1. Crude oil Atmospheric Distillation Unit

Transcription

1 Petroleum Refining Chapter 7: Distillation Chapter 7 Distillation The first processing units in the refinery are based on physical separation. There are two types: 1. Atmospheric distillation Units 2. Vacuum Distillation units 1. Crude oil Atmospheric Distillation Unit Introduction The first unit in the refinery. Based on physical separation at atmospheric pressure. Older units and those processing heavy crude oil tend to combine crude oil distillation and vacuum units in one. If the crude oil processed is heavy, then a pre-flash column is used which is followed by a vacuum tower (MAB CDU-01). Newer designs tend to separate the atmospheric and the vacuum distillation units. Here we will describe a modern crude oil distillation unit from mina Abdullah refinery. That will include the operating conditions and controls in such unit. Table 7-1. Atmospheric Distillation Units in Kuwait. Refinery Name Unit Throughput (BPSD) Feed MAA Distillation CDU-3,4,5 466,000 Kuwait export crude Eocene Unit 39 24,000 Eocene ZOR Distillation Unit-06 [615,000] Kuwait export crude Eocene [ ] Eocene MAB Distillation CDU ,000 Kuwait export crude Distillation CDU ,000 Total 760 [1,3750,000] Process Description MAB CDU The capacity of this unit is 150,000 BPD or 156,250 BPSD with operating factor of 0.96 The feed to the crude unit is a blend of three cruds (Table 7-2) The products are shown in Table 7-3 Heavier crude (lower API) produces more atmospheric residue and less distillate (Diesel and lighter). The opposite is also true. The process flow diagram can be divided in four parts, shown in Figure Crude exchanger (preheat) train with desalter & pre-flash. 2. Crude Heater. 3. Crude fractionator. 4. Stabilizer. 7-1

2 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Table 7-2: The feed to the crude oil distillation unit. Vol. % API S, wt% Kuwait export crude Ratawi Burgan Kuwait crude mix Table 7-3: Products from the crude distillation unit. Product Vol. % of Feed Destination Sour gas LPG Stabilized Naphtha Kerosene Diesel Atm. Residue Gas treating Gas treating PCN (petrochemical naphtha) tank Kerosene HTU or tank Diesel HTU or tank ARDS, Isomax, heavy cooling. Table 7-4 TBP Cut Points for Various Crude Oil Fractions Cut IBP(⁰F) EP(⁰F) Processing use LSR gasoline cut Min Light gasoline Normal LSR cut Max LSR cut HSR gasoline (naphtha) Kerosene Max reforming cut Max jet fuel mode Min reforming cut Max kerosene cut Max jet-50 cut Max gasoline operation Light gas oil * 610* 610* Max diesel fuel Max jet fuel Max kerosene Heavy gas oil ( HGO) Catalytic cracker or Hydrocracker feed Vacuum gas oil , SDA or catalytic cracker feed Catalytic cracker or hydrocracker feed Note: In some specific locations, economics can dictate that all material between 330 ⁰F IBP and 800 ⁰F EP (166 to 427 ⁰C) be utilized as feed to a hydrocracker. *For maximum No. 2 diesel-fuel production, end points as high as 650⁰F (343⁰C) can be used. 7-2

3 Petroleum Refining Chapter 7: Distillation Figure 7-1 Process flow diagram for a crude distillation unit 7-3

4 Temperature (F) API Gravity Temperature (F) API Gravity Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Example 7-1: CDU Material Balance From the TBP curve below, calculate the amount of Naphtha (C5 300 ºF), kerosene ( ºF), Diesel ( ºF), and Residue (650 + ºF) that can be produced in a 100,000 BPCD crude distillation unit, with a 31.7 API crude oil Percent Distilled Solution: Crude oil Calculate the mass flow rate of the crude oil charge SG = 141.5/( ) = Mass flow rate = BPD (SG) = 100,000 (0.867) = 1,265,000 lb/hr Percent Distilled 7-4

5 Petroleum Refining Chapter 7: Distillation Naphtha Volumetric flow rate = 12 0 = 12 V% = ,000 = 12,000 BPD Vaporization loss = 0 (given) Check your crude oil assay for other values API = 57 (at mid point of 6 V%) SG = = Mass flow rate = BPD (SG) = 12,000 (0.7507) = 131,400 lb/hr Kerosene Volumetric flow rate = = 18 V% = ,000 = 18,000 BPD API = 40 (at mid point of (30+12)/2 = 21 V%) SG = 141.5/( ) = Mass flow rate = BPD (SG) = 18,000 (0.8251) = 216,700 lb/hr Diesel Volumetric flow rate = = 33 V% = ,000 = 33,000 BPD API = 31 (at mid point of (30+63)/2 = 46.5 V%) SG = 141.5/( ) = Mass flow rate = BPD (SG) = 33,000 (0.8708) = 419,250 lb/hr Residue Volumetric flow rate = = 37 V% = ,000 = 37,000 BPD Mass flow rate = 1,265, , , ,250 = 497,650 lb/hr SG = (lb/hr) = 497, = API = = (BPD) 14.59(37,000) HW It is now time to do your HW on CDU and VRU material balance with cost(s) and utilities! Every time you submit a HW, you must include original copies of all your previous HWs corrected from mistakes if any! You must always tabulate your final answers following the same format in the textbook. Crude Preheating Train Crude oil is heated to about 700 ºF before it is introduced to the fractionator. For this purpose a set of heat exchangers is used to raise the temperature of the crude oil from storage temperature ( ºF) utilizing the hot products from the fractionator in addition to the pumparounds as heating mediums for heat recovery. The crude is first heated using the fractionator overhead (O/H) vapor and top pumparound (TPA). This is followed by the kerosene and diesel products that are used in parallel. This insures a crude temperature of 260 ºF which is the temperature required for the desalting process. After the desalter, crude is heated by the middle pumparound (MPA) then kerosene, diesel and atmospheric residue products respectively to the pre-flash drum temperature of 415 ºF. The flashed crude is then further heated by the bottom pumparound (BPA) and atmospheric residue to a temperature of 500 ºF. 7-5

6 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Example 7-2: Energy Balance around a heat exchanger Calculate the duty of kerosene product/crude heat exchanger in a crude oil distillation unit. The crude oil flows at 156,250 BPD and has an API of 31 and a Kw of The kerosene flows at 21,740 BPD and has an API of 46.5 and a Kw of 12.0 Kerosene 400 o F Crude oil 210 o F Crude oil? o F Solution: Kerosene 255 o F For the kerosene product side Sp.gr. = 141.5/ (API ) = 141.5/( ) = Kerosene flow in lb/hr = BPD (14.59) sp.gr. = (21,740) (14.59) (0.795) = 252,163 lb/hr From figure 7B4.5 in the API TDB for Kw = 12.0 Enthalpy of 46.5 API liquid at 400 ºF = 296 Btu/lb Enthalpy of 46.5 API liquid at 255 ºF = 205 Btu/lb No correction for pressure is needed since we are dealing with a low pressure unit (CDU). Q = ( ) Btu/lb (252,163) lb/hr = 23 MM Btu/hr Ans For the crude oil side Sp.gr. = 141.5/ (API ) = 141.5/( ) = Crude oil flow in lb/hr = BPD (14.59) sp.gr. = (156,250) (14.59) (0.871) = 1,985,608 lb/hr From figure 7B4.5 in the API TDB for Kw = 11.8 Enthalpy of 30 API liquid at 210 ºF = 170 Btu/lb No correction for pressure is needed since we are dealing with a low pressure unit (CDU). Q = (Hexit 170) Btu/lb (1,985,608) lb/hr = 23 MM Btu/hr The enthalpy of crude oil at the exit, Hexit = Btu/lb From figure 7B4.5 in the API TDB for Kw = 11.8 The temp of 30 API liquid at Btu/lb = 230 ºF Ans 7-6

