Article pubs.acs.org/IECR
Design and Control of a Reactive Distillation Process for Naphtha Hydrodesulfurization Ankit Sharma, Ojasvi, and Nitin Kaistha* Department of Chemical Engineering, Indian Institute of Technology Kanpur, Kanpur 208016, India S Supporting Information *
ABSTRACT: Increasingly stringent regulations on the maximum permissible sulfur levels in transportation fuels has inspired considerable interest in reactive distillation (RD) based fuel hydrodesulfurization (HDS) in the refining industry. In this work, the steady-state economic design and plantwide control of an RD naphtha HDS process is studied. The process consists of an RD column followed by an absorber-stripper unit. Engineering heuristics are employed to develop a near-optimum economic process design. The important trade-offs in the design of the process are explained. A simple decentralized plantwide control system employing single-point temperature inferential control in the RD column with the fresh naphtha feed as the throughput manipulator is shown to provide effective process regulation for large changes in throughput, fresh naphtha thiophene composition, and feed hydrogen composition.
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Recently, Estrada-Villagrana et al.14 reported an RD process for naphtha HDS to produce light naphtha and ultra-low-sulfur [part-per-billion (ppb) level] heavy naphtha for blending with petrol. RD based processes for diesel HDS have also been reported.15,16 The reports suggest that RD based HDS is a promising alternative for retrofitting existing refineries. The optimum process design along with associated trade-offs and control of the processes is however not addressed. In this work, the steady-state optimum design and control of an RD based naphtha HDS process is studied. In the following section, the HDS process and its modeling details are presented. An economic near-optimum process design is then obtained using engineering heuristics with an elaboration of the trade-offs involved. A plantwide control system for the process is devised and evaluated dynamically for large changes in throughput, naphtha feed thiophene composition, and hydrogen feed composition. The conclusions summarize the major findings from the work.
INTRODUCTION Reactive (catalytic) distillation (RD) is a process integration technology combining reaction and separation in a single vessel. It is advantageous over conventional sequential “reaction followed by separation” processes when the relative volatilities of the components allow reactant recycling using stripping or rectification and acceptable reaction rates are achieved at the tray bubble temperatures. For equilibrium-limited reactions, the continuous removal of products drives the reaction to nearcompletion. The reaction might also simplify the downstream separation task by reacting away azeotropes. Reduction of energy and capital costs,1,2 enhanced reaction conversions,3 removal of azeotropes for easier separation,4 and compact plants5 are possible major benefits of RD technology. It has been commercially applied for esterification,6 etherification,7−9 olefin metathesis,10 and cumene production.11 Sharma and Mahajani12 published a comprehensive review of potential RD applications. Newer RD applications continue to be reported in the literature. In recent years, RD for the hydrodesulfurization (HDS) of transportation fuels has attracted interest because of increasingly stringent fuel sulfur regulations. Conventional HDS processes use large trickle-bed reactors to convert the sulfur compounds present in natural gas, crude oil, and other petroleum products such as petrol/diesel into hydrogen sulfide by catalytic treatment with hydrogen. In addition to meeting fuel sulfur content specifications, HDS is necessary because sulfur poisons downstream noble-metal catalysts. The petroleum refining industry is particularly interested in RD based fuel HDS as it is a compact process that can be retrofitted to existing columns to meet stringent fuel sulfur specifications without a reduction in throughput. The combination of reaction and separation in a single unit is made possible by structured catalytic packing (including monoliths) that provides a large specific surface area without a prohibitively large pressure drop. For example, monoliths can provide surface areas as high as 1850 m2 /m3 with reasonable pressure drops.13 © 2016 American Chemical Society
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PROCESS AND MODELING DETAILS Figure 1 provides a schematic of RD process for HDS of naphtha.The process consists of an RD column followed by a closed-loop H2S absorption system using aqueous monomethyl diethanol amine (MDEA) absorbent. Fresh naphtha and a H2rich feed are fed to the RD column consisting of a reactive (catalytic) section with a stripping section below it. Naphtha is fed into the catalytic section, whereas H2 is fed to the lower portion of the column. In addition to liquid-phase hydrodesulfurization, the naphtha is fractionated in the catalytic section to recover desulfurized C7 free light naphtha as the liquid distillate. Noncondensables (H2, H2S, etc.) are vented out along with some light (C1−C6) hydrocarbons. The Received: Revised: Accepted: Published: 1940
August 24, 2015 January 9, 2016 January 20, 2016 January 20, 2016 DOI: 10.1021/acs.iecr.5b03108 Ind. Eng. Chem. Res. 2016, 55, 1940−1951
Article
Industrial & Engineering Chemistry Research
Figure 1. Schematic of RD based naphtha HDS process.
