Pilot Scale Gasification of Spent Cooking Liquor from Sodium Sulfite

Article pubs.acs.org/EF

Pilot Scale Gasification of Spent Cooking Liquor from Sodium Sulfite Based Delignification Erik Furusjö,*,†,‡ Ragnar Stare,‡ Ingvar Landal̈ v,†,‡ and Patrik Löwnertz‡ †

Division of Energy Sciences, Luleå University of Technology, SE-971 87 Luleå, Sweden Chemrec AB, Drottning Kristinas Väg 61, SE-114 28 Stockholm, Sweden

‡

ABSTRACT: This paper describes a pilot scale high pressure entrained flow gasification experiment with spent cooking liquor from a sodium sulfite based delignification process in the DP-1 black liquor gasifier in Piteå, Sweden. Approximately 92 tons of sulfite thick liquor were gasified during 100 h of operation without any operational problems despite the new feedstock. The syngas quality was found to be good for all operating points with the CH4 content below 0.3% and H2/CO ratio between 1.03 and 1.15. The experiment shows that the process capacity is limited by green liquor quality parameters primarily dependent on the presence of small amounts of unconverted carbon. The pilot plant capacity was found to be somewhat lower than for Kraft black liquor on mass basis but higher when measured as thermal load, due to the higher heating value of sulfite thick liquor. Mass and energy balances were made difficult by the unavailability of measured green liquor and syngas flow rates, which lead to the necessity of using alternative approaches for the estimation of these flows. Using these estimates, overall mass and energy balances were closed to within 5% for all operating points except one, and the process cold gas efficiency was 60−68% on sulfurfree lower heating value basis. Carbon balances indicate that 95−97% of feedstock carbon leaves with the syngas, mainly as CO and CO2 with the remainder being mostly green liquor carbonate. More than 95% of the feedstock sodium is found in green liquor, while 3−5% ends up in the gas condensate purge stream. The sulfur balance does not close as well as other elements but indicates that 70−73% of the feedstock sulfur ends up in the syngas as H2S and COS with the remainder being present in green liquor as dissolved sulfide salts.

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INTRODUCTION

problematic for the processes in the pulp mill that regenerates the pulping chemicals. Domsjö Fabriker is a sodium sulfite based biorefinery with three main products: specialty cellulose, lignin (as lignosulfonate), and ethanol. The biorefinery utilizes a two stage batch pulping process for the production of the main product: specialty cellulose. The liquor from the pulping process is used to generate the other two products; ethanol is produced by fermentation of hexose sugar in the liquor, and a part of the fermented liquor is dried and sold as lignin. However, the major part of the liquor from the pulping process is concentrated in evaporators to produce so-called sulfite thick liquor (STL), which is subsequently combusted in two RBs on site for chemical and energy recovery. This paper presents a study, which was carried out to investigate the possibility of replacing the RBs with gasification technology in order to add new biorefinery products and increase the energy efficiency. It is generally agreed in the pulp and paper industry that combustion of spent pulping liquors from sodium sulfite based pulping is more difficult than those from the sulfide based Kraft pulping processes, e.g., due to corrosion problems,10 and that the capacity is lower when STL is fired in the same equipment. Gasification of sulfite liquors is not a new approach. There was considerable interest in the development of sulfite-based pulping in Sweden around 1960. This led to a demand for an

Gasification of spent cooking liquors from the pulp and paper industry is a promising technology for production of renewable fuels and chemicals1 that has been attempted with a range of technology variations.2,3 The dominant pulping process is the Kraft process, and the spent pulping liquor from this process, black liquor (BL), has received most attention. When BL gasification (BLG) technology is integrated with a pulp mill, the recovery boiler (RB) that is normally used to recover energy and pulping chemicals is replaced with a gasifier. The BL energy is utilized to produce syngas, while the inorganic pulping chemicals are returned to the mill in an aqueous solution called green liquor (GL) similarly to how pulping chemicals are recovered in an RB based system. A 3 MWth pilot plant for pressurized oxygen-blown entrained-flow BLG is located in Piteå, Sweden and has been operated by Chemrec 2005−2012.4 Since January 2013 the plant is owned and operated by Luleå University of Technology, having a license agreement with the technology owner Chemrec. The plant has currently accumulated close to 25 000 operating hours on Kraft BL5 and also been used in a number of research studies to increase the understanding of various aspects of the technology.6−9 No complete mass and energy balances have been published, however. Practical experience from pilot plant operations shows that GL quality is normally the limiting factor for feasible operating conditions when Kraft BL is used as feedstock. Insufficient atomization, too low temperature, and/or too short residence time leads to unconverted carbon in the GL, which is highly © 2014 American Chemical Society

Received: August 19, 2014 Revised: November 11, 2014 Published: November 17, 2014 7517

dx.doi.org/10.1021/ef501753h | Energy Fuels 2014, 28, 7517−7526

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Figure 1. Schematic of pilot scale gasifier and gas cooler. Indicated streams are used in mass and energy balances.

temperature of 1000−1100 °C. A sketch of the geometry of the 0.6 m3 reactor has been published previously.6 The products from the gasification reactor are quenched in two stages using water sprays. The inorganic chemicals are separated from the gas by gravity before the syngas is indirectly cooled to approximately 40 °C in a counter current condenser. Separated inorganics, mainly sodium carbonate and sodium sulfide, are dissolved in water to form GL. Syngas is sent to downstream processes for production of bio-dimethyl ether (bioDME), but those processes are not further discussed here since they are out of the scope of this work. Some modifications in support systems were made, which were motivated by safety concerns related mainly to the low pH of the STL. These are not discussed further in this paper, since they do not influence process performance and characteristics. Flows of STL and oxygen are measured by Coriolis type mass flow meters. Flows of nitrogen, cooling water to gas cooler, water to GL dissolver, water to quench spray, and flashcooled condensate purge were measured through orifice flow meters. The flash-cooled condensate purge flow is recalculated to a hot pressurized condensate purge flow using temperature measurements before and after the flash vessel as well as balances over the flash vessel. As discussed below, flow measurements for GL and syngas are not reliable and were not used. Pressures and temperatures are measured in a large number of positions in the process, covering all important points for the major process and utility flows. Temperature measurements are available for all flows indicated in Figure 1. Pressure measurements are available for all flows indicated in Figure 1 except cooling air and cooling water. Temperatures in the gasification reactor are measured by seven thermocouples (TCs) protected by a ceramic encapsulation located in the cylindrical section of the reactor close to the top (3 TCs), in the middle (3 TCs), and close to the bottom conical section (1 TC) of the reactor.