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15 Petroleum Refining Chapter 7: Distillation Desalters Salt content is expressed as PTB (lbnacl equivalent /1000 bbl crude) If crude has a salt content more than 10 PTB it is necessary to desalt the crude before processing. Desalting criteria; PTB max. at desalter outlet. - 1 PTB max. in atmospheric residue to downstream. 1 PTB 3ppm (depending on crude oil density) [for sp.gr. = (17 ºAPI)] Example 7-3: Conversion from PTB to ppm 1 PTB = 1 lb / 1000 bbl = 1 lb / 42,000 US gal m = ρoil V = (sp.gr. ρw ) V = [sp.gr lb/gal] V sp.gr. = m = API API V PTB = 1 lb / [42,000 US gal API ] 1 PTB = (API+131.5) ppm Crude oil is desalted and dehydrated at KOC oil fields from 10,000 PTB to 10 PTB, then sent to the refineries for settling in storage tanks before desalting. There are two horizontal desalters in series (90% eff. ea.) Number of stages required for desalting is shown in Table 7-5 Table 7-5: Number of stages required for desalting crude oil. No. of stages Efficiency Salt in crude oil (PTB) Comment Not very common Very common in refinery Very common in oilfield Function The function of the desalters in the CDU is to reduce salt content of crude oil feed from 20 PTB to 0.47 PTB and the BS&W to 0.05 vol%. As an added bonus, metals (inorganic compounds dissolved in water emulsified with the crude oil) which can cause catalytic deactivation in downstream catalytic units are also removed in the desalter. A secondary but important function of the desalting process is the removal of suspended solids from crude oil. Those are; - Very fine sand, clay, soil particles (from ground) - Iron oxide, iron sulfide (from pipelines, tanks, other contaminants) (Solids removal criteria 60 to 80 % minimum) Desalting is necessary if salt content of crude is greater than 0.5 PTB (The specification is 1 PTB max. in atmospheric residue). 7-15

16 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Salt in crude originates from the underground salt formations and exists in crude in two forms; - Dissolved in the water emulsified in crude - Crystalline form in the crude. The problems associate with salt presence in crude oil are many among which are; - Fouling of heat exchanger and fractionator trays due to salt deposition. - Catalyst deactivation/poisoning in ARDs & other (downstream) catalytic units. - Corrosion due to high temperature decomposition of salts to HCl acid NaCl + H + Δ H + Cl - This is the major cause of corrosion in the fractionator O/H system 1 and heater tubes). Operation Principle 1. Water is first mixed with the crude feed, to dissolve the naturally existing salts in crude oil. 2. The crude is then passed through a mixing value to ensure an intimate mixing and the formation of a water-in-oil emulsion (water exists as fine droplets in oil). 3. Inside the desalter the crude oil passes through two horizontal grids through which a high voltage electrical field (20 to 24 kv) is applied This electrostatic desalting electrically charges the fine water droplets creating a dipole causing the water droplets to elongate. 5. Since the applied current is alternating (AC) with a frequency of Hz, the grids reverse their polarity 50 to 60 times per second. 6. This causes the polarity of the fine water droplets to alternate with the same frequency and the droplets to vibrate vigorously. 7. The dual effect of vibration and polarity causes the water droplets to combine (coalesce) forming larger droplets. 8. When the droplet is large enough, it settles down under the effect of gravity. 9. To aid the demulsifying process, chemical demulsifiers are injected to the the desalters 3 feed. These chemical compounds reduce the surface tension of the water droplets making it easier for them to combine. 1 This the reason cladding is used for the fractionator O/H in addition to special material of construction of the top trays (usually top 3 trays). In addition to the injection of corrosion inhibitors and NaOH in the O/H (NaOH neutralizes HCl according to the following reaction NaOH + HCl H2O + NaCl ppt ). Precipitates are washed in the reflux drum. 2 The actual voltage is determined by trial and error during unit start-up. The desalter is equipped with a voltage selector ranging from 16,000 to 35,000 volts. 3 Wetting agents (oxyalkylated phenols & sulfates) are frequently added to improve the water wetting of solids (so they may be easily removed from oil to the water) and reduce oil carryunder in desalters. Suspended solids are one of the major causes of water-in-oil emulsion. 7-16

17 Petroleum Refining Chapter 7: Distillation The feed to the 1 st desalter is preheated to 260 ºF which is essential to reduce the viscosity of the crude oil 1 and make the desalting process more efficient (ease the separation of water from crude in addition to the formation and breaking of water-in-oil emulsion). Higher temperatures are not desirable because of vaporization problems. This temperature increase, requires a corresponding increase in pressure (170 psia and 140 psia for 1 st and 2 nd stages, respectively) to eliminate any vapor formation in the desalters that would affect their operation. Any interruption in desalter operation will cause water carryover with the crude which appears as high salt and BS&W. Excessive water carryover can cause short circuiting of the electrodes and the desalter will trip. Example 7-4: Demulsifier injection rate Calculate the minimum amount of demulsifier (sp.gr. 0.8) in GPH required for desalting 156,250 BPSD crude oil. Solution: Amount of demulsifier ranges between to 0.01 lb/bbl of crude oil feed = (156,250) = 780 lb/day = 32.5 lb/hr = 5 GPH up to 10 GPH Ans. 1 For very heavy crude oils (< 15API ) addition of gas-oil as a diluents to the second stage is recommended to reduce viscosity and provide better separation efficiencies. 7-17

18 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Figure 7-2 A two-stage electrostatic desalting system. 7-18

19 Petroleum Refining Chapter 7: Distillation Figure 7-3 Schematic of a desalter (cross-section view). Figure 7-4 Two-stage Desalter system at an oilfield. 7-19

20 Desalted Crude Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Transformer 3 phase Oil layer Water-in-oil emulsion layer Water layer Crude oil feed Figure 7-5 Internal assembly of an electrostatic desalter. 7-20

