RULES OF THUMB FOR PROCESS ENGINEERS Revised 8/2002
Edited by H. R. Hunt
TABLE OF CONTENTS 1
SEPARATIO SEPA RATION N ............. .................... .............. ............. ............. .............. .............. ............. ............. .............. ............. ............. .............. .............. ............. ............. ........1-1 .1-1
A.
Vertical Vert ical Knockout Knoc kout Drum Prelimin Prel iminary ary Sizing Sizi ng ........................... .......................................... ............................. ............................ ............................. ............................. ................ 1-1
B.
Crude Cru de Oil Service Serv ice Separat Sepa rator or Sizing Siz ing ............................. ........................................... ............................ ............................. ............................. ............................ ............................. .................. ... 1-1
C.
Vertic Ver tical al Sep Separa arator tor Design Des ign ............................. ........................................... ............................. ............................. ............................ ............................. ............................. ............................ .................... ...... 1-1
D.
Horizon Hori zontal tal Sep Separa arator tor Design Des ign.............. ............................. ............................. ............................ ............................. ............................. ............................ ............................. ............................. ................ 1-2
E.
Solid/Li Soli d/Liquid quid Separati Sepa rations ons ............................. ........................................... ............................ ............................. ............................. ............................ ............................ ............................. ....................... ........ 1-3
F.
Solid Soli d / Liquid Liqui d Separatio Separ ations ns ............................. ........................................... ............................. ............................. ............................ ............................. ............................. ............................ .................... ...... 1-3
G.
Brown Brow n - Souder Sou derss Equ Equati ation on For Vessel Ves sel Sizing Siz ing ........................... .......................................... ............................. ............................ ............................. ............................. ................ 1-4
H.
Mist Extractor Extr actor Selection Selec tion ........................... .......................................... ............................. ............................ ............................ ............................. ............................. ............................ ......................... ........... 1-4
2 A.
HEAT TRANSFER TRANSF ER SHEL L & TUBE TUB E HEAT EXCHANGERS EXCH ANGERS ....... .......... ....... ....... ....... ....... ...... ....... ....... .....2-1 ..2-1 Heat Exchange Exch angerr Design Desi gn Practice Prac tices.................. s................................. ............................. ............................ ............................. ............................. ............................ ............................. .................. ... 2-1
LIQUIDS
SHELL SHE LL TUBE TUB E SIDE SID E ............................. ........................................... ............................. ............................. ............................ ............................. ............................. ........................... .............2-4 2-4
B.
Reboiler Rebo iler Ther Thermal mal Design Desi gn Practice Prac ticess ............................. ........................................... ............................ ............................. ............................. ............................ ............................. .................. ... 2-8
C.
Proces Pro cesss Condenser Con denser Ther Thermal mal Design Desi gn Practice Prac ticess ............................. ........................................... ............................ ............................. ............................. .........................2-1 ...........2-1 1
D.
Heat Transfer Tran sfer Units Unit s conversio conve rsions ns Table Tabl e 5 ............................ ........................................... ............................. ............................ ............................ ............................. .....................2 ......2-14 -14
E.
Heat Hea t Exchan Exc hanger gerss (Ge (Gener neral)............... al)............................. ............................. ............................. ............................ ............................. ............................. ............................ ............................. ................2 .2 -16
F.
Condense Cond ensers rs ........................... .......................................... ............................. ............................ ............................. ............................. ............................ ............................. ............................. ............................ ..................2 ....2 -16
G.
Reboiler Rebo ilerss and Chillers Chil lers ............................. ........................................... ............................ ............................. ............................. ............................ ............................. ............................. .........................2-1 ...........2-1 6
H.
Sizing Sizi ng Plate Plat e Heat Exchange Exch angers rs ............................ .......................................... ............................ ............................. ............................. ............................ ............................. ............................2-1 .............2-1 7
I.
Brazed Braz ed Aluminum Alumi num Plate Plat e Heat Exchange Exch angers rs ............................ .......................................... ............................ ............................. ............................. ............................ ..................2 ....2 -17
J.
Air Fin Heat Exchangers Excha ngers ............................. ........................................... ............................ ............................. ............................. ............................ ............................ ............................. .....................2 ......2-17 -17
K.
Fired Heaters ............................ .......................................... ............................. ............................. ............................ ............................. ............................. ............................ ............................. ............................2-1 .............2-1 8
L.
Cooling Cooli ng Towers ........................... .......................................... ............................. ............................ ............................. ............................. ............................ ............................. ............................. .........................2-1 ...........2-1 8
M.
Insula Ins ulation tion ............................ .......................................... ............................. ............................. ............................ ............................. ............................. ............................ ............................ ............................. .....................2 ......2-19 -19
N.
NGL Expander Expan der Plants Plant s ............................. ........................................... ............................ ............................. ............................. ............................ ............................. ............................. .........................2-2 ...........2-2 0
1
O.
Miscellaneous Plant Systems ................................................................................................................................2-2 1
P.
Method For Feasibility Study Sizing of Gas Plant Gas/Gas shell & Tube Heat Exchanger: .............2-24
3
TREATING.........................................................................................................................3-1
A.
Dehydration.................................................................................................................................................................3-1
B.
Amine Treating ..........................................................................................................................................................3-2
C.
Mol Sieve Treating .................................................................................................................................................... 3-4
D.
Corrosion ..................................................................................................................................................................... 3-5
E.
Copper Strip................................................................................................................................................................ 3-5
F.
Conversion Factors ................................................................................................................................................... 3-5
G.
Caustic Washer Design ............................................................................................................................................3-5
H.
Metallurgy Requirements For Amine Treaters .................................................................................................3-5
I.
H2S Gas Toxicity ....................................................................................................................................................... 3-7
J.
Liquefied Natural Gas (LNG) Plants ................................................................................................................... 3-7
K.
Gas Treating Iron Sponge ....................................................................................................................................... 3-8
L.