efficient recovery process for energy and chemicals; gasification was one of the options investigated. Tests carried out showed that a relatively high gasification temperature (>900 °C) was required to obtain good conversion of carbon, but also low temperature approaches were explored. The gasification pathway was subsequently abandoned due to material problems and poor performance; see Whitty2 and Furusjö et al.11 for more details. Little effort has been devoted to the recovery of energy and chemicals from sulfite pulping liquors in recent years. However, recent interest in biorefinery development and in particular the possibility to produce biofuels or green chemicals from lignocellulosic feedstocks12 has brought renewed interest in sulfite based delignification processes for further biochemical processing13 as well as thermochemical processing as in the present work. In 2009 a first pilot scale experiment concerning pressurized oxygen-blown high-temperature entrained-flow STL gasification was executed.11 The experiment showed that the process was feasible but practical problems with liquor preheating limited the attainable load to about 50% of design capacity. Furthermore, the relatively short duration of the experiment means that steady state was not achieved in the GL for most operating points. The objectives of the present study were primarily to determine the pilot plant capacity for STL in comparison to the Kraft BL capacity, to quantify and verify process performance during longer operating periods, and to present mass and energy balances for the process.

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EXPERIMENTAL SECTION The pilot gasifier test described in this report was performed in the DP-1 BLG pilot plant in Piteå, Sweden. The plant setup is not discussed in detail in this paper since it has been described previously.4,6−9 Briefly, the entrained flow gasification process shown schematically in Figure 1 is fed with spent pulping liquor and oxygen through a nozzle at the top of a ceramic-lined reactor operating at up to 30 bar pressure and a global 7518


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Discussion sections. A total of 92 t, or approximately 70 m3, of STL was gasified during the 100 h experiment duration (including start-up and transitions). Operating parameters not shown in Table 2 were kept at a level equal to what is normally used for Kraft BL, except for the primary quench spray flow, which was higher than normal and consisted of cold water instead of gas cooler condensate. This has an effect on GL CO2 absorption as discussed below. Gas cooler condensate was used in all other positions in the quench. Important parameters that were not varied were oxygen preheating at 50 °C, nitrogen to the central pipe of the burner at ∼20 Nm3/h, cool gas temperature ∼35 °C, and total flow of gas condensate to the quench (in different positions) at 6800 L/h. No water is added to the gasification reactor, except the water present in the STL as indicated by the STL dry solids content listed in Table 2. Samples of syngas, gas cooler condensate, and GL were taken shortly before the end of each operating point. Because of the short gas residence time in the 0.6 m3 gasification reactor, which is approximately 10 s for the conditions of Table 2, gas samples are representative of the time at which they were taken. The hydraulic residence time for GL is approximately 1.9 h, meaning that the GL volume is replaced 2.5−4.5 times during the 5−8 h duration of operating points 2−6 and 8−9. Gas condensate is replaced 10−16 times during each of these operating points. These values are considered adequate for the purposes of this work, although a longer duration of each point would make it possible to detect weaker effects of process changes. GL samples were subject to analysis by MoRe Research, Ö rnsköldsvik, Sweden. Carbonate and hydrogen carbonate were determined by acid titration. Sulfide was determined by silver nitrate titration. Total sulfur was determined by wet oxidation and ion chromatography. The solids concentration was determined by filtering and weighing. The solids volume was determined at the pilot plant using an Imhoff cone. Syngas was sampled from a position downstream the gas cooler in Tedlar bags that had been cleaned by purging with nitrogen several times. The gas sample was then analyzed with respect to CH4, CO, CO2, H2, N2, H2S, COS, and O2/Ar using a Varian CP-3800 gas chromatograph at Energy Technology Center, Piteå, Sweden. The water content was not measured but was estimated assuming saturation at sampling temperature and pressure. Gas cooler condensate was analyzed for Na by the Smurfit Kappa Kraftliner Laboratory, Piteå, Sweden, using AES (atomic emission spectroscopy). Two selected samples were subject to

Process data used are averages of data from a period of 30 min ending 10 min before the end of the operating point considered. This approach is used to avoid undesirable effects from data compression in the process database in association with operating point changes. STL was taken from the evaporation plant of the biorefinery, which is located after hexose fermentation, and transported daily in insulated trucks to the pilot plant (330 km). A lower dry solids (DS) content than normal (70% DS) was used to avoid chemical reactions in the liquor during transport; 58% DS content was used during the majority of the experiment, but the last truck load contained 63% DS STL. Table 1 shows liquor Table 1. STL Analysis and Typical Kraft BL Composition To Illustrate Differences