21 Petroleum Refining Chapter 7: Distillation Desalter Wash water Instead of using fresh water for both desalter stages, the fresh (wash) water is introduced to the second stage desalter first and the effluent water from the second stage is routed to the first stage. This was found to conserve water and increase the efficiency of the desalting process, since salt removal from the second stage is less than that from the first and the effluent water from the second stage desalter contains less salt than that from the first stage. If the water is injected into the first stage first, then most of the salt will go into the wash water which upon injection into the second stage will be brine (it will not be able to dissolve the salt from the second stage which is relatively low) and will reduce the desalting efficiency. In MAB refinery fresh water is used as wash water. In other refineries, the wash water is obtained from other refinery sources containing some contaminant (e.g. phenols) that are soluble in crude oil. This is desirable since it reduces the contaminants sent to the refinery waste water treatment system. Enough degree of mixing is provided to ensure enough contact between the wash water and the (dissolved and crystalline) salt in crude oil. The degree of mixing is controlled by the pressure drop across the mixing valve. - Too low pressure-drop will not suffice to mix the wash water with the crude and create the water-in-oil emulsion necessary to dissolve the salt in crude oil. - Too high pressure-drop will create unbreakable emulsion, which is also undesirable. In both cases, desalter efficiency will decrease and salt content of the crude oil at the desalter effluent will be high. In the latter case, the water content of crude will in addition be high which is also undesirable. The amount of wash water used is approximate 5 vol. % of crude feed. Using more than that will cause carry-over of the wash water (high in salt content) with the resultant - Increases the sodium content of the ARDS feed. - May result in discolored or dirty distillate products from the crude unit. Both water injection and desalter temperature are a function of crude API (Table 7-6) Example 7-5: Amount of wash water in desalters Calculate the amount of wash water (GMP) required for desalting 156,250 BPSD of 31 ºAPI crude oil at 260 ºF desalter temperature and the amount of caustic to maintain the PH. Solution From Table 7-6 below, the percentage of wash water = 5 vol % (of crude) Amount of wash water = 0.05 (156,250) = 7,813 BSPD = 7,813/34.3 = 228 GPM ans Caustic dosage is to lb NaOH / bbl of crude oil feed Amount of caustic = 0.001(156,250) = lb/day lb/day Table 7-6: Typical desalter operating conditions. API Wash water ¹ (vol.% of crude) Desalter Temperature (ºF) > < Desalter Pressure ² (psig) ¹ Use interpolation to begin with then optimize during operation. ² The pressure required to prevent vapor formation at the desalter operating temperature. This is equal to the crude oil bubble point pressure plus a safety factor

22 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Desalter Water ph Better dehydration is obtained in electrical desalters when operated in the ph range of 6-8 with the best dehydration and desalting obtained at a ph = 6 (neutral). If the ph of the brine exceeds 8, hard to break emulsions can form because of the (Sodium Naphthenate and Sodium Sulfide) present in crude, reducing desalting efficiency. The ph value is controlled by the addition of acid or base to the inlet or recycled water. NaOH (caustic soda) is used in MAB refinery to control PH (dosage is to lb NaOH / bbl of crude oil feed). See example below! When processing tank bottoms or slops 1 or when water or steam condensate contaminated with oil is used as injection water, a stable oil-in-water emulsion is formed. A demulsifying chemical is essential to break the oil-in-water emulsion (opposite to the above water-in-oil). Heavy naphthenic crudes form more stable (hard to break) emulsions than most other crude oils - Desalters usually operate at lower efficiencies when handling these. - Crude oil densities are close to that of water & temperature above 280 ºF are needed. Factors affecting desalting efficiency (separation ease) 1. Crude oil API. 2. Crude oil viscosity. 3. Crude oil olefin content (slops) 4. Volume of wash water used/volume of crude. 5. Wash water ph. 6. Degree of mixing (water/oil) ΔPmixing. 7. Degree of water washing of suspended solids. 8. Degree of separation of wash water from the oil. 9. Voltage applied. 10. Demulsifying chemicals applied. Dual Field Desalting: (not in Kuwait) A process that uses combined AC/DC currents to provide high dehydration efficiency (99% per stage is claimed). - An AC field is applied near the oil-water interface - A DC field is applied in the oil phase above the interface Advantages Provides efficient water separation at temperatures lower than other processes (save on the cost of heating especially for oil field desalting where fired heaters are used to heat-up the crude). Lower pressures can be used as a result of lowering the temperature (which reduces cost of pumping or NPSH requirement). 1 Using slopes (off-spec material from all around the refinery) from slop tanks reduces desalting efficiency because of the presence of cracked and olefinic material which act as emulsifying agents. 7-22

23 Petroleum Refining Chapter 7: Distillation Desalter Process Variables and Control Temperature Control The desalter should be operated in the range of 250 ºF to 260 ºF. This is controlled with the help of 1. The bypass controllers on heat exchangers E103 and E104 (Figure 7.2). 2. By increasing the heat recovery from TPA (by controlling the TPA flow and T). 3. If this is not enough then E105 and E109 may be bypassed to increase the inlet temperature thus heat recovery in desalter upstream heat exchangers E103 & E104. If the desalter inlet temperature is too high the opposite of the above procedure can be followed by 1. Lowering the heat recovery in top pumparound first. 2. Then by passing one of the shells of E-101 if necessary. Bypass TC Kerosene TI Crude oil Diesel Figure 7-6 Desalter inlet temperature control scheme Pressure Control The second stage desalter must be maintained at least 20 psi above the crude vapor pressure at the operating temperature to avoid vaporization. The pressure is usually controlled in the 2 nd stage desalter by a back pressure controller located on the effluent line. The pressure of the first stage is set by the pressure drop between the two desalters. The pressure drop across the mixing value is controlled by a field operated controller for both stages. The water and Demulsifying chemical injection rate is controlled through a manual flow controller. Controlling desalters interface level is very essential for efficient reliable operation. Low water level in desalter results in oil carry-under in the water effluent whereas very high interface level may result in short circuiting the electrodes. Therefore, the level in both desalters is control by a level controller which controls the flow of desalter effluent water. 7-23

24 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering PI PC LC LI Figure 7-7 Desalter pressure and level control scheme. Preflash drum (Flashed Vapor) low boiling material to fractionator flash zone. Heated Crude 415 o F 80 psig Preflash Drum (Flashed Crude) high boiling material to heater then fractionator flash zone Figure 7-8 The crude oil distillation unit pre-flash drum In the pre-flash drum, light ends produced during preheat of the crude oil are separated. 1. This reduces the pressure drop in the system (one phase flow instead of two). 2. Reduces the load on the heater and the heat exchangers following the flash drum. 3. Provides a better distribution of flow inside heater tubes, and thus, better heat transfer. 4. Smaller feed to HTR result in the design of a smaller & lower cost furnace. 5. Avoid excessive vapor generation or two phase flow, requiring special material of construction, thicker tubes, and extra mechanical support to avoid vibration all of which are costly. 7-24

