inside [7,8]. Different types of static mixers were designed (Sulzer, Lightnin, Ross, Kenics, etc.) providing good results in terms of mixing (their first application), but also in terms of mass transfer, which is interesting for scrubbing. Unfortunately, these gas–liquid contactors generate too high-pressure drop to be industrially viable. Consequently, the objective of this study was to design a compact apparatus, able to generate high interfacial areas and good mass transfer coefficients (high efficiency), without driving to a high-energy consumption. In order to conclude on the best operating conditions and design parameters for scale-up, pressure drop, interfacial area and volumetric mass transfer coefficients were studied in a laboratory-scale pilot unit. This step permitted to design a half-sized industrial pilot plant. 2. Experimental The lab-reactor (named Aquilair PlusTM ) consists of a transparent PVC pipe of 0.32 m length and 0.025 m inner diameter (volume of 1.53 × 10−4 m3 ), through which flow patterns as well as the type of dispersion can be observed (Fig. 1). The contactor is a new wire mesh packing structure which geometric characteristics are given in Table 1. Meshes are 12.5 mm side, made
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Fig. 1. Scheme of the experimental set-up.
of stainless steel. Then, the structure is bent (to avoid too much compactness, and so, high pressure drop) and rolled up so that it can be inserted in an empty tube of 0.025 m diameter. The packing geometric specific area, which is equal to 176 m2 m−3 of reactor, takes the surface offered by the packing itself and by the column wall into account; its packing factor calculation is higher than the ones of conventional packing, which supposes that a greater interfacial area will be offered to gas–liquid contact; the void fraction is also very high (97.5%) in order to minimize pressure drop. The volume gas flow rate QG , controlled with a rotameter, ranged from 10 to 50 m3 h−1 (superficial gas velocity USG from 5.6 to 28 m s−1 ). The liquid was injected into the reactor at liquid flow rates QL ranging between 0.0275 and 0.0975 m3 h−1 (superficial liquid velocity USL between 0.016 and 0.055 m s−1 ). The fluids circulated in vertical descendant (VD) co-current flow inside the pipe. pH and temperature were continuously measured with specific probes. The pressure drop P within the reactor was measured with two manometers: the first one was filled with water (for the low pressure drop values) and the second one was filled with mercury (for the high-pressure drop values). They were connected just before and after Aquilair PlusTM and the pressure drop was measured for the different liquid and gas flow rates mentioned previously. The interfacial area a was determined by transferring CO2 (CG,CO2 = CG,A = 4000 ppmv) into an aqueous sodium hydroxide solution (CL,OH − = 1N). Details are given in the Annex. The CO2 gas concentrations (inlet and outlet) were measured with an IR analyser type Beryl 100 from COSMA and in the liquid phase by the measurement of the alkalinity and total alkalinity (AFNOR NF T 90-036, July 1977). The liquid-side volumetric mass transfer coefficient kL a was measured thanks to the implementation of a physical transfer
of gaseous 2-butanone in an aqueous phase (water, pH equal to 7). 2-Butanone, chosen because of its small solubility (Henry’s constant = 6.7 × 10−2 atm L mol−1 at 25 ◦ C), was injected as a liquid in the gas flow with a syringe distributor, vaporised in it in order to reach a concentration of about 50 mg m−3 in the gas phase (see Eq. (5) in Appendix B). 2-Butanone concentrations in the gas phase (inlet and outlet) were indirectly measured after trapping it in a 2,4-dinitrophenylhydrazine (2,4-DNPH) solution. Indeed, the trapping yields for 2-butanone exceed 99%. This method is based on the reaction between 2-butanone and 2,4-DNPH, giving an hydrazone which is then analysed by HPLC (C18, mobile phase composed of 80% methanol and 20% water, UV detection for a wavelength of 365 nm). The gas-side volumetric mass transfer coefficient kG a was obtained by implementing an irreversible and instantaneous reaction between NH3 and HCl (see Eq. (6) in Appendix B). The condition that must be verified to implement this method is kG EkL /H. This was confirmed once kG a was measured (EkL /H varied from 500 to 1080, kG values range between 0.05 and 0.28 m s−1 ). Concerning the half-size pilot plant, it was of 0.3 m diameter and 0.6 m length, with slightly different internal packing characteristics (void fraction, geometric surface) from the labpilot ones. Liquid and gas were also distributed by the top of the contactor, in co-current flow. Gas velocities ranged between 5 and 20 m s−1 for a chosen liquid flow rate of 8 m3 h−1 , in order to reproduce the lab operating conditions. P was measured with pressure sensors type Cerabar S PMC71 located before and after the contactor, and the gas velocity was obtained thanks to a Pitot probe from Endress Hauser. The H2 S gas concentration was measured with Odalog specific probes (App-Tek type I), which range from 0 to 200 ppm.