HHV C H S O Na K Cl N DS

MJ/kg DS kg/kg DS kg/kg DS kg/kg DS kg/kg DS kg/kg DS kg/kg DS kg/kg DS kg/kg DS kg DS/kg

low DS

high DS

typical Kraft BL

uncertainty (absolute)b

18.3 42.5% 4.2% 9.6% 33.99%a 8.7% 0.17% 0.01% 0.83% 58.4%

18.4 42.5% 4.1% 9.2% 34.84%a 8.3% 0.13% 0.01% 0.92% 63.1%

13.4 33.9% 3.5% 5.0% 36.2% 19.0% 2.3% 0.1% 0.1% 73%

±0.2 ±0.9% ±0.25% ±1.1% ±0.7% ±0.1% ±0.001% ±0.1% ±1.2%

a

By difference. bAccording to laboratory documentation and recalculated from relative where applicable.

composition and properties (SP Technical Research Institute of Sweden, Borås, Sweden). An average of the two liquor compositions for all parameters except DS was used in the calculations since the differences between liquors are not significant and no changes which would be expected to have major impact on liquor composition were made in the pulping process. Each of the trucks contained up to 25 m3 of liquor. The liquor was unloaded to a buffer tank upon delivery. Care was taken to clean the buffer tank before each new load of liquor to avoid “old” liquor, which could possibly have started to decompose/react, from being fed to the pilot plant. Table 2 gives an overview of the pilot experiment’s operating points. The values of individual parameters for these operating points were partly motivated by the analysis results obtained during the experiment as described in the Results and Table 2. Overview of Operating Points op. point

duration (h)

STL (wet) (kg/h)

pressure reactor (bar g)

O2 (kg/h)

atom.a N2 (Nm3/h)

CH4 cold gasb (vol %)

liquor feed (°C)

liquor DS (%)

1 2 3 4 5 6 7 8 9

2.7 8.1 8.8 7.9 7.7 5.2 21.0 6.0 5.0

649 833 1018 1017 1018 1019 949 939 1043

28.3 27.4 29.2 27.2 27.1 27.1 27.1 27.1 27.2

219 273 332 336 337 342 315 314 342

50 0 0 16 0 0 27 27 6

0.2% 0.18−0.23% 0.30−0.35% 0.20−0.25% 0.15−0.20% 0.09−0.15% 0.20−0.25% 0.25−0.30% 0.20−0.25%

129 129 129 129 129 129 129 129 135

58.4% 58.4% 58.4% 58.4% 58.4% 58.4% 58.4% 62.6%c 62.9%c

Atomization support nitrogen mixed with oxygen before the burner. bIncluded as an indication of reactor temperature; oxygen flow used to control temperature, see text. c63.1% dry solids not fully reached due to presence of low DS liquor in the day tank after point 7. a

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Table 3. Results from Chemical Analysis of GL Samples op. point

TTAa (mol/L)

[HCO3−] (mol/L)

[CO32−] (mol/L)

[HS−] (mol/L)

tot. S (mol/L)

S red.b (%)

S/(Na+K)2 (mol/mol)

TOCc (mg/L)

1 2 3 4 5 6 7 8 9

0.82 1.66 2.91 2.73 3.13 2.97 2.51 2.40 2.70

0.28 0.51 0.59 0.45 0.50 0.45 0.41 0.32 0.26

0.17 0.38 0.84 0.84 0.95 0.91 0.76 0.75 0.89

0.20 0.40 0.63 0.61 0.72 0.70 0.59 0.58 0.65

0.20 0.40 0.66 0.64 0.71 0.71 0.52 0.55 0.58

100 100 96 96 102 98 113 105 113

0.48 0.48 0.43 0.45 0.46 0.47 0.47 0.48 0.48

n.a. 47 99 71 116 20 22 29 25

Total titratable alkali calculated as [HCO3−] + 2[CO32−] + [HS−], approximately equal to [Na+] + [K+]. bSulfur reduction efficiency calculated as hydrogen sulfide concentration divided by total sulfur concentration. cTotal organic carbon in filtered GL, i.e., excluding sludge. a

TTA. HCO3− concentrations of about 10−20% of TTA (for samples that do not have very low TTA) are higher than for normal operation on BL. This is due to CO2 absorption caused by the high primary quench spray flow7 used in this experiment. The process was operated with high primary quench spray flow since CO2 absorption, leading to the presence of hydrogen carbonate in GL, is advantageous to sulfite chemical recovery based on either the Stora or Tampella process. Hence, it can be noted that the GL composition in Table 3 is not representative of Kraft BLG where the process is operated to minimize CO2 absorption, since it leads to a higher causticizing requirement in recovery of Kraft pulping chemicals and is thus a disadvantage in that case. High sulfur reduction efficiency is important for the reuse of pulping chemicals in the mill. Efficiencies given in Table 3 are sometimes over 100% due to analysis errors, but the overall conclusion is that reduction efficiency does not deviate from 100% within measurement accuracy, which is in contrast to recovery boilers that normally give around 95% for Kraft BL14 and lower values for STL. Data in Table 4 show that the solids volume and concentration are well-correlated except for operating point 3 that has GL of poorest quality (based on visual appearance). This indicates that a more voluminous sludge is produced under such conditions. Syngas. The composition of cold syngas is shown in Table 5. The main gas components are H2, CO, and CO2 with smaller amounts of H2S, COS, and CH4. Trace components such as benzene, HCl, and NH3 were not measured in this campaign but have been measured previously for BLG.8 Gas Cooler Condensate. The gas cooler condensate is water that is condensed from the saturated raw syngas in the gas cooler. It contains dissolved gases and inorganic salts that are washed from the syngas. The amount of alkali metals is a good measure of the amount of alkali that is carried over from the quench to the gas cooler. Table 6 shows the Na contents in gas condensate samples. Table 7 presents an analysis of selected metals in two selected gas condensate samples. As expected, Na is the dominant element. C was not analyzed, but HCO3− and/ or CO32− are the expected counterions with smaller amounts of hydrogen sulfide being present as well.

elemental analysis by ICP-MS by ALS Scandinavia, Luleå, Sweden.