Distribution of Sulfur Compounds in NGL Product ....................................................................................... 3-8
4
FLUID FLOW .....................................................................................................................4-1
A.
Misc. .............................................................................................................................................................................. 4-1
B.
NGL Expander Plants ..............................................................................................................................................4-1
C.
Piping ............................................................................................................................................................................ 4-2
D.
Physical Fan Laws ..................................................................................................................................................... 4-4
E.
Control Valves ............................................................................................................................................................4-5
F.
Two Phase Flow: ........................................................................................................................................................4-6
5
FRACTIONATION.............................................................................................................5-1
A.
Minor Components Non Ideal In Hydrocarbons .............................................................................................. 5-1
B.
Columns........................................................................................................................................................................5-1
6
COMBUSTION ..................................................................................................................6-1
2
A.
Flare............................................................................................................................................................................... 6-1
B.
Fired Heaters .............................................................................................................................................................. 6-1
C.
Fuel Requirements .................................................................................................................................................... 6-2
7
PHYSICAL PROPERTIE S ..............................................................................................7-1
A.
Standard Conditions ................................................................................................................................................. 7-1
B.
Characterization of Liquid Refinery Feeds and Products ............................................................................. 7-1
C.
Physical Properties of Selected Liquids ...............................................................................................................7-2
D.
Physical Properties of Selected Gases/Vapors ...................................................................................................7-8
E.
Physical Properties Recommendations ..............................................................................................................7-1 3
F.
Simulation Techniques for Characterization of Oils .....................................................................................7-1 4
8
COMPRESSORS, EXPANDERS & PUMPS................................................................8-1
A.
Reciprocating Compressors .................................................................................................................................... 8-1
B.
Comp ressor Quickies................................................................................................................................................8-1
C.
Liquefied Natural Gas (LNG) Plants ................................................................................................................... 8-2
D.
Energy Conservation Natural Gas Engines ........................................................................................................ 8-2
E.
Fuel consumption....................................................................................................................................................... 8-2
F.
NGL Expander Plants ..............................................................................................................................................8-2
G.
Gas Processing – Simulation guidelines ...............................................................................................................8-3
H.
Pump sizing ................................................................................................................................................................. 8-3
I.
Pumps ............................................................................................................................................................................ 8-4
J.
General: ........................................................................................................................................................................8-5
9
REFRIGERATION ............................................................................................................9-1
A.
Condensers .................................................................................................................................................................. 9-1
B.
Propane Refrigeration Systems .............................................................................................................................9-1
C.
Gas Processing – Simulation Guidelines .............................................................................................................9-1
D.
Condensing Temperature Effects: ........................................................................................................................ 9-1
3
10 MISCELLANEOUS........................................................................................................ 10-1 A.
Large Production & Processing Platforms .......................................................................................................1 0-1
B.
Water and Steam Systems .....................................................................................................................................1 0-3
C.
Economics ..................................................................................................................................................................10-3
D.
Hydrates .....................................................................................................................................................................10-4
E.
NGL Expander Plants ............................................................................................................................................10 -4
F.
Miscellaneous Plant Systems ................................................................................................................................10 -4
G.
Liquified Natural Gas (LNG) Plants ..................................................................................................................10 -4
H.
Gas Processing – Simulation Guidelines ...........................................................................................................10-4
I.
Offshore Pipeline Gas Specifications ..................................................................................................................10 -4
J.
Offshore Crude Oil Specifications ......................................................................................................................1 0-5
K.
Wind Loadings .........................................................................................................................................................10 -5
L.
Steam Leaks @ 100 psi...........................................................................................................................................10-5
M.
Composition of Air ..................................................................................................................................................1 0-5
N.
Platform Deflection .................................................................................................................................................1 0-5
O.
Kinetics .......................................................................................................................................................................10 -6
P.
Storage, Vessel Capacity ........................................................................................................................................10-6
Q.
Pipeline Volume: ......................................................................................................................................................10-6
R.
Pressure Vessels .......................................................................................................................................................10-6
S.
NACE Requirements .............................................................................................................................................10 -6
T.
Pressure Waves (e.g. water hammer).................................................................................................................10 -6
U.
Insulation Types .......................................................................................................................................................10-7
V.
Absolute Pressure of Atmosphere at Height ‘H’ feet above Sea Level .....................................................1 0-7
4
1 A.
SEPARATION
Vert ical Knockout Drum Preliminary Sizing
1. Size for Vapor W=1100 [Density Vapor (Density Liquid-Density Vapor)] 1/2 Where: W = maximum allowable mass velocity - pounds/hour/ft2 1100 is an empirically determined constant Densities - pounds/cubic foot at process T & P 2. Size for Liquid Should be able to contain maximum slug expected depending on pipe configuration. Never size for less than 1 minute liquid holdup. Size 8 to 10 ft tall B.
Crude Oil Service Separator Sizing
1. Use vertical separator for high vapor to liquid ratios and for two phase separation. 2. Use horizontal separator for high liquid to vapor ratios and for three phase separation. Vessel L/D 3 to 5. 3. Check size for both gas and liquid handling (i.e. gas superficial velocity and liquid residence time). 4. Use 3 min liquid residence time for the hydrocarbon phase in a crude oil system. 5. Use 3 to 6 min residence time for the water phase in a crude oil system. 6. Estimate 15 to 30 minutes water residence time for electrostatic coalescers (100% filled). Vessel L/D 4 to 6. C.