C4 H4S Thiophene
Table 1. Hydrodesulfurization Reaction Kinetic Expressionsa no.
reaction
rate expression
k = k0 exp(−Ea/RT)
1
C4H4S + 2H2 → C4H8S
−rT = kTCCH−0.12CH20.42CTH0.52
8.09 × 10−4 exp(−6.53 × 104/RT)
2
C4H8S + 2H2 → C4H10 + H2S C6H10 + H2 → C4H12
−rTHT = kTHTCTHT0.84
1.2 × 10−4 exp(−8.87 × 104/RT) 9.77 × 10−4 exp(−9.21 × 104/RT)
3
−rT = kTCCH−0.12CH20.42CTH0.52
+ 2H 2 → Hydrogen
C4 H8S Tetrahydrothiophene
The tetrahydrothiophene irreversibly hydrodesulfurizes in the presence of hydrogen C4 H8S Tetrahydrothiophene
+ 2H 2 → C4 H10 + Hydrogen
n‐Butane
H 2S Hydrogen Sulfide
The H2 also hydrogenates unsaturated components as a side reaction. This is modeled as the irreversible side reaction of cyclohexene to cyclohexane
a
Thermodynamic package: Nonrandom two-liquid (NRTL) model. b Cj, bulk concentration of component j (mol/m3); k0, pre-exponential factor [mol/(s kg of catalyst)]; Ea, activation energy (J/mol).
C6H10 Cyclohexene
stripping section prevents C6 from dropping down to the bottoms. This ultra-low-sulfur (part-per-billion-level) heavynaphtha bottom product is the valuable one used for downstream blending with gasoline to bring its sulfur content below the desired specification. The vapor vent stream is desulfurized in an aqueous amine (MDEA) solvent H2S absorber, compressed, and recycled back to the RD column. A small fraction of the recycle gas is purged to prevent the buildup of light inerts in the loop. The H2S-rich amine is sent to the stripper where the H2S and absorbed hydrocarbons are stripped off. The hot H2S-lean amine from the stripper is sent back to the absorber after being cooled in a process-to-process heat exchanger (with the cold H2S-rich amine) followed by cooling to 40 °C in a cooler. Thiophene is used as the model desulfurization compound in the naphtha feed. It undergoes hydrogenation to form tetrahydrothiophene (THT) in the irreversible reaction
+
H2 Hydrogen
→ C6H12 Cyclohexane
This model reaction scheme was taken from the work of Estrada-Villagrana et al.14 The reaction kinetics used in our work were taken from the work of Irandoust and Gahne17 on naphtha HDS using a monolith catalyst and are reproduced here in Table 1 for completeness. Aspen Plus was used for steady-state and dynamic process simulations. The nonrandom two-liquid (NRTL) equation of state was used to model the liquid-phase activity coefficients. The vapor phase was modeled as ideal. The ASTM cut-point data and API gravity in Table 2 were used to obtain 11 pseudocomponents for the full-range naphtha feed. The pseudocomponents are also included in the table. The actual naphtha feed was thus assumed to contain these 11 pseudocomponents plus 686 ppm thiophene and 5 wt % cyclohexene. 1941
DOI: 10.1021/acs.iecr.5b03108 Ind. Eng. Chem. Res. 2016, 55, 1940−1951
Article
Industrial & Engineering Chemistry Research
process. Finally, the inlet temperatures of the naphtha feed (T F1 ) and the hydrogen feed (T F2) must be chosen appropriately. The rationale for the proper choices of the the design/operating variables is discussed next. Total Catalyst Loading (Wcat). The stringent part-perbillion-level sulfur specification on the bottoms product implies that essentially all of the thiophene must be converted to nbutane. This requires sufficient catalyst on the RD column reactive trays to ensure complete thiophene conversion. The desired light-naphtha/heavy-naphtha split in the column is such that the distillate mole fraction of PC188F (86.7 °C boiling point) and the bottoms mole fraction of PC162F (72.2 °C boiling point) are small. Because thiophene boils at ∼84 °C, it is distributed both above and below the naphtha feed. It is therefore reasonable to have catalytic trays both above and below the naphtha feed so that all of the thiophene can be hydrodesulfurized. We assume that the catalyst is equally distributed onto all of the reactive trays and adjust Wcat to meet the 90 ppb bottoms sulfur specification. Column Operating Pressure. To minimize the catalyst requirement, a higher column operating pressure is recommended so that the column temperatures are higher for enhanced reaction rates. Higher column pressure also minimizes the loss of light hydrocarbons in the vapor distillate and achieves a denser vapor for a lower column diameter. It also favors H2S absorption in the downstream absorber. However, higher pressure would cause the reboiler duty to increase as the separation becomes more difficult. Invoking Douglas’ doctrine18 of designing a process to minimize loss of precious raw materials/products as energy