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RESULTS Operating Conditions. In the previous experiment,11 reactor load was limited by fouling in the liquor preheater, but the use of a redesigned heater in this experiment eliminated those problems. As noted above, the operating conditions of a BLG plant are normally limited by GL quality. A major purpose of this work was to determine the pilot plant capacity on STL. Hence, operating conditions (Table 2) were adjusted in order to try to obtain conditions that would give an acceptable GL quality using as high load as possible. Due to the long residence time for GL in the process, at least 5−8 h duration for an operating point is required in order to be able to quantify the effects on GL quality, as discussed above. Green Liquor. GL samples were subject to chemical analysis (Table 3) and visual inspection and solids quantification (Table 4). Since the process was intentionally Table 4. Appearance and Solid Content of GL Samples op. point

solids vol. (ml/L)

solids conc. (g/L)

1 2 3

10 25 80

n.a. 1.1 2.3

4 5 6

40 55 20

1.6 2.6 0.9

7 8 9

30 45 45

2.1 2.0 2.3

appearance transparent transparent; oil film dark transparent; floating sludge transparent; oil film transparent; minor oil film nonsettling particles; minor oil film transparent; minor oil film transparent; very minor oil film transparent; minor oil film

operated on the border of acceptable GL quality, some samples inevitably have unacceptable quality, e.g., sample 3, cf. Table 4. See the Discussion section for more details about the evaluation of GL quality. With the exception of the first two operating points, total titratable alkali (TTA) is 2.4−3.1 mol/L, which is somewhat lower than what is expected of a commercial plant. The reason for this is the difficulty in controlling TTA during the short experiment duration under varying load. This is not considered to make analysis of GL composition etc. less representative. The adjustment to other TTA levels is straightforward. Sulfidity in GL, measured as S/(Na + K)2 ratio, is 0.45−0.48. This corresponds to a HS− concentration that is 22−24% of

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DISCUSSION Gasifier Operating Conditions and Syngas Properties. There were no operational problems related to the process during the experiment. The process was operated on the limit of acceptable GL quality, which is largely determined by small amounts of unconverted carbon. Despite this, there were no 7520


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Table 5. Cold Gas Composition for Operating Points According to Table 2 op. point 1 2 3 4 5 6 7 8 9 a

H2 (%) 29.3 30.2 28.5 29.5 28.9 27.7 28.8 30.4

CO (%) 25.4 26.2 25.1 26.3 26.3 24.6 27.7 29.0

CO2 (%)

CH4 (%)

30.3 30.8 29.7 30.8 31.0 29.4 27.0 26.9

H2S (%)

0.20 0.25 0.21 0.14 0.06 0.17 0.19 0.25

2.42 2.63 2.49 2.59 2.54 2.41 2.31 2.33

COS (ppm) No Data 334 359 361 297 305 287 297 266

O2/Ar (%)

N2 (%)

H2Oa (%)

sum (%)

H2/CO

CO/CO2

0.02 0.02 0.02 0.02 0.02 0.05 0.02 0.02

10.8 9.1 13.1 9.8 10.7 14.6 12.8 9.4

0.2 0.2 0.2 0.2 0.2 0.2 0.2 0.2

98.7 99.4 99.5 99.4 99.8 99.3 99.0 98.5

1.15 1.15 1.13 1.12 1.10 1.13 1.04 1.05

0.84 0.85 0.85 0.85 0.85 0.84 1.03 1.08

Estimated based on saturation.

Table 6. Na Content in Gas Condensate Samples operating point

Na (mg/L)

1 2 3 4 5 6 7 8 9

1000 1480 1820 1820 2160 2280 1600 2000 2280

Table 7. Element Composition for Two Selected Gas Condensate Samples Ca Fe K Mg Na S

mg/L mg/L mg/L mg/L mg/L mg/L

op. point 7

op. point 9

1.15 0.465 53.3 1.83 1800 230

1.41 0.471 61.6 2.49 1910 276

Figure 2. CO/CO2 (●) and CH4 (○) in cold syngas as a function of the ratio between feeds of oxygen and thick liquor solids (TLS).

sulfite liquor experiment11 but the reason for this is not understood. As shown in Figure 2, the CH4 concentration is strongly dependent on O2/fuel ratio for constant liquor DS which is in agreement with previous experience for BLG.6,15 However, for the two operating points with higher liquor DS, a much lower oxygen flow can be used to reach the same CH4 concentration in the gas, which is likely explained by the fact that water behaves as a ballast in the reactor leading to higher oxygen consumption in order to reach the same temperature. CO/CO2, on the other hand, is not dependent on oxygen/fuel ratio for constant DS, indicating that the extra oxygen added oxidizes CH4 and/or hydrogen rather than CO. It has been shown previously that the water gas shift equilibrium controls bulk gas composition for BLG.9 No comparison between experimental syngas concentrations of H2, CO, and CO2 and the concentrations expected from water gas shift equilibrium in the rector is made, since thermodynamic equilibrium simulation of the gasification process is outside the scope of this paper. As shown in Figure 3, the CH4 concentration is more correlated to measured temperatures at the top and in the middle of the reactor than at the bottom. The correlation coefficients between CH4 and the top, middle, and bottom temperatures are 0.966, 0.945, and 0.751 respectively. Hence, strong correlation is present for temperatures in the top and middle part of the reactor in which a flame is present.16 Thermocouples are mounted level with the ceramic surface and are influenced by radiation and ceramic temperature. Hence temperature readings may not give an accurate representation