Vertical Separator Design
1. The disengaging space - the distance between the bottom of the mist elimination pad and the inlet nozzle, should be equal to the vessel internal diameter or a minimum of 3' -0". 2. The distance between the inlet nozzle and the maximum liquid level should be equal to one-half the vessel diameter, or a minimum of 2' -0". 3. A mist eliminator pad should be installed. Otherwise the separator should be designed so that actual gas velocity should be no greater than 15% of the maximum allowable gas velocity as calculated by the following equation: Vg = (11.574) * (AMMcfd) /A
1-1
Where: Vg = Vertical velocity of gas, ft/sec AMMcfd = Actual gas volume at operating conditions, MMcfd A = Cross Sectional Area of vessel, ft2 4. The dimension between the top tangent line of the separator and the bottom of the mist eliminator pad should be a minimum of 1' -0". 5. Inlets should have an internal arrangement to divert flow downward. 6. Liquid outlets should have antivortex baffles. 7. Mist eliminator pads should be specified as a minimum of 4 inches thick, nominal 9 lb/ft3 density and stainless steel. 8. Normal practice for calculating liquid retention time is to allow for the volume contained in the shell portion of the vessel only. No credit is taken for any liquid retention time attributable to the volume contained in the vessel head. Sump height should be a minimum of 1' -6". to allow for liquid level control. D.
Horizontal Separator Design
The following are commonly used rules of thumb for s izing Horizontal separators. 1. Oil level is usually controlled by a weir, which is commonly placed at a point corresponding to 15% of the tangent-to-tangent length of the vessel. This results in 85% of the vessel being available for separation. Height of the weir is commonly set at 50% of the internal diameter of the vessel. 2. The maximum liquid level should provide a minimum vapor space height of 1' -3" 3. but not be substantially below the centerline of the vessel. 4. Separators designed for gas-oil-water separation should provide residence time and separation facilities for removal of the water. 5. For separators handling fluids where foaming is considered a possibility, additional foam disengagement space and foam control baffling should be provided. Mist eliminator devices should be located external to the vessel to maximize foam disengagement potential within the vessel. 6. The volume of dished heads should not be taken into account in vessel sizing calculations. 7. Inlet and outlet nozzles should be located as closely as possible to vessel tangent lines. 8. Liquid outlets should have antivortex baffles.
1-2
E.
Solid/Liquid Separations
Recommended Feed Solids Content for Separation Processes Feed Solids Content in Vol % Max 25 Max 4 Max 10 Max 0.2 Max 20 Max 15 15-20 40 10 Max 5/10 Max 25 Max 40 25-40 30-50 15-40 10-20 10-30 20-40 20-40 20-40 20-40 2-30 10-40 20-40
Decanter centrifuge Self-cleani ng separator Disc-nozzle centrifuge Tube centrifuge Conical hydrocyclone Circulating bed hydrocyclone Tables and spirals Cone concentrato r Heavy media cyclone, jigs Clarifiers/thickeners Hydrosep arators Bowl classifiers Upstream classifiers Rake/spiral classifiers Filter press Vacuum disc filter Vacuum drum Vacuum band Horizontal filters Sieve bends Vibro screen Basket/peeler centrifuge Pusher centrifuges Screen (scroll) centrifuges Vibro screen centrifuges
F.
20-40
Solid / Liquid Separations Comparative Performance
UNIT OPERATION
FILTRATION SEDIMENTATION
PRODUCT PARAMETER SOLID IN LIQUID IN WASH LIQUID SOLIDS POSSIBILITIES STREAM STREAM FAIR TO GOOD GOOD GOOD
CENTRIFUGATION
FAIR TO EXCELLENT FAIR
POOR POOR
CYCLONING
POOR
POOR
SCREENING
POOR
ULTRAFILTRATION
EXCELLENT
POOR TO FAIR POOR TO FAIR
FAVORABLE FEED CONDITIONS SOLIDS SOLIDS CONCENTRATION CHARACTERISTICS HIGH/MEDIUM
LIGHT, COURSE TO MED. FLOC. FINE
MEDIUM/LOW
DENSE, MEDIUM OR FLOCCULATED FINE DENSE, FINE
LOW EFFICIENCY FAIR TO EXCELLENT POOR
LOW/MEDIUM
POOR
HIGH/MEDIUM
POOR
LOW
1-3
MEDIUM/LOW
DENSE, COURSE TO MEDIUM COURSE TO MEDIUM VERY FINE
G.
Brown - Souders Equation For Vessel Sizing
W = C DV ( D l − DV )
W = Vapor Loading - #/Hr. /Sq. Ft. DV = Vapor Density - #Cu. Ft. at operating condition D1 = Liquid Density - #Cu. Ft. C = Constant (a) For Absorbers use 600 (b)For Scrubbers use 1100 (c)For Still use 500 Example: Size Scrubber For Field Engine Discharge Dv
=
29, 423# / mol 63 psia 520 R x x 380.6 ft 3 / mol 14.5 psia 550 R
= 0.333# / ft 3
D1 = 0.82 Sp. Gr. X 62.3 #/Cu. Ft. (H2 0) = 51.0 #/Cu. Ft. W = 1100 .333(51.0 − .333) = 1100v16.9 = 4520 #/hr./Sq. Ft.
Gas Flow = 158,311 MPD x 29.423 #/Mol = 194,000 #/hr. Cross Section Area Required = 194,000 #/hr. 4520 #/hr/Sq. Ft Dia. =
43 x 4 / π
= 43 Sq. Ft.
= 7.4 Ft .
Use 8 Ft. Diameter Scrubber H.
Mist Extractor Selection
1. The stainless steel mesh pad type mist extractor is generally less expensive than the vane type and is adequate for most clean service applications. Similar liquid removal efficiencies can be achieved (within certain velocity constraints) with mist particle sizes of 10 microns and larger. 2. The pad type usually has less clean pressure drop than the vane type. 3. The vane type usually performs better than the pad type where tacky solids such as iron sulfide are present in the flowing gas stream. The liquid flow from the mist extractor is at right angles to the gas flow in vane type and it tends to wash solids away better.