is significantly cheaper, the qualitative arguments suggest that the RD column should be operated at as high a pressure as possible. The use of high-pressure steam as a cheap reboiler heating medium, however, imposes an upper limit on the maximum column pressure. At 10 atm, the bubble temperature of heavy naptha is about 233 °C, giving a sufficient reboiler temperature driving force for high-pressure steam (255 °C) as the heating medium. The column operating pressure is thus fixed at 10 bar. Total Number of Trays. The RD column primarily acts as a PC162F/PC188F splitter with a bit of reaction for converting the small amount of thiophene in the naphtha feed. The desired split thus sets the number of fractionation trays required. From the Fenske equation (DSTWU module in Aspen Plus), the minimum number of trays for the desired split is 16. The total number of trays is then set as twice the minimum number of trays, or 32 trays. Column Feed Temperatures. The naphtha is fed to the RD column as a hot vapor. The naphtha feed temperature significantly impacts the required total catalyst loading to meet the bottoms 90 ppb sulfur specification, which decreases as the temperature is increased. For example, at an inlet temperature of 185 °C, the total catalyst requirement is more than 4 tons, which decreases to less than 1 ton for an inlet temperature of 250 °C. The hotter naphtha feed also reduces the column reboiler duty at the expense of a higher feed heater duty. Because the special structured RD catalyst packing is expensive ($30 per kilogram per year), the naphtha feed temperature is fixed at 235 °C (the highest feasible temperature using high-pressure steam) to minimize the catalyst requirement (and cost). Feed-Tray Locations (NF1 and NF2), Number of Reactive Trays (NRX), and RD Column Specifications. The fresh naphtha should be fed into the reactive zone, whereas the
Table 2. Feed Naphtha API and ASTM Cut-Point Data and Pseudocomponent Boiling Ranges API Data API gravity density at 15 °C (kg/m3) Distillation (ASTM D86) Cut Points
66.4 715
cumulative yield (vol %)
temperature (°C)
0 5 10 30 50 70 90 100
41 55 62 77.5 97.5 114 137.5 170 Generated Pseudocomponents
pseudocomponent
normal boiling point (°F)
specific gravity
molecular weight (g/mol)
PC93F PC137F PC162F PC188F PC213F PC237F PC261F PC286F PC312F PC337F PC355F
92.8 136.9 162.4 187.8 212.9 237.5 260.9 286.3 312.3 337.3 355.3
0.676 0.694 0.703 0.713 0.722 0.730 0.739 0.747 0.756 0.764 0.770
69.5 80.2 86.8 93.5 104.3 111.3 118.3 126.3 134.9 143.6 150.1
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STEADY-STATE PROCESS DESIGN The RD HDS process is to be designed for processing fresh naphtha containing 686 ppm thiophene and 5 wt % cyclohexene at a flow rate of 100 kmol/h. The product heavy-naphtha sulfur should be 1.6 wt % pseudocomponent PC162F (PC162F has a boiling range that matches those of C6 hydrocarbons) light key impurity with less than 90 ppb sulfur. The light liquid distillate naphtha should be 2.4 wt % pseudocomponent PC188F (PC188F has a boiling range that matches those of C7 hydrocarbons) impurity. The absorberstripper unit is to be designed for 4 ppm H2S in the fuel-gas purge. RD Column Design. The RD column (including column feed heaters) is arguably the most important process unit for this HDS process. Its design variables are the column pressure (P), total number of trays (Ntot), number of reactive trays, locations of the naphtha and hydrogen feed trays (NF1 and NF2, respectively), and catalyst loading per reactive tray (Wcat). As in previous reports,14−16 all trays above the naphtha feed are taken as reactive. Because unconverted thiophene drops down toward the bottoms, a few trays below the naphtha feed should be loaded with catalyst (reactive) to ensure complete thiophene conversion. The number of reactive trays below the naphtha feed (NRX) is thus a more convenient equivalent specification of the total number of reactive trays and is used here. In addition to these design variables, the hydrogen-to-thiophene molar ratio fed to the column (RF), the vent temperature (Tvent) (or, alternatively, the vent rate), the reflux ratio (RL), and the reboiler duty (QReb) are operating RD column specifications that must be chosen appropriately to meet product specifications while maximizing the total annual profit for the 1942
DOI: 10.1021/acs.iecr.5b03108 Ind. Eng. Chem. Res. 2016, 55, 1940−1951
Article
Industrial & Engineering Chemistry Research
Figure 2. Effects of (a) NRX and (b) NF1 on total catalyst loading and reboiler duty.