25 Petroleum Refining Chapter 7: Distillation The flashed vapors are sent directly to the fractionator whereas the remaining crude is sent to the heater at 500 ºF after exchanging heat with bottom pumparound (BPA) and atmospheric residue. The crude is flashed at a temperature of 415 ºF and a pressure of 80 psia. This temperature and pressure are calculated to have a flashed vapor in the range between 5 and 10% of the crude oil. - Less than that will cause crude vaporization before entering the heater. - Higher than the will require an increase in the heaters outlet temperature (to maintain the same amount of heat in the feed to the fractionator) which might cause cracking and discoloration problems in the distillate product. The Pressure in the flash drum is maintained by a back Pressure controller which controls the vapors going to the flash zone of the fractionator. The temperature is controlled through bypassing of part of atmospheric residue around E110. The level in the Flash drum is controlled by a level controller that sets the two flow controllers on E103 and E104. PC atm Residue ResdResidue 415 F Vap.to frac. Desalted Crude From E109 From E114 T C 80 psi 415 F 415 F V101 L C To E113 S To FC on E103, 104 Figure 7-9 Pre-flash drum control system 7-25

26 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Example 7-6: Equilibrium Flash Vaporization (EFV) Calculate the amount of vapor and liquid produced in a flash drum operating at 600 ºF and 80 psig when 20,000 BPSD of the following (30 API) fraction is introduced. Also calculate the API and reconstruct the ASTM curve for the top product (Flash) and bottom product (reduced crude - this is the feed to the heater). ASTM D86 Distillation Vol. % Temp (ºF) ºF Solution: 1. Correct the ASTM D86 Distillation temperatures above 475 ºF for cracking: log D = ( )T log D = ( )(520) = D = 7 ºF The correct 90%T = = 527 ºF 2. Find the EFV 50%T from Figure 3B1.1 in the API technical data book: ASTM (10% to 70%) slope = = In figure 3B1.1 start with ASTM D86 50%T = 440 (x-axis) go up to the line with slope of 2.1 (this line does not exist so try to draw it between lines 2 and 3). The y- axis gives F = - 7 ºF (correction to be added) The EFV 50%T = ASTM D86 50%T + F = = 433 ºF 3. Obtain the temperature difference for each portion of the EFV curve from Figure 3B1.2: Segment of ASTM T (ºF) EFV T (ºF) EFV T (ºF) Vol% Curve (% Vol.) 10 to to 50 (Given above) (Figure 3B1.2) % 30% 50% 70% 90% 50 to to The EFV temperatures (at 1 atm) are calculated from the EFV 50%T of 433 ºF as follows: EFV 50%T = 433 ºF. EFV 30%T = EFV 50%T EFV T30 to 50 = = 411 ºF. EFV 10%T = EFV 30%T EFV T10 to 30 = = 381 ºF. EFV 70%T = EFV 50%T + EFV T50 to 70 = = 450 ºF. EFV 90%T = EFV 70%T + EFV T70 to 90 = = 476 ºF. 5. Correct the EFV temperatures (at 1 atm) to (95 psia) as follows: ASTM (10% to 90%) slope = = VABP = T T T T T90 = = ºF. 5 5 Calculate the value of the ratio Focal point = ASTM VABP = = 24.1 (10 to 90% slope)

27 Temperature (F) Petroleum Refining Chapter 7: Distillation From the right scale of the x-axis in figure 3B3.1 locate the VABP of 440 ºF. Move up to 30 API, then left to 2 ASTM (10% to 90%) slope, and finally move up to a ratio value of 24. This is called the focal point. From the focal point draw straight lines connecting the atmospheric EFV temperatures (the left scale) at 15 psia. This is called the phase diagram. Each of these lines represents a constant vol.% vaporized, and the EFV temperatures can be read for various pressures. EFV curve at 80 psig Vol. % Temperature (ºF) Draw the EFV curve (at 80 psig): At 600 ºF (y-axis), the vol.% (x-axis) is equal to. This is the volume 30 % vaporized. Top product rate = 20,000 (0.3) = 6,000 BPD Bottom product = 20,000 (1-0.3) = 14,000 BPD ASTM D86 Temperature (ºF) EFV Temperature (ºF) at 1 atm EFV Temperature (ºF) at 95 psia % Volume % 7. The ASTM D86 Distillation for the top product are calculated using Figure 3B4.1 from the API technical data book. From 30 vol.% vaporization (x-axis), we go to the lines of ASTM 10-70% slope of charge 2.1 (calculated in step 2 above), then move horizontally to the 10% line, then 7-27

28 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering vertically to the 80 psig pressure, then horizontally again to the y-axis and read the temperature correction value of 25 ºF. The same procedure is repeat for the 30, 50, 70, 90% lines and tabulated below. Distillation Temp (ºF) Vol% Charge Correction Top prod The ASTM D86 Distillation for the bottom product are calculated same wise but using Figure 3B4.2 from the API technical data book. Distillation Temp (ºF) Vol% Charge Correction Bottom prod The API gravity for the top and bottom products are obtained from figure 3B4.3 ASTM D86 10% to 30% slope for charge = = Starting from 30% vaporization on the x-axis moving up to the lines of constant charge slope of 2.5 (for overhead) then horizontally to a charge API of 30 then moving up to an overhead API of 35. The same is repeated but using the lines of constant charge slope of 2.5 (for bottoms). This gives a bottoms API of

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42 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Feed API = 30 ASTM D86 V% IBP FBP % Liquid P = 80 psig T = 600 F 30% Vapor 70% Liquid Top Product API = 35 ASTM D86 V% T (F) IBP FBP Bottom Product API = 28 ASTM D86 V% T (F) IBP FBP Figure 7-10 Example solution 7-42

43 Petroleum Refining Chapter 7: Distillation Crude direct-fired Heater Figure 7-11 CDU heater system simplified. The crude heater raises the temperature of the crude from 500 ºF to about 700 ºF. This is sufficient to vaporize all the products withdrawn above the flash zone plus the over-flash (10 to 20% of the bottoms product) This is equivalent to the total gas, naphtha, kerosene, diesel products and over-flash (10-20% of and fractionators bottoms) minus the flash drum vapors. At heater outlet Amount of liquid = 0.9 (atm reside) Amount of vapor = (Crude oil - Flash drum vapors) 0.9 (atm reside). Amount of vaporization = amount of vapor / amount of liquid The crude enters the heater through six passes into the convection section where it is heated by the flue (combustion) gases, and then enters the radiation section for further heating. The six passes join together at the heater outlet and enter the flash zone of the fractionator. The heater is equipped with an air preheat system to increase the efficiency (91.5% guaranteed) The heater is both gas and oil fired (for flexibility) and can be operated on natural draft with a by-pass around the air preheated during emergencies. 7-43

44 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Heater Process Control Figure 7-12 CDU Heater Process Control System. The total flow to crude heater is divided among six passes by a flow control valve on each pass, reset by the total crude flow controller. The skin temperature of each coil can be taken as an indication to the flow in each coil. If the skin temperature of one pass is much higher than the others, the flow in that pass is insufficient and must be increased. Heater outlet temperature is controlled by a temperature controller that regulates the amount of fuel (gas or oil) burned in the heater. This temperature should not exceed 700ºF to prevent cracking of the feed or damage to the heater tubes. The temperature controller is also connected to a flow controller that regulates the amount of air to the combustion chamber in the heater. Example 7-7 Stoichiometric equation for combustion & Emissions Factor IFP P179 Example 7-8 Heater efficiency 7-44