1-4
4. If the vane type is used in corrosive service (hydrogen sulfide, carbon dioxide, or oxygen with water wet gas), the vanes should be 316 stainless steel. Experience has shown that a small amount of corrosion with carbon steel vanes roughs the surface and solids tend to accumulate and plug the vanes rapidly. 5. For retrofit or sometimes new applications, it is possible to use a smaller dia meter vessel for the vane type as it may be fitted in different orientations to limit the velocity to acceptable ranges. The pad type is usually installed horizontally. 6. It is usually cheaper to retrofit vessels with the pad type as both would have to be cut and match marked to fit through an 18” or smaller manway and reinstalled inside the vessel. The vane type usually has boxing that must be welded together inside the vessel while the pad type can usually be bolted. 7. The vane type may be used for small in-line applications where the pad type usually can not. 8. If the pad type plugs with solids or hydrates, the pressure drop will likely dislodge the mist extractor and plug downstream piping or equipment. 9. For tough separation applications where it is necessary to remove mist particles smaller than 10 microns (such as inlet to glycol or amine systems where the foreign liquid may cause foaming or chemical contamination), often a combination of pad type (for coalescing) and vane type (for mist removal) is used.
1-5
Separation References
GPSA Engineering Data Book, Vol. 1, Section 7 - Separators and Filters Engineering Standard 10.48-2: Process Vessel Sizing- Entrainment Reduction SPS Design Report - DR15 Selection of Equipment for a Solid-Liquid Separation Process
1-6
2
HEAT TRANSFER SHELL & TUBE HEAT EXCHANGERS
Material presented herein is intended to supplement Phillips Engineering Standards, General Design Specifications, and Recommended Design Practices listed in the Section VI and does not supersede these documents. A.
Heat Exchanger Design Practices
1. Heat Exchanger Selection The selection logic shown in Fig. 1, at the end of this section, may be used as a guide in selecting heat exchanger types. 2. General Design Practices General design practices governing the design of shell-and- tube heat exchangers are summarized, as follows:
• • • • • • • • • • • • • • • • • •
High pressure stream should be located on the tubeside. Stream requiring special metallurgy should be located on the tubeside. Stream exhibiting highest fouling should be located on the tubeside. More viscous fluid should be located on the shellside. Lower flowrate stream should be placed on the shellside. Consider finned tubes when shellside h is less than 30% of tubeside h. Do not use finned tubes when shellside fouling is high. Design exchanger for maximum utilization of allowable pressure drop. Do not design heat exchanger for operation in transition flow. Do not provide thermal overdesign by increasing fouling factors. Provide thermal overdesign by increasing bundle length, not diameter Avoid multiple tubepass exchangers with close temperature approaches. Vertical shellside condensation should be in downflow. Vertical tubeside boiling should be in upflow. Use RODbaffle exchangers when tube vibrations are predicted. Use RODbaffle exchangers for low shellside pressure loss processes. Avoid triple-segmental plate baffles, disk-and-doughnut baffles, and orifice baffles. Horizontal shellside condensers should be specified with vertically cut baffles.
3. Overall Heat and Film Transfe r Coefficients Overall heat transfer coefficients suitable for feasibility design estimates are provided in Table 1, and film coefficients are contained in Table 2. 4. Heat Exchanger Velocities
2-1
Recommended shellside and tubeside liquid velocities for various tube materials are summarized as follows: Permissible tubeside velocities for dry gases range from 50 to 150 feet/sec. The recommended minimum shellside liquid velocities are 1.5 feet/sec. TUBE Material
Velocity (FT/SEC)
ADMIRALTY, CARBON STEEL COPPER, BRASS (85-15) NICKEL, COPPER-NICKEL STAINLESS STEEL, MONEL TITANIUM
4 TO 8 2 TO 4 5 TO 10 6 TO 12 6 TO 15
5. Allowable Pressure Losses Recommended maximum allowable shellside and tubeside pressure losses are 10 to 15 psi for plate-baffle exchangers. Allowable shellside pressure losses for RODbaffle should range from 4 to 8 psi. 6. Cooling Water Temperatures Maximum cooling water and tube wall temperatures to minimize fouling deposition are 125F and 145F, respectively. 7. Mean Temperature Differences Log Mean Temperature Difference (LMTD) correction factors (F) for single shellpass, multiple tubepass exchangers should be greater than 0.75 to avoid temperature approach problems. 8. Recommended Fouling Factors Recommended Tubular Exchanger Manufacturers Association (TEMA) fouling factors are provided in Table 3. 9. Vibration Considerations
•
Shellside baffle tip and average crossflow velocities should not exceed 80% of the calculated Connors Critical Velocity in order to avoid fluidelastic instability tube vibration.
•
The vortex shedding-to-tube natural frequency ration should not exceed 0.50 to avoid vortex shedding tube vibration.
•
In exchangers flashing gas on the shellside, the turbulent buffeting- to-tube natural frequency ratio should not exceed 0.50 to avoid turbulent buffeting tube vibration.
2-2
•
In exchangers where acoustic resonance (noise) is predicted, a triangular tube pitch may eliminate the problem Detuning plates may also be necessary in certain cases.