Table 1 Geometric characteristics of Aquilair PlusTM
3. Results and discussion
Height (m) Internal diameter (m) Void fraction (%) Geometric surface ag (m2 m−3 ) Packing factor (m2 m−3 ) Volume (m3 )
0.32 0.025 97.5 176 6667 1.53 × 10−4
3.1. Lab-plant study Fig. 2 shows the pressure drop per reactor length unit versus superficial gas and liquid velocities. As expected, P/L increases with both operating parameters. Moreover, it can be
A. Couvert et al. / Chemical Engineering and Processing 47 (2008) 1793–1798
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Fig. 2. Evolution of pressure drop vs. superficial gas and liquid velocities in the VD configuration.
Fig. 3. Evolution of the interfacial area vs. the superficial gas and liquid velocities in the VD configuration of Aquilair PlusTM .
noticed that P/L in our reactor is three times lower than the one in the Lightnin SM, and far lower than those generated by SMV4 Sulzer SM, which can reach values up to 50,000 Pa m−1 for USG = 15 m s−1 [10]. This seems logical if one compares void fractions in the different gas–liquid contactors (78% for the Sulzer SM, 87% for the Lightnin SM, 97.5% for Aquilair PlusTM ). The more the void fraction, the less the resistance opposed to the fluid flow. Otherwise, the interfacial area a measured ranged from 600 to 2800 m2 m−3 (Fig. 3). These values are twice higher than the ones measured in the Lightnin SM and the empty tube, and far higher than the ones obtained in packed towers (around 100–300 m2 m−3 ). Moreover, the interfacial area measured in towers filled with structured packing and operating in counter-current flows can reach values ranging from 250 up to 750 m2 m−3 [11,12]. The increasing of this parameter with the superficial liquid velocity was foreseeable since when USL increases, the number of liquid drops or trickles in contact with the gas phase rises. Several authors observed the same phenomenon in different gas–liquid contactors: Benadda [13] and Lara-Marquez [14] in packed towers, Metha et al. [15] in spraying towers. In the same time, when USG increases, turbulence is more important, so that liquid trickles are broken into small drops, driving to a higher gas–liquid contact area.
Concerning mass transfer parameters, the VD configuration led to promising values of kL a coefficients, up to 0.20 s−1 (Fig. 4). As a comparison, kL a measured in the empty tube does not overcome 0.05 s−1 for USG < 15 m s−1 and USL < 0.055 m s−1 , kL a found in PT varies between 2.5 × 10−3 and 0.035 s−1 [16], whereas the values found in structured PT are expected to be higher depending on the geometric characteristics of the packing elements. As an example, Raynal et al. [17] found kL a values ranging from 0.15 up to 0.25 s−1 with their packing made of smooth stainless sheets, but at very high L /G
Fig. 4. Evolution of the liquid-side volumetric mass transfer coefficient vs. the superficial gas and liquid velocities in the VD configuration of Aquilair PlusTM .
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Fig. 5. Evolution of the gas-side volumetric mass transfer coefficient vs. the superficial gas and liquid velocities in the VD configuration of Aquilair PlusTM .