signs of accumulation or deposits on the gas side such as increased pressure drop or decreased heat transfer coefficients, e.g., due to tars. The tar concentration in cold gas has previously been shown to be extremely low for Kraft BLG.8 The H2/CO ratio (1.04−1.15, see Table 5) is lower than for normal Kraft BL operation (H2/CO 1.16−1.40 in cold gas7). This is in agreement with experience from the previous sulfite liquor experiment.11 However, H2/CO in cold gas has been shown to be dependent on quench operation;7 the higher than normal primary spray flow used in this experiment may thus have contributed to the lowering of the ratio. The published Kraft BLG syngas CO/CO2 reaches up to 1.03 for an operating point with 1250 kg/h BL flow with 73% DS BL.6 This CO/CO2 ratio is similar to operating point 8 of this experiment with 939 kg/h and 63% DS STL. For operating point 9 of this experiment with 1043 kg/h STL, CO/CO2 is 1.08. Hence, CO/CO2 ratios (Table 5) are generally higher than for typical Kraft liquor operation at similar reactor conditions. This is thought to be caused by the higher heating value of the sulfite liquor (cf. Table 1). CO/CO2 is also higher than in the previous sulfite liquor experiment (CO/CO2 0.70− 0.87)11 which is likely due to the lower load used in those experiments leading to higher relative heat losses. The CH4 concentration (0.06−0.25%) is much lower than for typical Kraft BLG in the same reactor temperature range (0.7−2.7% CH4),6 which is in agreement with the previous 7521


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originates in a recovery boiler or gasifier, and consist of mainly inorganic salts that are not soluble in GL. The GL quality was judged as acceptable or unacceptable using an experience-based assessment of the possibility to separate solids in the pulp mill, which is required prior to further processing of GL. Nonsettling particles, foam, or floating sludge hinders separation and is indicative of lower quality GL. Also a very high solids volume is a reason for nonacceptance. A film on the surface is normally formed when GL is cooled and is considered acceptable. On the basis of these criteria, operating points 3, 5, and 6 were judged to have unacceptable GL quality, whereas remaining GL samples were considered acceptable. It should be noted that the amount of unconverted carbon found in GL is very small in all cases. Using TOC (Table 3) and solids content (Table 4) in combination with GL flow estimates (see below) and making the approximation that 20% of GL solids are organic carbon, in agreement with a span for RB GL literature values for loss on ignition (LOI),18 leads to an estimated 0.1−0.2% of total carbon leaving the process as organic carbon with the GL for operating points 2−9. It has previously been shown for BLG that GL solids separation is not more difficult than for RB green liquor.19 It is not possible to completely separate effects of reactor residence time from atomization effects, since these two factors are not independent in this experiment. Atomization is dependent on many factors of which the most important is nozzle design. Operating parameters influencing atomization are primarily reactor load, reactor pressure, nitrogen atomization support, and liquor preheat temperature, of which the last influences through temperature dependence of liquor viscosity. All operating points are considered to give an atomization “power” that is possible to reach in a commercial scale gasifier without using support nitrogen. In addition, the last operating point, which has the highest load and produced acceptable GL, used very little nitrogen support but a higher liquor preheating temperature, which indicates that liquor preheating is important. It would have been beneficial to be able to test liquor preheating to >130 °C for 58% DS liquor and to try to increase load further with 63% DS liquor, but this was not possible within the restricted time of the experiment. Approximately the same wet liquor load gave acceptable GL for 58% DS and 63% DS, i.e., operating points 4 and 9. This indicates that wet load, rather than the liquor solids load, often used to measure capacity may be more relevant and suggests that gas residence time is important and that capacity is greater for high DS liquor. Approximately 1040 kg/h wet load corresponds to 17.5 t DS/d if 70% DS liquor is assumed as would be the case when a gasifier is integrated with the biorefinery. This number can be compared to the nominal capacity of 20 t DS/d on Kraft black liquor. However, because of the higher energy content of the STL compared to Kraft black liquor, the situation is reversed if energy is used as a capacity limiting parameter. The thermal input at operating point 9 is about 3.3 MWth compared to the nominal plant capacity of 3 MWth. Mass Balances. Mass balances include all flows shown in Figure 1, except cooling air and cooling water that are not mixed with the process flows. In addition, there are sealing water flows entering the process through the GL circulation pump and condensate recirculation pump. The GL circulation sealing water flow was not measured and was estimated to be 300 kg/h of water based on measurements when the plant was not in operation, but this estimate is uncertain. Nitrogen purge

Figure 3. Cold gas CH4 content as a function of reactor temperature in the top (blue, ●), middle (black, ○), and bottom (red, ×) of the reactor.

of gas temperature but are considered useful for detection of changes in temperature and flame intensity between operating points. Figure 3 thus indicates that flame conditions rather than global reactor temperature determines CH4 content in agreement with the findings of detailed simulations of CH4 degradation reactions in BLG.15 The temperature reading at the bottom of the reactor is believed to be the measurement most relevant for global temperature, i.e., the temperature of the gas and smelt leaving the reactor, since the measurement point is located relatively close to the reactor outlet. A possible way of estimating how well the global temperature is approximated by the reactor bottom temperature reading is to calculate the global temperature using a thermodynamic equilibrium model of the reactor. Such a calculation is, however, outside the scope of this paper. H2S and COS concentrations are higher than for Kraft BL operation (H2S 0.7−1.5%6), which is in agreement with the previous sulfite liquor experiment.11 The explanation is the higher sulfidity of STL compared to BL (Table 1). The sulfur split between syngas and GL is further analyzed and discussed below. COS concentrations in the hot gas are determined by equilibrium reactions in the reactor, which leads to an expected dependence on gas H2S concentrations and temperature. Furthermore, COS is known to be hydrolyzed to CO2 and H2S after leaving the reactor until reactions are quenched.7 The large flow of primary quench spray in this experiment can make gas quenching faster, leading to less COS hydrolysis. Plant Capacity. A major purpose of the experiment was to determine the pilot plant capacity on STL. From the results of the experiment, it is clear that GL quality (primarily TOC and properties of GL solids) is the factor limiting reactor capacity in agreement with previous experience from BLG. The results, specifically CH4 content, indicate that syngas quality is good at all operating points. Although higher hydrocarbons were not measured in this experiment, the overall concentration has been shown to be very low for BLG,8 and the concentration of the major component, benzene, has been shown to correlate well with CH4.17 Table 2 in combination with Table 3 and Table 4 contain relevant data for the determination of pilot plant capacity. Solids are always formed in GL, regardless of whether it 7522


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Figure 4. Method used for estimation of syngas flow based on gas cooler duty. The top row inputs are data from instrumentation in the plant or gas analysis. Shaded steps are performed based on Peng−Robinson thermodynamics (see text).