TABLE 1
HOT FLUID
COLD FLUID
WATER AMMONIA MEA OR DEA FUEL OIL FUEL OIL GASOLINE HEAVY OIL HEAVY OIL REFORMER STREAM LIGHT ORGANICS MEDIUM ORGANICS HEAVY ORGANICS GAS OIL GASES GASES CONDENSING STEAM CONDENSING STEAM CONDENSING STEAM CONDENSING STEAM CONDENSING STEAM STEAM LIGHT ORGANICS MEDIUM ORGANICS HEAVY ORGANICS CRUDE OIL GASOLINE (CONDENSING)
WATER WATER WATER WATER OIL WATER WATER HEAVY OIL REFORMER STREAM WATER WATER WATER WATER WATER GASES WATER LIGHT ORGANICS MEDIUM ORGANICS HEAVY ORGANICS PROPANE (BOILING) GASES LIGHT ORGANICS MEDIUM ORGANICS HEAVY ORGANICS GAS OIL CRUDE OIL
2-3
OVERALL U BTU/HR-FT2-F
250-500 250-500 140-200 15-25 10-15 60-100 15-50 10-40 50-120 75-120 50-125 5-75 25-70 2-50 2-25 200-700 100-200 50-100 6-60 200-300 5-50 40-75 20-60 10-40 80-90 20-30
TABLE 2. APPROXIMATE FILM HEAT TRANSFER COEFFICIENTS
LIQUIDS Oils, 20º API 200º F average temperature 300º F average temperature 400º average temperature Oils, 30º API 150 º F average temperature 200º F average temperature 300º F average temperature 400º F average temperature Oils, 40º API 150º F average temperature 200º F average temperature 300º F average temperature 400º F average temperature Heavy Oils, 8-14º API 300º F average temperature 400º F average temperature Diesel oil Kerosene Heavy naphtha Light naphtha180 Gasoline Light hydrocarbons Alcohols, most organic solvents Water, ammonia Brine, 75% water VAPORS Light hydrocarbons Medium HCs, organic sol. Light inorganic vapors Air Ammonia Steam Hydrogen — 100% Hydrogen — 75% (by volume) Hydrogen — 50% (by volume) Hydrogen — 25% (by volume)
SHELL
TUBE SIDE
40-50 70-85 80-100
15-25 20-35 65-75
70-85 80-100 110-130 130-155
20-35 50-60 95-115 120-140
80-100 120-140 150-170 180-200
50-60 115-135 140-160 175-195
20-30 0- 50 115-130 145-155 145-155 180 200 250 200 700 500
10-20 20-30 95-115 140-150 130-140 200 250 200 700 500
Shell or tube sides 10 psig 50 psig 100 psig 300 psig 25 60 100 170 25 70 105 180 14 30 60 100 13 25 50 85 14 30 55 95 15 30 50 90 40 105 190 350 35 80 150 280 30 70 130 240 25 55 100 180
VAPORS CONDENSING Steam Steam, 10% non-condensables Steam, 20% non-condensables Steam, 40% non-condensables Pure light hydrocarbons Mixed light hydrocarbons Gasoline Gasoline-steam mixtures Medium hydrocarbons Medium hydrocarbons with steam Pure organic solvents Ammonia
Shell or tube sides 1,500 600 400 220 250-300 175-250 150-220 200 100 125 250 600
LIQUIDS BOILING
Water Water solutions, 50% water or more Light hydrocarbons Medium hydrocarbons Freon Ammonia Propane Butane Amines, alcohols Glycols Benzene, tolune
1,500 600 300 200 400 700 400 400 300 200 200
2-4
500 psig 200 220 120 100 110 135 420 340 310 270
NOTES: 1. Where a range of coefficients is given for liquids, the lower values are for cooling and the higher are for heating. Coefficients in cooling, particularly, can vary considerably depending upon actual tube wall temperature. 2. Tube side coefficients are based on 3 /4 –in diameter tubes. Adjustment to other diameters may be made by multiplying by 0.75/actual outside diameter. Shell side coefficients are also based upon 3/4 –in diameter. Precise calculations would require adjustment to other diameters. The accuracy of the procedure does not warrant it. 3. Coefficients can vary widely under any one or combination of the following: a. Low allowable pressure drop. b. Low pressure condensing applications, particularly where condensation is not isothermal. c. Cooling of viscous fluids particularly with high coefficient coolants and large LMTDs. d. Condensing with wide condensing temperature ranges — 100º F and larger. e. Boiling, where light vapor is generated from viscous fluid. f. Conditions where the relative flow quantities on shell and tube sides are vastly different (usually evidenced by difference in temperature rise or fall on shell and tube sides). g. Wide temperature ranges with liquids (may be partly in streamline flow).
TABLE 3 EXCHANGER SERVICE
FOULING (HR-FT2-F/BTU) LESS 125 F GREATER 125F COOLING TOWER WATER 0.001 0.002 BRACKISH WATER 0.002 0.003 SEA WATER 0.0005 0.001 BOILER FEEDWATER 0.001 CONDENSATE 0.0005 STEAM 0.0005 COMPRESSED AIR 0.001 NATURAL GAS & LPG GAS 0.001 - 0.002 ACID GASES 0.002 - 0.003 REFORMER FEED-EFFLUENT GAS 0.0015 HYDROCRACKER FEED-EFFLUENT GAS 0.002 HDS FEED-EFFLUENT GAS 0.002 MEA AND DEA SOLUTIONS 0.002 DEG AND TEG SOLUTIONS 0.002 HEAT TRANSFER FLUIDS 0.002 PROPANE AND BUTANE 0.001 GASOLINE 0.002 KEROSENE, NAPTHA, & LIGHT DISTILLATES 0.002 - 0.003 LIGHT GAS OIL 0.002 - 0.003 HEAVY GAS OIL 0.003 - 0.005 HEAVY FUEL OIL 0.005 - 0.007 VACUUM TOWER BOTTOMS 0.010 NATURAL GAS COMBUSTION PRODUCTS 0.005
9. TEMA Shell Configurations Single shellpass, TEMA “E” shells are preferred for most single-phase and condensing applications. Two shellpass, TEMA “F” shells with two tubepass bundles are preferred when pure counterflow conditions and maximum mean temperature difference (MTD) are required. “F” shell exchangers should be specified with welded longbaffles or “ Lamiflex” longbaffles seals. Bundle should may also be utilized to minimize longbaffle leakage. TEMA “G” and “H” split-flow shells are preferred only for horizontal shellside thermosiphon reboilers. Dividedflow TEMA “J” shells with RODbaffle tube bundles are preferred for low pressuredrop, single-phase and condensing services. TEMA “K” shells are used exclusively for horizontal, kettle reboilers. 10. Return Head Types Fixed tubesheet exchangers are preferred for services where thermal expansion, shellside mechanical cleaning, and tube bundle removal are not concerns. U-tube
2-5