45 Petroleum Refining Chapter 7: Distillation Figure 7-13 CDU Heater Passes. Figure 7-14 Process heater with air preheat system 7-45

46 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Crude Fractionator The atmospheric fractionator normally contains trays. The crude fractionator at MAB has about 41 trays bubble cap trays above the flash zone (fractionation zone or enriching section). - 6 sieve trays below the flash zone (stripping zone or stripping section). - Despite their higher efficiency, bubble caps are avoided in the stripping zone of crude oil fractionators for 2 reasons; 1. High velocity of the steam tends to rip-off the caps. 2. High viscosity of the residue makes the caps obsolete. Both the crude oil from the heater (700 ºF) and the flashed vapors from the flash drum (415 ºF) enter the fractionator at the flash zone where vapor-liquid separation occurs. Stripping steam is introduced at the bottom of the fractionator to: - Strip out light materials from the bottom product (atmospheric residue). - Lower the partial pressure and, thus, the boiling point of the residue to avoid thermal cracking and degradation of the bottom product - (Thermal cracking produces coke, which tends to block heat exchangers and other equipment resulting in poor heat transfer and lower efficiencies). The fractionator operates at slightly above atmospheric pressure psia at the fractionator O/H reflux drum psia at the fractionator O/H psia at the flash zone. - There is usually 5 psi pressure drop between the flash zone and the O/H. Reflux (heat removal) is provided though out the tower by; 1. Condensing the tower overhead vapors and returning a portion of the liquid to the top of the tower as a reflux. 2. Withdrawing a liquid stream from some intermediate location in the enriching section of the tower, cooling it, then returning it back to the tower a couple of trays above (this is called pumparound) this provides an internal reflux within the tower. Without pumparounds, a bigger fractionator would be required to achieve the same degree of separation. 3. The over-flash allows some fractionation to occur on the trays just above the flash zone by providing internal reflux in excess of the side stream withdrawal. 4. Pump-back streams lower in the tower (not in Kuwait) Product Draws Kerosene Product is drawn from tray 29 and is introduced to the top tray of Kerosene stripper. Diesel product is drawn from tray 17 and is introduced to the top tray of diesel stripper. Each of the side stream products removed from the tower decrease the amount of liquid traffic below the point of draw-off. Increasing the amount of kerosene withdrawal, for example, will result in a corresponding decrease in the amount of liquid traffic going down through the tower. This will lead to less condensation of vapors going up and more of the diesel cut will rise up and go into kerosene. The opposite is also true. The over flash is a liquid drawn from tray 7 (above the flash zone) and after external flow measurement is returned to tray 6 (below the flash zone). 7-46

47 Petroleum Refining Chapter 7: Distillation The over flash (which is 5 vol. % of the crude) helps to remove the heavy material from the diesel product and improves the fractionation (separation) between the diesel and the atmospheric residue cuts. Figure 7-15 CDU Distillation Tower simplified. Pumparounds Additional reflux is provided by three pumparound systems (top, middle, and bottom pumparounds). Liquid is removed from the tower, (all or portion of the side stream) is cooled by heat exchangers (exchanging heat with the crude oil feed), and returned to the tower. The top pumparound (TPA) is withdrawn from tray 39 and returned after heat exchange to tray 41. The middle pumparound (MPA) is withdrawn from tray 27 and returned after heat exchange to tray 28. The bottom pumparound (BPA) is withdrawn from tray 15 and returned after heat exchange to tray

48 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering This cold stream condenses more of the vapors coming up through the tower thereby increasing the liquid flow (reflux) below that point. Advantages of Pumparound What is the difference between a reflux and a pump-around? Better heat recovery (at higher temperature) and better energy efficiency - If all the heat in the tower were to be removed in overhead condensers, which is not truly possible, all of the heat energy would be exchanged at the bubble-point temperature of the overhead stream (105 ºF). - By using pumparounds at lower points in the column, the heat transfer temperatures are higher (290º, 420º and 580ºF for the top, middle, and bottom pumparounds, respectively) and a higher fraction of the heat energy can be recovered by preheating the feed. - The pumparounds remove all the heat from 2/3 the column. Better fractionation between product cuts - When a portion of the liquid traffic flowing down through the column is removed, cooled, and routed back the column, this cooled stream condenses more of the vapor coming up through the tower (especially the heavy material in the vapor) allowing only the light material to rise up the tower in the vapor phase, thus, fractionation (separation) is enhanced. Better tower design (more proportionate diameter and smaller height) - If all the heat in the tower is to be removed in overhead condensers, the amount of liquid reflux to the tower will be huge. - This would result in an inverted cone-type liquid loading which requires a very large diameter at the top of the tower. - The tower diameter will have to be reduced below each side product draw-off point along the tower to correspond to the decrease in liquid flow to conserve materials of construction and maintain balanced V/L traffic through the tower. - To reduce the top diameter of the tower and even the liquid loading over the length of the tower, intermediate heat-removal streams (pumparounds) must be used. Better vapor-liquid traffic along the tower - Each side stream product withdrawal decreases the amount of internal liquid (reflux) flowing down below that point in the column. To generate uniform liquid-vapor load through the column a portion of liquid is removed, cooled, and routed back the column, this cooled stream condenses more of the vapor coming up the tower and provides the reflux (liquid traffic) below the product draw-off point. 7-48

49 Petroleum Refining Chapter 7: Distillation Figure 7-16 CDU Distillation Column 7-49

50 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Side Draw Strippers The liquid side stream withdrawn from the tower will contain low-boiling components, which lower the flash point of the product. (Because the lighter products in the vapor phase pass through the heavy products in the liquid phase and the two are in equilibrium on every tray). These light ends are stripped from each side stream in a separate side-stripper. In MAB CDU there are two side strippers mounted on each other. The top being for Kerosene (10 bubble-cap trays) and the bottom one being for diesel (4 bubble-cap trays). There function is to remove light ends from the product stream thus improving (increasing) their flash points. For Kerosene, the stripper is provided with a reboiler, whereas for diesel stripping is done by super-heated steam. - This is because Kerosene product should meat ATK specifications which is very stringent on water content to meet freezing point requirement (water tends to freeze at high altitudes). - For diesel, a coalescer is used eventually to remove water from diesel product. Increasing the degree of stripping causes more light ends to be removed thus making the flash of the product stream even higher. The strip-out /light ends (along with the stripping steam in case of diesel) leave the stripper at the top and enters the fractionator in the vapor zone directly above the tray of the side draw. Atmospheric Residue Section Several trays are generally incorporated below the flash zone (6 trays) and stream is introduced below the bottom tray to - Strip any remaining gas oil (more valuable) from the liquid in the flash zone (residue). - Produce high-flash-point bottoms (by stripping the low boiling point material). Stripping Stream Crude towers do not normally use reboilers because of the tendency of the residue to crack at high temperatures clogging the heat exchanger. Superheated steam is used instead to reduce the partial pressure of the hydrocarbon and thus lower the required vaporization temperature. O/H Condensing System The fractionators O/H vapors (usually consisting of gas, unstabilized naphtha and the stripping steam) The pentane and heavier fraction (C5+) is condensed in the overhead cooling system. The butane and lighter (C4-) remain in the gas phase. Ammonia solution is injected to crude column vapor line and on tray 41 to control acids formed by hydrolysis of the salt present (e.g. Na + Cl - + H2O HCl + NaOH). Corrosion inhibitor is also added to prevent corrosion due to HCl and H2S present. Part of the condensate (the naphtha and the water) collected in the O/H accumulator is returned to the top of the tower as reflux, and the remainder is sent to the stabilizer section of the crude unit (or the refinery gas plant). The reflux (being at a lower temperature) controls the temperature of the top of the tower thus the quality (end point) of the O/H product. 7-50