Table 8. Overall Mass Balance for Operating Points 2−9 op. point

O2/N2a (kg/h)

STLb (kg/h)

waterc (kg/h)

sum in (kg/h)

syngas (kg/h)

GL (kg/h)

cond. (kg/h)

sum out (kg/h)

in−out (kg/h)

(%)

2 3 4 5 6 7 8 9

336 399 417 404 408 417 412 414

833 1018 1017 1018 1019 949 939 1043

1302 1259 1617 1257 1307 1376 1345 1343

2471 2676 3051 2679 2734 2741 2696 2800

953 978 1025 947 941 891 901 950

1108 793 834 734 769 844 952 931

847 1005 1028 1042 1037 950 928 973

2908 2776 2887 2724 2746 2685 2780 2854

−437 −100 164 −45 −12 56 −84 −53

−18% −4% 5% −2% 0% 2% −3% −2%

Sum of oxygen to burner, nitrogen to burner, and nitrogen instrument purge. bWet flow, DS content according to Table 1. cSum of water to primary quench spray, GL dissolver, and sealing water to pumps. a

flows are also present for some instruments and included in the balances. A GL flow meter is present in the plant but is considered to give very uncertain measurements due to a highly fluctuating nature of the flow, which is caused by an unfortunate configuration of the GL level control loop and/or a malfunctioning valve used in this control loop. An estimated GL flow based on a closed Na+K balance of the unit has been used, as shown in eq 1, where ρGL is the density of GL measured online, Qcond is volumetric flow rate of gas condensate, FI is the mass flow rate of flow I, xNa+K is mass I fraction of Na and K in flow I, and cNa+K is the volumetric I concentration of Na and K in flow I. The choice of the elements Na and K is motivated by the fact that approximately 96% of Na and K leaves through this stream. FGL = ρGL

temperature and pressure in the bottom and top of the cooler, respectively. The estimation procedure is shown in Figure 4. The gas cooler duty was estimated to be 2.1−2.6 MJ/Nm3 dry gas for operating points 3−9. Small differences in measured raw gas temperature have large influence on calculated cooling duty, since the temperature is used to calculate water content at saturation in the gas before the cooler, which contributes to uncertainty in gas flows estimated by this method. The simulations do not account for alkali particles that exist in the raw gas and are dissolved in the condensate. Gas condensate samples were not fully analyzed for all operating points (Tables 6 and 7). The average sulfur content for the two samples analyzed (Table 7, 253 mg/L) was used for all operating points in the balances. On the basis of sodium content (Table 6) and sulfur content, the carbon content of the condensate was estimated by assuming that hydrogen carbonate is the counterion to all sodium that is not balanced by sulfide sulfur. Table 8 shows that the overall mass balances for operating points 3−9 close well. The largest deviation is for operating point 2 that was in the beginning of the experiment and has a lower GL concentration than the other samples (Table 3). The reason for the poor closure for this point is not understood. Accumulation of material is not likely, since the estimated flows from the process are larger than the flows into the process. Table 9 shows element balances for the most important elements from a process understanding point of view. Na+K balances close exactly since GL flows were estimated based on that assumption. 3.0−4.7% of the total Na+K entering the process is found in the condensate purge stream, which means

Na + K Na + K FTLSx TLS − Q condccond Na + K cGL

(1)

Similarly, the syngas flow measurement, which is based on orifice pressure drop, is unreliable due to difficulties in cleaning pressure sensor piping. Therefore, syngas flows were estimated based on gas cooler heat loads that were calculated using temperatures and flows of cooling water. The gas flow was then estimated based on calculations of gas side cooler duty using the Peng−Robinson equation of state in UniSim Design software (Honeywell). Measured cool gas composition for each operating point was used as dry gas composition in the Peng− Robinson simulations, and the gas was assumed to enter and leave the gas cooler in a saturated state at measured 7523


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Table 9. S, Na+K, and C Element Balances for Operating Points 2−9 op. point

STL (kmol/h)

sum in(kmol/h)

1.80 2.20 2.20 2.20 2.20 2.05 2.17 2.43

1.80 2.20 2.20 2.20 2.20 2.05 2.17 2.43

Na+K 2 3 4 5 6 7 8 9 C 2 3 4 5 6 7 8 9 S 2 3 4 5 6 7 8 9

17.2 21.0 21.0 21.0 21.0 19.6 20.8 23.2

syngas(kmol/h)

17.2 21.0 21.0 21.0 21.0 19.6 20.8 23.2

1.43 1.74 1.74 1.74 1.74 1.62 1.72 1.92

GL(kmol/h)

cond.(kmol/h)

sum out(kmol/h)