and floating head bundles are required when thermal expansion, shellside mechanical cleaning, and tube bundle removal provisions must be made. Fixed tubesheet exchangers should be considered if shellside-to-tubeside inlet temperature differences are less than 100F. Fixed tubesheet exchangers having shell expansion joints should be avoided. U-tube and floating head exchangers are required when fixed tubesheet units cannot meet above requirements, with U-tube bundles being preferred over floating head bundles if tubeside mechanical cleaning is not required. Split-ring floating head bundles are preferred over pull-through floating head bundles in general refinery service because of higher thermal performance and lower cost. Outside packed floating head exchangers are not recommended. 11. Shellside Baffle Types Baffles types recommended for Phillip’s plant services include single-segmental plate-baffles, double-segmental plate-baffles, no-tube-in-window(NTIW) baffles, and RODbaffles. Single-segmental plate baffles, having a single chordal cut, are preferred for single-phase services where higher shellside pressure losses (15 psi) may be tolerated. Double-segmental plate baffles, having two chordal cuts, are preferred for single-phase and condensing services, where modest shellside pressure losses (10 psi) are allowed. RODbaffles are preferred for single-phase and two phase services, where low shellside pressure losses (5 psi) are required or where flow-included tube vibrations are likely in plate-baffle exchangers. Triple segmental disk-and-doughnut, and orifice baffles are not recommended. NITW baffles may be used as an alternate to RODbaffles where economics are favorable. 12. Tube Type, Size, and Layout The preferred tube size for shell-and-tube heat exchangers in medium to heavy tubeside fouling service (.001 hr-ft2 -F/Btu or greater) is 1.00 inch O.D. For light tubeside fouling services (less than .001 hr - ft2 -F/Btu), 0.750 inch O.D. tubes are preferred. Generally 30 or 60 degree triangular layouts are preferred for clean, single-phase services (<.001 hr- ft2 -F/Btu) in which chemical cleaning maybe used. For medium or heavy fouling services (> .001 hr-ft2 -F/Btu) in which mechanical cleaning is required, 90 square or 45 rotated square layouts are preferred. Minimum TEMA tube pitch-to-diameter ratio is 1.25. For kettle and internal reboiler services and all RODbaffle exchangers, 90 square layout is required. 13. Recommended Material Tubes: Inhibited Admiralty tubes are strongly recommended for non-chromate containing, cooling water services where tubewall temperatures range from 145F to 450F. Inhibited Admiralty tubes are also recommended for conventionally treated cooling water service for tubewall temperature between 165F and 450F. Do not use admiralty or other copper bearing alloys when cooling tower water may become contaminated with ammonia or where copper is incompatible with the process fluid. Carbon steel tubes are recommended for cooling water services where tubewall temperature is below 165F. Low-chrome steel tubes are recommended for hightemperature, sulfur-bearing streams. Austenitic stainless steel alloys are
2-6
recommended for low temperature services (below - 150F). Monel tubes are recommended for HF acid-containing streams above 160F, while Titanium tubes are recommended for brackish and sea water services. Welded, fully killed carbon steel (ASTM A-214) should be avoided in low pH water soluble hydrocarbons, furfural, phenol, sulfuric acid, amine service, HF alkylation, and in final overhead crude tower coolers. Seamless carbon steel tubes (A-179) should be used where welded tubes are not permitted. Duplex 2205 tubes should be used instead of austenitic stainless tubes in high chloride services. The table below contains recommended tube wall thicknesses.
Material
¾ Inch OD
Wall Thickness
1 Inch OD
Wall Thickness
Carbon Steel
14 BWG avg wall
.083
12 BWG avg wall
.109
Non-Ferrous (Inhibited Admiralty)
16 BWG min wall
.065
14 BWG min wall
.083
Nickel Base Alloy
16 BWG avg wall
.065
14 BWG avg wall
.083
Ferrous Alloy Steel
16 BWG avg wall
.065
14 BWG avg wall
.083
Baffles, Tie Rods, & Spacers : should be constructed of minimum quality material compatible with tube and tubesheet material. Tube Sheets : must be compatible with service conditions. In services requiring welded tube-to-tubesheet joints, strength welds a re preferred over seal welds. Shell & Channels : must be compatible with service conditions. Specify TEMA “A” type heads when access to the tube ends is desirable or when frequent tubeside cleaning is expected.
Direct question about material suitability should be directed to Engineering Materials and Services. 14. U-Bend Support U-tube exchangers having bundle diameters greater than 36 inches should have U bend tube supports. I designing a new U-tube exchanger, it is preferred to specify a full support baffle at the U-bend tangent, and avoid flowing through the U-bend entirely. 15. Nozzles, Impingement Plates, and Annular distributors Momentum criteria (pv2 ) above which shellside impingement plates and annular distributors and tubeside solid distributor plates should be used are summarized in Table 4. Impingement rods can be utilized in lieu of a solid impingement plate. Rod
2-7
diameter should be identical to the tube O.D. Perforated impingement plates should not be used. B.
Reboiler Thermal Design Practices
1. Reboiler Selection Logic The choice of reboiler type is governed by thermal performance, fluid properties, fouling tendencies, and surface area requirements, as shown in the logic diagram provided in Fig. 2 at the end of this section. 2. Internal or Column Reboilers Internal reboilers, consisting of multi- tubepass, U-tube bundles, should be used for relatively-clean, moderate-viscosity fluids, in small surface-area applications, where periodic column shutdown for cleaning may be tolerated.