ratios (ours vary from 0.5 to 5). Otherwise, kL a values measured in Statiflo and Lightnin SM are lower than the ones obtained in Aquilair PlusTM as soon as USG ≥ 10 m s−1 [18]. Like the other parameters, kL a is an increasing function of the superficial gas and liquid velocities. Indeed, on the one hand, a increases with USG and USL , and on the other hand, the higher these two parameters, the higher turbulence in the contactor, which lets suppose higher values of kL . Finally, kG a was found to vary between 40 and 530 s−1 (Fig. 5). These values are similar to those found in other classical reactors like packed towers or venturis [19] and show that the mass transfer resistance is located in the liquid phase. kG a increases with USG , but not with USL . In 1968, Jhaveri et al. [20] have noticed the same phenomenon. 3.2. Extrapolation parameters Based on the study of hydrodynamic and mass transfer parameters, but also of efficiency performances revealed by the lab-contactor [21], a half-sized industrial pilot plant was designed in order to treat a polluted gas flow rate of 4000 m3 h−1 (Fig. 6). Here, it is important to specify that to treat the same gas flow rate in a packed tower, the volume of the contactor would be far higher; indeed, a pilot packed tower was implemented in the same area, in order to treat the same gas flow rate, and its volume was 25 times higher than the Aquilair PlusTM one. The first step was to choose a superficial gas velocity according to an economically acceptable pressure drop. This pressure drop was reached for USG = 15 m s−1 in the lab plant. At this superficial gas velocity, mass transfer parameters were satisfactory, and the sulphur compounds well removed according to the objectives defined. Taking these considerations into account, and basing the extrapolation on a contact time Tc of 0.04 s, a larger and longer reactor (0.3 m diameter, 0.6 m length) was implemented in the WWTP. Its void fraction was slightly higher (98.2% instead of 97.5%), and its geometric specific area lower (ag = 120 m2 m−3 instead of 176 m2 m−3 ) than the ones of the lab-plant; moreover, wall effects are lower in the industrial plant because of the higher diameter compared to the packing (45% in the lab-plant against 7% in the pilot plant), which lets suppose that the pressure drop would be lower in the industrial plant than in the lab-plant if all the other geometric and operating characteristics were similar.
Fig. 6. Photo of the half-size industrial plant.
First, the pressure drop was measured and compared to the one predicted by the correlation stemmed from the lab experimental data. Fig. 7 shows the variations of the pressure drop with USG for QL = 8000 L h−1 (USL = 0.03 m s−1 ). It can be observed that the experimental values are slightly lower than the ones measured in the lab-plant (see correlation of Sanchez et al. [22]). This result can be related to the higher void fraction and the lower wall effects in the pilot plant. No measurement of the mass transfer parameters (a or kL a) was done on this pilot plant, but observing the gas and liquid flows inside the contactor, it could be suspected that the dispersed regime, favourable to good mass transfer performances, was not completely reached. Then, the performances of the pilot plant in terms of sulphur hydrogen removal were characterised. As expected, the efficiency was not as good as the one found in the lab-plant (76% instead of 95% for L /G ≈ 3). This could be attributed to the mass transfer performances of the industrial plant, lower than the ones measured in
Fig. 7. Evolution of the pressure drop in the half-size industrial plant vs. the superficial velocity (QL = 8000 L h−1 ).
A. Couvert et al. / Chemical Engineering and Processing 47 (2008) 1793–1798
the lab-plant, certainly because of its higher void fraction. This could be verified by realising one test on a configuration more similar to the lab one (void fraction of 97.7%, ag = 160 m2 m−3 ). Indeed, the H2 S removal for USG = 12 m s−1 was found equal to the one measured in the lab-plant (about 90%). Unfortunately, the pressure drop was higher than the one obtained in the lab-plant. Some improvements are being realising. 3.3. Economical evaluation A study consisting in replacing a classical system in a WWTP (two or three packed towers designed to treat a gas flow rate of 100,000 m3 h−1 ) by five Aquilair PlusTM units was led by the technical department of Veolia Water. The difference between both systems comes from the fact that the Aquilair PlusTM units are located near each source of odorous compounds (four to treat sulphur compounds of the pre-treatment, biological treatment, sludge treatment and relieving post; one more to treat ammonia of the sludge treatment post), whereas the two or three packed towers are situated in a building dedicated to deodorisation. Therefore, in the classical system, the odorous gas has to be collected and carried to these premises, which implicates more ventilators and ventilation pipes. This is