1.74 2.12 2.11 2.10 2.10 1.98 2.09 2.33

0.05 0.08 0.08 0.10 0.10 0.07 0.08 0.10

1.80 2.20 2.20 2.20 2.20 2.05 2.17 2.43

0.9 1.0 1.0 1.0 1.0 0.9 0.9 1.0

0.05 0.07 0.07 0.09 0.09 0.06 0.07 0.09

0.42 0.46 0.47 0.48 0.49 0.46 0.50 0.56

0.007 0.008 0.008 0.008 0.008 0.008 0.007 0.008

21.4 22.4 22.3 21.5 21.2 19.0 20.0 22.2

1.43 1.74 1.74 1.74 1.74 1.62 1.72 1.92

0.94 1.04 1.02 0.98 0.95 0.85 0.86 0.93

22.3 23.5 23.4 22.6 22.2 20.0 21.1 23.2

in−out(kmol/h) 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

0% 0% 0% 0% 0% 0% 0% 0%

−5.1 −2.5 −2.4 −1.6 −1.2 −0.4 −0.2 0.0

1.36 1.51 1.51 1.48 1.45 1.33 1.37 1.50

(%)

−30% −12% −11% −7% −6% −2% −1% 0%

0.06 0.23 0.23 0.27 0.30 0.30 0.36 0.42

4% 13% 13% 15% 17% 18% 21% 22%

Table 10. Overall Energy Balance for Operating Points 2−9 op. point

sens. inA (kW)

STL HHV (kW)

total in (kW)

sens. outB (kW)

GL HHV (kW)

syngas HHV (kW)

gas cooling (kW)

reactor cooling (kW)

total out (kW)

in−out (kW)

(%)

2 3 4 5 6 7 8 9

51 73 58 67 65 58 56 68

2482 3030 3028 3031 3033 2826 2999 3344

2532 3102 3086 3097 3098 2884 3054 3412

347 352 346 345 341 324 330 339

119 131 135 138 139 132 143 160

1906 2026 1995 1918 1853 1680 1862 2108

391 552 587 585 596 518 480 551

90 90 90 90 90 90 90 90

2853 3151 3153 3075 3020 2743 2906 3248

−321 −48 −67 22 78 141 148 164

−13% −2% −2% 1% 3% 5% 5% 5%

A

Sum of sensible heat of inputs; dominated by STL. O2, N2, and water have small negative values since they are below the reference temperature. Sum of sensible heats of products; approximately equal contributions from GL and condensate purge.

B

that it has been carried over with the raw syngas from the gasifier to the gas cooler. Since this amount of loss cannot be tolerated in a commercial gasification plant, such a plant would have to include a method for recycling that alkali, most probably by using the condensate purge as the dissolving liquid for GL generation. Note that the separation of alkali from the gas in the gas cooler is highly efficient as has been shown previously.8 Hence, alkali carry-over to the gas cooler is not expected to lead to impurities in cold syngas. Carbon balances close fairly well for operating points 5−9 but less well for points 2−4. The large deviations cannot be explained by nonsteady state in the GL since GL does not contain an appreciable amount of C. Similarly, the accumulation of carbon containing material in the process is not the cause, since the estimated flows of carbon out from the process are larger than those into the process (Table 9). The most likely reason for the deviations is that syngas flows are not accurately estimated by the gas cooler energy balance approach

used. For points 7−9 with good C balance closure, it is estimated that 95−97% of total C leaves with the syngas. The remainder is mostly found as sodium carbonate in GL. As noted above, the amount of unconverted carbon in GL is 0.1− 0.2%. Sulfur balances show deviations of up to 22% of total sulfur, which is a fairly high number and lead to suspicions of sulfur accumulation in the system. The total accumulated amount of sulfur that would be required to explain the differences between inputs and outputs in Table 9 is approximately 900 kg during the 100 h experiment, but no signs of such accumulation could be detected when the plant was inspected after the experiment. A more likely reason is errors in analyses and/or flow estimates. It can be noted that in a scenario where C balances are used to estimate syngas flow, i.e., where C balances close exactly, the S balances get a very similar relative mismatch of 20−22% for all operating points, which may indicate that there is a systematic error in the sulfur analysis method for one of the streams. The 7524


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Table 11. Syngas Heating Values and Cold Gas Efficiencies for Operating Points 2−9 op. point

HHVa (MJ/Nm3)

LHVb (MJ/Nm3)

SF-LHVc (MJ/Nm3)

H2+CO LHVd (MJ/Nm3)

CGE HHVa

CGE SF-LHVc

2 3 4 5 6 7 8 9

7.75 7.98 7.57 7.85 7.70 7.38 7.91 8.36

7.11 7.32 6.95 7.21 7.08 6.78 7.29 7.69

6.53 6.69 6.36 6.60 6.48 6.20 6.74 7.14

6.46 6.60 6.28 6.54 6.46 6.15 6.67 7.05

76.8% 66.9% 65.9% 63.3% 61.1% 59.4% 62.1% 63.0%

78.3% 67.8% 66.8% 64.2% 62.1% 60.4% 64.2% 65.5%

a

Higher heating value including sulfur compounds. bLower heating value including sulfur compounds. cLower heating value excluding sulfur compounds. dLower heating value of H2 and CO components.

about 91% of total gas LHV. Hence, almost 10% of syngas LHV is due to the presence of H2S and COS in the gas. For operating points 3−9, CGE values based on HHV are 59−67%, and the more relevant CGE values for SF-LHV are 60−68%. The CGE values in Table 11 show some correlation with temperatures measured in the reactor as expected (not shown) but are of course uncertain due to the uncertainty of syngas flow as discussed above. The estimation of CGE for a commercial scale gasifier would have to address the higher specific heat losses in pilot scale, the abnormally low STL DS content used here, and the 10% nitrogen in the synthesis gas due to pilot specific solutions. Such estimation would lead to a significantly higher number but is outside the scope of this paper. Impact on a Sodium Sulfite Mill Operation. The chemical recovery processes for sodium sulfite liquors that are currently used industrially are the Stora process, which is used at Domsjö Fabriker, and the Tampella process. Sodium sulfite liquors from a neutral sulfite semi-chemical (NSSC) pulping processes are commonly recovered together with a colocated Kraft process. Both the Stora and the Tampella processes include the separation of sulfur from the GL as H2S by lowering the pH. H2S is then oxidized to SO2 and recombined with sodium to form the cooking liquor for the sulfite process. With a gasification based recovery, this is facilitated both by the sulfur split in the gasification, as shown in Table 9, and the opportunity to choose the operating conditions yielding partially carbonated the GL, as shown in Table 3. As shown for a Kraft BLG processes, a full change to gasification based recovery has a large influence on the mill steam balance, since the energy in spent cooking liquor is used to generate chemicals instead of supplying process heat.1 The same situation applies for sulfite based delignification and process steam supply from another source will be required to operate the process when gasification based recovery is used. The gasification process enables a certain amount of low pressure steam production, primarily from gas cooling, and also the downstream processes for syngas upgrading can give excess heat that can be used in the delignification process. Quantification of this heat and thus the change in the mill energy balance requires assumptions about the configuration of the processes used for production of biofuels or chemicals from the syngas, which is outside the scope of this paper.