• • • • • •
Internal Reboiler Recommended Design Practices Tube Bundle shall be U-Tube Type Tubes Shall be Oriented on 90 Degree Square Pitch Minimum Clearance Between Tubes Shall be 0.25 inches Tube Bundle Diameter Shall Not Exceed 36 Inches Use Two Bundles Side-by-Side in Column for Large Area Requirements Limit Design Heat Flux < 0.7 Maximum Heat Flux
2-8
NOZZLES
TABLE 4 FLUID
MAXIMUM (Pv 2) (LB/FT 2_SEC2)_
SHELLSIDE NOZZLES
CLEAN, NONCORROSIVE, NONABRASIVE SINGLEPHASE GAS, VAPOR, LIQUID
SHELLSIDE NOZZLES
ALL OTHER LIQUIDS
SHELLSIDE NOZZLES
TWO-PHASE MIXTURES, SATURATED VAPORS, ALL OTHER GASES AND VAPORS
IMPINGEMENT PLATE OR ANNULAR DISTRIBUTOR
BUNDLE/SHELL ENTRANCE & EXIT
ALL FLUIDS
4000
TUBESIDE
TUBESIDE
CLEAN, NONCORROSIVE NON-ABRASIVE LIQUIDS TWO-PHASE MIXTURES, SATURATED VAPORS, ALL OTHER GASES AND VAPORS
2-9
1500
500
6000
AXIAL NOZZLES WITH PERFORATED DISTRIBUTOR PLATES
3. Kettle Reboilers Kettle reboilers consisting of multiple tubepass, U-tube bundles, installed inside enlarged TEMA K type shells, are preferred for medium viscosity fluids in moderately heavy fouling services, where large surface areas are required.
• •
• • • • • • • • • •
Kettle Reboiler Recommended Design Practices Tubes Shall be on 90 Degree Square Pitch Tube Pitch Depends on Temperature Difference Temperature Tube Pitch (inch) Difference ¾” O.D. Tube 1” O.D. Tube <35F 1.000 1.250 <60F 1.125 1.375 >60F 1.250 1.500 Shell Diameter (D s) > 1.6 Bundle Diameter (D b) Kettle diameter should be sized for desired liquid entrainment ratio Column Liquid Height (Hd) > Bundle Diameter, (D b) Weir height (Wh ) > Bundle Diameter (D b) Use Two Feed & Return Lines for Boiling Range ∆ T br >100F Use Two Feed & Return Lines for Bundle lengths>12 feet Limit Design Heat Flux < 0.7 Maximum Heat Flux Limit Mixture Wall-to-Bulk Fluid ∆ Twb < Half Boiling Range ∆ T br Use RODbaffle Bundles if Tube Vibration Likely Use Small Diameter, Long Tube Length Bundle when practical
4. Vertical Thermosiphon Reboilers Vertical thermosiphon reboilers, consisting of single-tubepass, single-shellpass, TEMA E shells and having upflow boiling on the tubeside, should be used for moderate-fouling low viscosity (M< 50 cP), Wide boiling- range ( ∆ T br > 100F) mixtures, at above atmospheric pressures, where moderate surface areas are required. Vertical Thermosiphon Reboiler Recommended Design Practices • Single-Tubepass, Single-Shellpass, Fixed Tubepass Exchanger • Design Exit Weight Fractions Vapor range from 0.10 to 0.15 for Hydrocarbons • Maximum Exit Weight Fraction Vapor less than 0.30 for Hydrocarbons • Suited for Wide Boiling Range ( ∆ T b > 100F), low Viscosity (M< 50cP) Fluids • Liquid Driving Head (Hd) = 60 to 100% of Tube Length (Lt ) • Liquid Sensible Heating Zone Length (Lsh) < 25% of Tube Length (Lt ) • Exit Pipe Flow area (A po ) ~ (Total Tubeside flow area (At ) • Inlet Pipe Flow area (A pi) – 25% of total tubeside Flow Area (At ) • Exit Line Pressure Drop ( ∆ Po) equal to 30% of total Hydrostatic Head ( ∆ Ph) • Use Sweep and Long Radius Elbows in Two-Phase Exit Lines
2-10
• •
Limit Design Heat Flux < 70% of Maximum Nucleate Boiling Heat Flux Pr < 0.2 Consider inlet tubeside distribution baffles for cases where two-phase process streams enter exchanger
5. Horizontal Thermosiphon Reboilers Horizontal thermosiphon reboilers, consisting of multiple-tubepass, U-tube or floating- head RODbaffle bundles, in either TEMA J,G or H shells, should be considered for viscous fluids, in moderate fouling service, where larger surface areas are required. Horizontal Thermosiphon Reboiler Recommended Design Practices • Multiple-Tubepass, U-Tube or Floating Head Design • Use TEMA H Shell Configuration for Tube Length (Lt )>12 feet • Column Liquid Driving Heat (H d) > Bundle Diameter (D b) • Design Exit Weight Fraction Vapor 0.10 to 0.20 for Hydrocarbons • Maximum Exit Weight Fraction vapor <0.30 for Hydrocarbons • Limit Design Heat Flux < 0.7 Maximum Nucleate Boiling Heat Flux • Tube Diameter-to-Pitch (Dt /Pt ) Ratios same as Kettle Reboilers • Exit Line Pressure Drop ( ∆ Po) < 0.3 Total Hydrostatic Head ( ∆ Ph) • Use Sweeps and Long Radius Elbows in Two-Phase Exit Piping 5. Forced Circulation Reboilers Forced circulation reboilers, having vaporization on the tubeside, are recommended for highly viscous fluids in heavy fouling service, where large surface areas and low exit weight fraction vapor are required. Forced Circulation Reboiler Recommended Design Practices • Used for Highly Viscous, Heavy Fouling Fluid Services • Vertical Tubeside Vaporization Preferred • Entering Liquid Velocities of 5 to 7 fps • Design Exit Weight Fraction Vapor from 0.05 to 0.10 • Bubble Flow, Two-Phase Flow Regime Preferred • Avoid Slug and Stratified Flow in Horizontal Tubeside Reboilers C.