51 Petroleum Refining Chapter 7: Distillation Fractionator process variables and control Pressure The pressure of the fractionator must be kept constant (close to atmospheric pressure) for the stable operation of the column. This done by a pressure controller (at the suction of the compressor) in the O/H system. High fractionators pressure increases the flash zone pressure hence the heater outlet temperature (to get the same separation) Lower pressure, lowers the flash zone temperature and improve fractionation to an extent. O/H Temperature The temperature at the top of the column is controlled by a temperature controller that adjusts the flow of the reflux naphtha. The O/H temperature affects the Naphtha end-point temperature. Naphtha end point If the amount of reflux is increased (or its temperature is decreased; which is not practical since its temperature is already close to ambient temperature), more of the heavy material in the vapor leaving the top of the fractionator will be condensed. The O/H vapor (mainly naphtha) will therefore contain only the light material and upon condensation naphtha will have a lower final boiling point (FBP). To summarize, increasing the fractionator O/H temp increases the FBP of naphtha Heat Removal by TPA For a constant O/H temperature, the heat removal in TPA affects the overhead condenser duty, and vice versa. The heat removal in TPA is controlled by the amount of pumparound flow and the temperature difference between TPA draw-off and return points (according to the relation Q = m Cp T). The return temperature is controlled by a temperature differential controller (TDC) which bypasses part of the TPA around E102 (that heats the crude oil feed). Naphtha / Kerosene Cut Point The draw-off temperature of Kerosene (370 ºF) gives an indication of the cut point between Naphtha and Kerosene. A lower temperature indicates cutting into Naphtha and a higher temperature indicates cutting into Kerosene. The draw-off temperature is maintained by controlling the kerosene draw (flow) and the fractionator overhead temperature. Increasing Kerosene draw will take some of the diesel with it and increase the final BP, while the reducing the O/H temperature will cut into naphtha and increase the amount of liquid to be drawn as kerosene. IBP and FBP of Kerosene The IBP of Kerosene is maintained by a temperature controller that controls the reboiler duty (by bypassing part of the atmospheric residue around the reboiler). The TC is located on the stripper bottoms. The FBP of kerosene is controlled by the amount of heat removal from the MPA (both temperature and flow) and the amount of kerosene withdrawal. 7-51

52 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering - The higher the heat removal by the MPA (e.g. lower return temperature or higher flow rate of the MPA) will provide more internal reflux at a lower temperature thus more cooling. This will condense the heavy compounds in the rising vapor from the diesel section to the kerosene section resulting in a lower FBP of kerosene. - Less kerosene withdrawal will result in more internal reflux (more liquid will flow down in the tower below that draw point). This will result in more condensation of the vapors rising to the kerosene draw-off tray. The heavy molecules will condense first according to their boiling point and only the lighter portion of the vapor will make it to the kerosene draw-off tray. The net result is a lower FBP of kerosene. Kerosene / Diesel cut Point The draw-off temperature of diesel (523 ºF) gives an indication of the cut point between Kerosene and Diesel. This is maintained by controlling the diesel draw-off rate (flow). IBP and FBP of Diesel The IBP of diesel is maintained by controlling the flow rate of the stripping steam to the diesel stripper. The steam lowers the partial pressure of the H/C thus vaporizes the light compounds from diesel. The FBP of diesel is controlled by the amount of heat removal from the BPA (both temperature and flow) and the amount of diesel withdrawal. - The higher the heat removal by the BPA (e.g. lower return temperature or higher flow rate of the BPA) will provide more internal reflux at a lower temperature (more cooling). This will condense the heavy molecules in the rising vapor from the residue section to the diesel section resulting in a lower FBP of diesel. - Less diesel withdrawal will result in more internal reflux (more liquid will flow down in the tower below that draw point). This will result in more condensation of the vapors rising up to the diesel draw-off tray. The heavy molecules will condense first according to their boiling point and only the lighter portion of the vapor will make it to the diesel draw-off tray. The net result is a lower FBP of diesel. If the color of diesel is off spec, you might be cutting into residue. Reduce the amount of diesel draw and/or increase heat removal by BPA to improve fractionation between diesel and residue. Both of these solutions will increase the amount of liquid traffic down the tower, which condenses (and gets rid of) the heavy matter rising up and leaving with diesel. The two variables are usually adjusted by trial and error until desired color is obtained. If neither solution works then there might be an internal problem in the distillation column such as tray valve or tray damage, which reduces fractionation. Diesel / Atmospheric residue cut point Temperature of the over flash indicates the cut point between diesel and atmospheric residue. It is maintained by controlling the diesel product draw-off rate. IBP of atmospheric residue The IBP of residue is maintained by controlling the stripping steam quantity to the fractionator bottom by using a flow controller. The level inside the fractionator is controlled by a level controller that sends a signal to the control value on the atm. Residue product line. 7-52

53 Petroleum Refining Chapter 7: Distillation Summary - CDU Product Quality Control Naphtha Increasing naphtha stabilizer reboiler temp. Removes light gases (heavier) IBP RVP Flash Increasing frac. OH temp. FBP (heavier) API (heavier) Naphtha production Kerosene Increasing kerosene side stripper reboiler temp. Removes light materials (heavier) IBP Flash Increasing Kerosene draw FBP (heavier) API (heavier) Freeze (freeze faster) Diesel Increasing kerosene side stripper reboiler temp. Removes light materials (heavier) IBP Flash Increasing diesel draw FBP (heavier) API (heavier) Color (darker) Residue Increasing stripping steam to frac. bottom FBP does not change Removes light materials (heavier) IBP 7-53