analysis of S in STL has an estimated relative uncertainty of approximately 12% as noted in Table 1. Furthermore, sulfur compounds are known to be susceptible to absorption by gas sampling system surfaces. The splitting of STL sulfur between the syngas and GL streams is an important aspect of process performance since it influences the chemical recovery in the biorefinery. Due to fact that S balances in Table 9 do not close exactly, different means of estimation of the S split leads to different results. Use of the STL and GL flows, which are considered more accurate than the syngas flow, leads to an estimate of 70−73% of total S in the syngas flow. Energy Balances and Cold Gas Efficiency. Energy balances include all flows included in the mass balances. In addition, gas cooler duty and reactor air cooling are also included. Gas cooler duty is estimated based on cooling water flow and temperatures as discussed above. Reactor heat loss was not based on measurements of cooling air flow and temperature for each point due to the large thermal inertia of the reactor system; instead an historic average of 90 kW heat loss from the reactor was used for all operating points. A reference condition defined by 25 °C, 1 atm, and C, H, and S in fully oxidized form was used when energy balances were calculated. This leads to the possibility of using higher heating values to represent stream chemical energy content, which makes interpretation easier. Table 10 shows that the overall energy balances close within 5% except for operating point 2, which also had poor closure for the mass balance. A small positive difference is expected since diffuse heat losses were not accounted for. Gas heating values and cold gas efficiencies (CGE) for operating points 2−9 are shown in Table 11. Similarly to overall energy and mass balances, results for operating point 2 are believed to be inaccurate. Four different heating values are shown; the sulfur-free LHV (SF-LHV) is a representation of useful energy, since in a commercial application H2S and COS would normally be removed from the syngas in a gas cleaning process before the syngas is utilized, which is also done in the pilot plant.5 In a gasification process integrated with a pulping process, all sulfur is returned in reduced form for the pulping chemical recovery process. For other integration options the removed sulfur may be used to produce sulfuric acid or elemental sulfur. H2 and CO are the useful components for production of fuels or chemicals from syngas, and hence, the energy in those components (as LHV) is presented separately in Table 11. Due to the low content of CH4 and other hydrocarbons, the values of sulfur-free LHV and LHV of H2+CO are very similar at

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CONCLUSIONS A 100 h high pressure entrained flow gasification experiment with spent cooking liquor from a sulfite-based delignification process has been described and analyzed. STL differs in 7525


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composition and properties from Kraft BL, which gives some differences in process behavior. Nevertheless, it was possible to operate the BLG pilot plant on the new feedstock without other modifications than the redesign of the liquor preheater. This indicates a large robustness and feedstock flexibility of the BLG process, which is being investigated further through planned pilot scale cogasification of pyrolysis oil and BL within the LTU Biosyngas Research Program. The process was deliberately operated on the limit of acceptable GL quality in order to determine plant STL capacity, which was found to be somewhat lower than for Kraft BL on mass basis but higher when measured as thermal load, due to the higher heating value of STL. The syngas quality was good for all operating points with CH4 content below 0.3% and H2/ CO ratio between 1.03 and 1.15. Mass and energy balances were made difficult by unavailability of GL and syngas flow rates, which necessitated the use of alternative approaches for estimation of these flows. By using these estimates, reasonably good closures of overall mass and energy balances were obtained, and CGE was estimated to be 60−68% on SF-LHV basis. Carbon balances indicate that 95−97% of feedstock carbon leaves with the syngas as CO, CO2, COS, or CH4. The remainder is mostly GL carbonate, while 0.1−0.2% is estimated to be unconverted carbon in GL. The sulfur balance closes less well than other elements but indicates that 70−73% of the feedstock sulfur ends up in the syngas with the remainder being present in GL as dissolved sulfide salts. 3.0−4.7% of feedstock alkali metals are carried over with the raw syngas to the gas cooler and ends up in the gas cooler condensate. The duration of this experiment and the analysis presented herein are not fully sufficient to enable a commercial deployment of the STL gasification process. Further tests with longer operating periods would be required to detect longterm problems such as material issues or accumulation of minor components in the system. In addition, analysis of syngas trace component content would be required to design downstream equipment in a biofuels plant based on STL gasification. It should be noted, however, that the close resemblance to Kraft BLG with 20 000+ h of pilot scale operation makes the requirement less severe than for many other new technologies.

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Article

REFERENCES

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel.: +46-920-492545. Present Address

R.S.: Scandinavian Biogas Fuels AB, Holländargatan 21A, SE111 60 Stockholm, Sweden. Notes

Chemrec AB has submitted patent applications (publication WO2011123034 A1) relevant for the research described in this article. The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Domsjö Fabriker AB is acknowledged for cooperation around the experiment described in this article and the supply of STL. The operating staff at the pilot gasification plant is acknowledged for their work during the experiment. 7526


Pilot Scale Gasification of Spent Cooking Liquor from Sodium Sulfite

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