location of the naphtha feed tray (NF1; first subplot in panel b). Similar to ordinary distillation, NF1 affects the reboiler duty, and the feed location for the minimum reboiler duty is low in the column at tray 24 (second subplot in panel b). Increasing NRX causes a significant reduction in Wcat, with the reduction flattening out beyond NRX = 4 (first subplot in panel a). Also, NRX has very little effect on the reboiler duty (second subplot in panel a). We therefore choose NF1 = 24 and NRX = 4. For this choice, feeding hydrogen on tray 30 (one tray below the last reactive tray) appears reasonable and is used. With the hydrogen being fed close to the bottom, having its temperature to be the same as that of the naphtha feed (i.e., 235 °C) gives a temperature that matches the temperature of the bottom section of the column (220−233 °C) and is appropriate. Hydrogen-to-Thiophene Ratio (RF) and Vent Temperature (Tvent). Exploratory simulations showed that, as the hydrogento-thiophene ratio (RF) was increased, the percentage of hydrogen sulfide carried away by the vent gas increased. This implies that the amount of H2S in the liquid distillate was reduced. The vent gas, however, also carries away more light hydrocarbons (up to C6), thereby reducing the liquid distillate rate. Increasing Tvent has a very similar effect. For an appropriate choice of these two variables, Figure 3 shows the variations in the liquid distillate rate and the amount of H2S in the liquid distillate with RF for different feasible values of Tvent. A strict economic optimization would balance the premium earned by reducing the H2S content in the liquid distillate with the revenue lost by decreasing the liquid product rate. Such a trade-off requires reliable data on the variation in
fresh hydrogen should be fed below the reactive zone so that H2 is available on all of the reactive trays. An initial converged steady-state solution for the RD column is obtained with an arbitrary naphtha feed-tray location at tray 16 (middle of the column), with hydrogen (fresh plus recycle) fed much lower at tray 30. The total hydrogen feed rate is chosen so that the hydrogen-to-thiophene molar ratio to the RD column is high (initially chosen as 120 on a mole basis). The column is first converged without any reactive trays (i.e., ordinary distillation with zero catalyst per tray). The three RD column specification variables, chosen as the vapor vent rate, the reflux ratio, and the reboiler duty, are adjusted for a reasonable vent temperature (∼50 °C for a water-cooled condenser), a 2.4% PC188F impurity in the distillate, and a 1.6% PC162F impurity in the bottoms. The Design Spec Vary feature in Aspen Plus is used to meet these specifications. With no reaction, all of the thiophene drops down to the bottoms. To desulfurize the thiophene, the number of reactive trays below the feed tray is initially chosen as three (NRX = 3). Wcat is then increased until the bottoms thiophene impurity is 90 ppb. The location of the naphtha feed affects the reboiler duty. This location is adjusted to achieve the minimum reboiler duty for three reactive trays below the feed with Wcat determined for 90 ppb sulfur in the heavy-naphtha bottoms. The number of reactive trays below the naphtha feed tray is then further adjusted to reduce the total catalyst loading while meeting the target 90 ppb sulfur specification. Figure 2 illustrates the effects of NF1 and NRX on the reboiler duty and total catalyst loading. Note that, at a given value of NRX, Wcat is insensitive to the 1943
DOI: 10.1021/acs.iecr.5b03108 Ind. Eng. Chem. Res. 2016, 55, 1940−1951
Article
Industrial & Engineering Chemistry Research
Figure 3. Effects of RF and Tvent on the (a) flow rate and (b) H2S content of the liquid distillate.
Table 3. RD Column Design Variables and Associated Variables and Selected Values variable
value
Design Variables pressure (P) total number of trays (Ntot) naphtha feed-tray location (NF1) hydrgen feed-tray location (NF2) catalyst loading (Wcat) Associated Variables hydrogen-to-thiophene molar ratio (RF) vent temperature (Tvent)
10 atm 32 24 30 729 kg 250 80 °C
the price of light naphtha with sulfur content, which is difficult to obtain or even conjecture. We therefore use simple qualitative arguments to choose reasonable values for the two variables. Given the premium associated with low-sulfur-content fuels, the percentage of H2S removed in the vent should be high (say, >60%) so that the sulfur content in the liquid distillate is low (