why the economical evaluation took the costs related to civil engineering, instrumentation, reactive and electricity consumption into account, but also ventilation expenses. This estimation led us to conclude that the gas–liquid contactor developed in this study drives to savings in terms of investment and working (15–40%), without consuming more energy, and decreasing ground hold in the same time. 4. Conclusion Hydrodynamic and mass transfer parameters were quantified in a new compact scrubber, Aquilair PlusTM . This gas–liquid contactor proved to be economically interesting in term of pressure drop compared with conventional structured packing reactors, but also in terms of performances (high-interfacial areas and liquid-side volumetric mass transfer coefficients). Moreover, Aquilair PlusTM presents the plus to be compact, therefore to limit ground hold, and costs related to it. Thanks to the lab study, a half-size industrial plant was designed. First results concerning the pressure drop it generates are promising. Efficiency experiments will confirm the potentiality of Aquilair PlusTM as a scrubber for gas containing sulphur compound. Finally, a comparison between a classical system and the Aquilair PlusTM solution proved that our reactor is economically more advantageous than two or three conventional packed towers. Appendix A. Nomenclature
A a C D
solute (CO2 , 2-butanone or NH3 ) volumetric interfacial area (m2 m−3 ) concentration (mol m−3 ) diffusivity coefficient (m2 s−1 )
E G H k kL , kG L L n N Q S Us v V
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enhancement factor mass gas flow rate (kg s−1 ) Henry’s constant (Pa m3 mol−1 ) kinetic constant ((mol m−3 )1−n s−1 ) film mass transfer coefficient (m s−1 ) contactor length (m) mass liquid flow rate (kg s−1 ) reaction global order molar mass flux transferred (mol s−1 ) volume flow rate (m3 s−1 ) section area of Aquilair PlusTM (m2 ) superficial velocity (m s−1 ) reaction rate (mol m−3 s−1 ) reactor volume (m3 )
Symbols P pressure drop (Pa) μ dynamic viscosity (Pl) Index G, L i, o
gas, liquid inlet, outlet
Appendix B. Annex B.1. Interfacial area The interfacial area a was determined by transferring CO2 into an aqueous sodium hydroxide solution. It was deduced from Eq. (3) considering the reaction between CO2 and NaOH as rapid (Ha > 5, see Eq. (1)), irreversible (v = kOH− CL,A CL,OH− with kOH− = 6.4 m3 mol−1 s−1 ) and of pseudo-first order (CL,OH− CL,A , so that CL,OH− ≈constant): Ha2 =
DA kOH− CL,OH−
(1)
kL2
Then, the molar mass flux of CO2 transferred can be written as the following equation: ∗ −C N = kL aVE (CL,A L,A )
(2)
In such conditions, CO2 concentrations in the liquid are considered equal to zero (all the CO2 absorbed is consumed by the rapid reaction), and Eq. (2) gives the following equation: a=
V
N
∗ ∗ ∗ DA kOH− CL,OH− (CL,A,i /CL,A,o )/ ln((CL,A,i
∗ − CL,A,o ))
(3) A: solute (CO2 here), N: molar mass flux of CO2 trans∗ : CO ferred (mol s−1 ), V: contactor volume (m3 ), CL,A 2 liquid concentration at the gas–liquid interface (mol m−3 ), kOH− = kinetic constant of the reaction between CO2 and NaOH (1 N), DA = 1.62 × 10−9 m2 s−1 : CO2 diffusivity in sodium hydroxide calculated from the CO2 diffusivity in water DA/water μwater = DA/solution μsolution [9]
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In these conditions, we assume that gas-side resistance is negligible (HA kG kL with HA = 2630 Pa m3 mol−1 Henry’s constant of CO2 ). Then interfacial concentration can be calcu∗ ) and molar mass transfer lated with Henry’s law (PA = HA CL,A flux from the variation of the CO2 concentration in the gas phase and in the liquid phase (Eq. (4)). N = QL (CL,A,o − CL,A,i ) = QG (CG,A,i − CG,A,o )
(4)
B.2. Liquid-side volumetric mass transfer coefficient The liquid-side volumetric mass transfer coefficient kL a was measured thanks to the implementation of a physical transfer of gaseous 2-butanone in an aqueous phase. In this case, the mass transfer resistance is located in the liquid film, and the 2-butanone liquid concentration at the inlet is zero; then kL a is calculated from the following equation: kL a =
N ∗ E.V.(CL,A,i
∗ − (CL,A,o
∗ ∗ − CL,A,o )/ ln(CL,A,i /(CL,A,o − CL,A,o ))
(5) A: solute (2-butanone here); E: enhancement factor = 1 when there is no reaction in the liquid phase; N: molar mass flux of 2-butanone transferred (mol s−1 ); CL,A,o : 2-butanone liquid con∗ centration in the liquid phase at the outlet (mol m−3 ); CL,A,i , ∗ CL,A,o : 2-butanone liquid concentration at the gas–liquid interface (inlet and outlet) (mol m−3 ) B.3. Gas-side volumetric mass transfer coefficient The gas-side volumetric mass transfer coefficient kG a was obtained by implementing an irreversible and instantaneous reaction between NH3 and HCl. kG a =
QG CG,A,i ln SL CG,A,o
(6)
A: solute (NH3 here): S: section area of Aquilair PlusTM L: Aquilair PlusTM length (m); CG,A,i , CG,A,o : ammonia gas concentration, inlet and outlet (mol m−3 ).
(m2 );
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