54 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Pump-around All variables kept the same, increasing the pump-around circulation flow rate or T (between draw-off and return) or both, will improve diesel-residue fractionation (separation) for the bottom pumparound (BPA), kerosene-diesel separation for the middle pumparound (MPA), and naphtha-kerosene separation for the top pumparound (TPA). Improving separation means less overlap between the naphtha, kerosene and diesel cuts. Increasing heat removal from fractionator by increasing the flow or T of the TPA will have the same effect as increasing fractionator reflux flow rate or reducing fractionator reflux temperature in reducing the fractionator OH temperature and decrease naphtha FBP as explained above. This will result in less yield of naphtha product (with more yield of kerosene) Increasing MPA will result in more heat removal from that section of the tower and thus increased condensation (of the heavier portion of the rising vapors through that section) resulting in more liquid traffic down that can be drawn with the diesel product. The net result will be less diesel rising up to the kerosene section improving separation between kerosene and diesel as well as more yield of diesel product (with less yield of kerosene product). Effect of Reflux and Pumparround on Product Quality (and BP curve) Increasing fractionator OH temp (by reducing reflux flow, increasing reflux T, reducing TPA flow or T removes light materials from (the lighter portion of) Kerosene. Naphtha Resulting in BP increase in the heavier portion of naphtha boiling-curve. Naphtha yield will increase. Kerosene Resulting in BP increase in the lighter portion of the kerosene boiling-curve. Kerosene yield will decrease. FBP not affected Increasing MPA flow or T removes the heavy materials from (the heavier portion of) Kerosene to (the lighter portion of) diesel. The same effect will be obtained when decreasing kerosene draw Kerosene Resulting in BP decrease in the heavier portion of Kerosene boiling-curve. Kerosene yield will decrease. Diesel Resulting in BP decrease in the lighter portion of the diesel BP curve. Diesel yield will increase. FBP not affected 7-54

55 Petroleum Refining Chapter 7: Distillation Increasing BPA flow or T removes the heavy materials from (the heavier portion of) diesel to (the lighter portion of) residue. The same effect will be obtained when decreasing diesel draw Diesel Resulting in BP decrease in the heavier portion of diesel boiling-curve. Diesel yield will decrease. This will improve the color of diesel if such a problem exists Residue Resulting in BP decrease in the lighter portion of the residue boiling-curve. Residue yield will increase. FBP not affected This is not advisable since its more economical to recover all diesel from residue, while maintaining diesel colorless. To increase the IBP For residue (i.e. from 680 to 690 F), we increase stripping steam flow rate to the fractionator bottom to remove the light materials. IBP residue stripping steam Same principle applies to naphtha, kerosene and diesel therefore IBP diesel stripping steam (in diesel side stripper) IBP kerosene reboiler temperature (in side kerosene stripper) IBP naphtha reboiler temperature (in side naphtha stabilizer), RVP To increase the FBP Residue FBP cannot be changed. FBP Diesel diesel draw-off from main fractionator (this means more heavy materials leading to Color making sample colored instead of colorless) FBP Kerosene kerosene draw-off from main fractionator (this means more heavy materials leading to freezing point say from -55 to -50 C and smoke point both of which is not desirable) FBP Naphtha naphtha draw-off (by fractionator top temperature, say from 230 to 231 F, this means more heavy materials and octane number). Sour Gas Compressor and stabilizer section process variables and control Stabilizer feed surge drum pressure The stabilizer feed drum pressure is maintained by a pressure controller on the sour gas flow from the surge drum. This pressure sets the compressor discharge pressure and should be kept to the minimum required to maintain C3/C4 LPG in the liquid phase later in the stabilizer. Compression results in the condensation of the heavy materials from the gas (light naphtha) which is sent to the stabilizer along with the heavy naphtha. 7-55

56 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Figure 7-17 Compressor and Stabilizer feed drum Sour gas compressor Two compressors are provided (one is a spare). Each is a single stage reciprocating type 1 with capacity control at 0, 25, 50, 75, and 100% load for flexibility during unit turndown 2. Crude unit has a turndown ration of about 60%. However, the compressor is known to operate at 25-75% according to the amount of gas produced, which may vary widely depending on the crude oil feed and whether other streams are added to it (like slops for example). For intermediate flow rates, a spill-back controller is provided which re-circulates part of the compressed and cooled gas to the suction KO (knock out) drum to maintain the pressure. When the spill-back is completely closed and the pressure is not controlled, the excess gases are vented to the flare (emergency) from the compressor suction through a pressure controller. Stabilizer The unstabilized naphtha from the fractionator overhead reflux drum contains propane & butanes (C3 & C4) which makes the vapor pressure (RVP) much higher than is acceptable for gasoline blending. To remove these, wild naphtha is charged to a stabilizing tower (fractionator) where the RVP is adjusted by removing the propane & butanes from the raw naphtha (LSR gasoline) stream. Later, in the product blending section of the refinery, n-butane is added to the gasoline stream to increase the ON and regulate the RVP as desired. In some refineries, the stabilizer is incorporated in the crude distillation unit while in other refineries it is placed in the refinery gas plant. 1 High pressure units use centrifugal type multistage compressors. 2 When the unit is operated at lower throughput 7-56

57 Petroleum Refining Chapter 7: Distillation The feed to the stabilizer consists of heavy naphtha from fractionator reflux drum and light naphtha from stabilizer feed drum. The feed enters the stabilizer in the middle at tray 20 after heating by stabilizer bottoms for heat recovery. Stabilizer feed temperature should be maximized in order to reduce the load on stabilizer reboiler. The stabilizer has 40 trays and is operated at 120 psia to allow the recovery of C3/C4 as LPG at ambient temperature (130 ºF in summer). Stabilizer Reboiler Control Tray 4 temperature is controlled by a temperature controller which bypasses part of the hot BPA steam around the reboiler. This temperature is a measure of the light ends in the stabilized naphtha (indicated by naphtha product RVP). Stabilizer Pressure Control This is done by a pressure controller on stabilizer reflux drum to control sour gas flow. The pressure is better kept low to reduce heat requirement on the reboiler and improve separation, but still high enough to allow recovery of C3/C4 as LPG. Reflux The reflux flow is controlled to give the desired reflux ratio. The higher the reflux the lower the C5 content of LPG. This is because the reflux condenses the C5 vapors rising up to the top of the tower. This requires higher reboiler duty to vaporize the increase in the cooler and higher liquid traffic. 7-57

58 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Figure 7-18 Naphtha stabilizer process variables and control 7-58

59 Petroleum Refining Chapter 7: Distillation 2. Vacuum Rerun unit (VRU) Introduction Vacuum distillation is performed at a pressure lower than atmospheric pressure to take advantage of the fact that reducing the pressure lowers the boiling point of liquids. This permits the distillation of liquids that are temperature sensitive and avoids any degradation of such liquids. Petroleum crude oil is a complex mixture of hundreds of different hydrocarbon compounds generally having from 3 to 60 carbon atoms per molecule (although there may be small amounts of hydrocarbons outside that range with carbon numbers up to 85). The refining of crude oil begins with distilling the incoming crude oil in a distillation process operating at (slightly above) atmospheric pressure. In atmospheric distillation, it is important not to subject the crude oil to temperatures above C ( F) because the high molecular weight components in the crude oil will undergo thermal cracking and form petroleum coke at higher temperatures. The furnace outlet temperature required for atmospheric pressure distillation of the heavy fractions of residual oil are so high that thermal cracking 1 would occur with the resultant 1. Loss of product (giving gases and coke) 2. Equipment fouling (coke deposition) in furnace tubes, distillation tower and the transfer line between them. This constraint yields a residual oil from the bottom of the atmospheric distillation column consisting entirely of hydrocarbons that boil above 370 to 380 C. Lowering the pressure reduces the boiling point of HCs. Therefore, to further distill the residual oil from the atmospheric distillation column, the distillation must be performed at absolute pressures as low as 10 to 40 mmhg (Torr) to limit the operating temperature to less than 370 to 380 C. Figure 7-19 vacuum distillation column in a petroleum refinery. 1 Cracking usually starts at 485 ºF at atmospheric conditions and becomes significant at higher temperatures. 7-59

60 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering Feed and Products The feed to the unit is the atmospheric residue (usually after desulfurization in the ARD unit). The purpose of vacuum unit is to further recover useful distillates from desulfurized atmospheric residue by physical separation. Table 7-7: Vacuum Rerun Units Capacity in Kuwait. Refinery Unit Trains Throughput (BPSD) MAA VRU 1 85,000 MAB VRU 2 150,000 ZOR MAB refinery has two identical and independent vacuum units each capable of processing 75,000 BPSD of desulfurized atmospheric residue. For full utilization of the design capacity of the vacuum units, the atmospheric residue from ARDS and Isomax units are supplemented by imported desulfurized atmospheric residue from Ahmadi Refinery. MAA refinery has one vacuum unit capable of processing 85,000 BPSD of desulfurized atmospheric residue and one bitumen (القار) unit capable of processing 11,000 BPSD. The vacuum unit in Shuaiba refinery is part of the crude unit. There is another vacuum unit part of the H-Oil unit. LVGO and HVGO are drawn as side products. HVGO is used to heat up the feed to make use of its high temperature, and then mixes with the LVGO to make vacuum gas oil (VGO) feed to the Hydrocracker. Table 7-8: Products from MAB vacuum rerun unit (VRU). Product API TBP (⁰F) Wt% Destination 1. Gases 0.3 Burned in the heater 2. Oily water (slops) 1 Crude oil unit 3. VGO (13+35) Hydrocracker unit FCC (Fluid catalytic cracking) 4. Vacuum Residue LSFO via HOC Delayed coker unit Asphalt manufacture Process description The feed to the vacuum unit from the ARD is received at high temperature (500 ºF) in a feed surge drum. It is further heated to 615 ºF before it enters the heater. There it is heated to the flash zone temperature of 770 ºF. The surge drum is blanketed by fuel gas to prevent contamination with air. Furnace outlet temperature ( ºF) is a function of 1. The boiling range of the feed. 2. The fraction vaporized. 7-60

61 Petroleum Refining Chapter 7: Distillation 3. The feed coking characteristics (Conradson carbon). To reduce thermal cracking of the HC and thus coke formation inside the heater tubes, 1. Steam is added to the furnace inlet to increase the furnace tube velocity (reduce residence time) by increasing the volumetric flow rate inside the tubes as well as decreasing the total HC partial pressure in the vacuum tower (Thus lowering the furnace outlet temperature required for separation). 2. A liquid stream (called wash oil) from the chimney tray immediately above the flash zone (in the heavy gas oil range) is recirculated to the heater inlet to increase the furnace tube velocity and reduce residence time and thus cracking. Vacuum Fractionator The vacuum tower has 4 stripping trays (below the feed) and 6 wash trays (above the feed). The tower in addition has a packed bed (2 metal ball rings) to enhance fractionation and condense the heavy vacuum gas oil (HVGO) and the wash oil. Three chimney trays are also provided. For LVGO, HVGO, and heater recycle. These are not intended for fractionation but merely for collecting liquid product for withdrawal. A mist eliminator is also provided to prevent liquid entrainment with the vapor due to excessive vaporization generated by the vacuum. Absolute pressures in the vacuum tower flash zone ranges form mmhg. Steam injection to the furnace inlet and the fractionator bottom lower the effective pressure even further (to 10 mm Hg or less) which further improves vaporization. The amount of stripping stream used is a function of the boiling range of the feed and the fraction vaporized (Generally ranges from 10 to 50 lb/bbl feed). In MAB 0.29 lb steam/gal of bottom product is used. Introducing less steam to the fractionator, will not strip all the gas oil from the vacuum bottoms (vacuum residue). Heat is removed from the tower by three circulating pumparound systems. Top pumparound for light vacuum gas oil (LVGO) to improve fractionation and remove heat from the tower. Middle pumparound for heavy vacuum gas oil (HVGO) to improve fractionation, remove heat from the tower, and heat recovery. In the vacuum tower boot, the residue product is partially recycled to quench (cool) the residue and reduce the residence time at high temperature to avoid cracking. No improvement of fractionation here since the pumparound is at the bottom below the feed, the steam, and the last tray. Strainers are provided for the LVGO and HVGO pumparounds to remove scale which could block the spray system inside the tower. Portion of the HVGO is refluxed back to the tower to provide better fractionation between the HVGO and Vacuum residue in the trays 5-10 and improve HVGO end point. The 10 to 40 mmhg absolute pressure in a vacuum distillation column increases the volume of vapor formed per bbl vaporized. Larger diameters are used in the vacuum distillation columns than the atmospheric columns to control the space velocity of the vapor. Distillation columns may have diameters of 15 meters or more, heights ranging up to about 50 meters, and feed rates ranging up to about 160,000 barrels per day. The vacuum distillation column internals must provide good vapor-liquid contacting while, at the same time, maintaining a very low pressure-increase from the top of the column top to the bottom. Therefore, the vacuum column uses distillation trays only where withdrawing products from the side of the column (referred to as side draws). Most 7-61

62 Prof. Tareq Albahri 2018 Kuwait University Chemical Engineering of the column uses packing material for the vapor-liquid contacting because such packing has a lower pressure drop than distillation trays. This packing material can be either structured sheet metal or randomly dumped packing such as Raschig rings. Figure The internal packing of vacuum distillation column. Vacuum Inducing System The effective pressure 1 at the flash zone determines the fraction of feed vaporized for a given furnace outlet temperature This is essential for the design of the fractionator tower overhead lines, and condenser to minimize the pressure drop between the vacuum-inducing device (the steam jet ejectors) and the flash zone. A few mmhg decrease in pressure drop will save many $$$ in operation cost. The desired operating pressure is maintained by compressing the overhead vapors using steam which is partially condensed in the first stage condenser (in a multistage steam ejectors 2 and barometric condensers system). The size & number of ejectors and condensers used is determined by 1. The vacuum needed 2. The quality of vapors handled For a flash zone pressure of 25 mm Hg 3 ejector stages are usually required. - The first stage condenses the steam and compresses the non-condensable gases. - The second and third stages remove the non-condensable gases from the condensers. The vacuum produced is limited to the vapor pressure of the water used in the condensers. If cooler water is supplied to the condensers, a lower absolute pressure can be obtained in the vacuum tower. 1 The HC's partial pressure = total absolute pressure partial pressure of steam. 2 Vacuum pumps and surface condensers (more expensive but less water contamination). 7-62

63 Petroleum Refining Chapter 7: Distillation Figure 7-21 Steam jet ejector installed Figure 7-22 Steam jet ejector taken apart Example 7-9: Sizing a steam jet ejector system Figure 7-23 Vacuum Distillation Unit Simplified 7-63