Process Condenser Thermal Design Practices
The major choices to be made in the selection of shell-and-tube condensers used in the petrochemical industry are between shellside and tubeside condensation and between horizontal and vertical orientation. 1. Condenser Selection Logic Process condenser selection should be go verned by thermal performance, allowable pressure loses, operational pressure, condensing temperature range, condensing medium corrosiveness, mechanical cleaning considerations, and integral condensate
2-11
subcooling requirements, as shown in the logic diagram Fig. 3 at the end of this section. 2. Horizontal Shellside condensers Horizontal TEMA E shellside condensers are preferred for noncorrosive, low pressure and vacuum service, where the single-phase tubeside cooling medium must be placed in on the tubeside because of high fouling deposition. In low shellside pressure loss services, TEMA J Shell, divided-flow, condensers containing RODbaffle bundles are preferred. TEMA G shell condensers may be considered in cases where temperature pinch problems occur in E or J type shell. H and K type shells are not recommended for use in horizontal condensers..
• • • • • • • •
• • • • •
Horizontal Shellside Condenser Recommended Design Practices Use for Noncorrosive, Low Pressure and Vacuum Service Use when Frequent Mechanical cleaning is required Limit Shellside Pressure Loss to less than 10% of Inlet Pressure Evaluate Effects of Shellside ∆ p on MTD for Mixtures Use TEMA E Shell if Shellside ∆ p Permits Use Multiple “E” Shells in Series if Temperature Pinch Predicated Use TEMA J Shell with RODbaffle Bundles for Low Pressure Service Shell and Baffle Type Governed by Following Shellside ∆ p : §
Single Segmental E Shell:
§
Double Segmental E Shell:
§
RODbaffle J Shell:
∆ p s = 10-20 psi ∆ p s = 5-10 psi ∆ p s = 1-5 psi
Design condenser for Shear Flow Regime Use Vertical Baffle Cuts With Drainage Notches Vary Baffle Spacing at Exit To Achieve Shear Flow Design Outlet Nozzles to Avoid Shellside Condensate Flooding Use Separate Exchanger for Condensate Subcooling
3. Vertical Shellside Condensers Vertical TEMA E Shell downflow shellside condensers are preferred for noncorrosive, low-to-moderate pressure services, where two-phase upflow boiling is occurring on the tubeside.
• • • • • •
Vertical Shellside Condenser Recommended Design Practices Use for Noncorrosive, Low-to-Moderate Pressure Services Use with Two-Phase Boiling on Tubeside Use for Close Approach Temperatures Preferred for Vertical Thermosiphon Reboilers and Feed-Effluent Exchangers Design for Downflow Condensation and Upflow Boiling Limited to Single Shellpass, TEMA E Shells
2-12
•
Double Segmental Plate Baffles and RODbaffles Preferred
4. Horizontal Tubeside Condensers Horizontal tubeside condensers are preferred for kettle and horizontal shellside thermosiphon reboilers and in corrosive services where the condensing medium requires special metallurgy. Multiple tubepass tubeside condensers should be avoided because of potential liquid-dropout and inerts accumulation in floating head channels. Horizontal Tubeside condenser Recommended Design Practices • Use in Kettle and horizontal Thermosiphon Reboiler service • Preferred when Condensing Medium Requires Special Metallurgy • Limited to Single Tubepass and Two Tubepass U-Tube Designs • Do Not Design as Multiple Tubepass, Floating Head Unit • Design for Operation in Shear-Controlled Flow Regime §
5. Vertical Tubeside Condensers Vertical TEMA E shell downflow tubeside condensers are preferred for high pressure, wide condensing range, corrosive fluids, where single-phase fluids are used as cooling medium. Vertical Tubeside Condenser Recommended Design Practices • Preferred for High Pressure, Corrosive Condensing Media • Use with Single-Phase Cooling Medium • Design as Single Tubepass TEMA E Shell Configuration • Consider when Inerts Removal is Critical • Consider for Wide Condensing- Range Fluids
2-13
D.
Heat Transfer Units conversions MULTIPLY
LENGTH
kj/kg C
0.23885
Btu/lb F
m
3.2808
ft
HEAT
kcal/kg kcal/kg C
1.0000
Btu/lb F
in
0.083333
ft
Btu/lg F
4.1868 4.186 8
kj/kg C
in
25.400
mm
THERMAL
W/m C
0.57779
Btu/hr ft F
ft
12.000
in
CONDUCTIVITY CONDUCTIVITY cals/s cm C
241.91
Btu/hr ft F
0.30480
m
2
m
10.764 0.15500
in
0.0069444
2
m /m 2
ft /ft ft
0.092903
3
m
In
3.2808 0.30480
2
35.315
3
ft
Btu/hr ft2F/in 0.083333 0.08333 3
Btu/hr ft F
2
Btu/hr ft F
1.7307 1.730 7
W/m C
2
kcal/hr m C
0.67197
Btu/hr ft F
2
W/cm C
57.779 57.77 9
Btu/hr ft F
in ft
2/
ft ft
DYNAMIC
Pa s
2419.1
lb/hr lb/hr ft
2
VISCOSITY
m /m
cP
2.4191 2.419 1
lb/hr ft
2
kg/hr m
0.67197
lb/hr ft
3
lb/s ft
3600.0
lb/hr lb/hr ft
115827.0 115827. 0
lb/hr ft
m ft
3
lbf s/ft
3
lb/hr
0.00041338 0.00041338 Pa s
3
lb/hr lb/hr ft
0.41388
cP
m /s cSt m2 /s kj
38750.0 0.038750 10000.0 0.94782
ft /hr ft /hr Stokes Btu
lbf lbf
KWhr kcal
3412.1 3412. 1 3.9683
Btu Btu
0.00057870 ft
3
ft
0.028317
gal
0.13368
gal (IMP) litter kg lb
1.2009 0.26417 2.2046 0.45359
N kp
0.22481 2.2046
m ft
gal (US) gal (US) gal kg
KIMEMATIC VISCOSITY ENERGY
2
2
lb ft/s2
ft lbf
0.0012851 Btu
kp
9.80665 9.80665
N
hp hr
2544.4
0.0040218 4.0218
in of water in of water
Btu Btu
1.0551 kj 0.0002931 0.0002 931 kWhr
0.14504
lbf/in lbf/in2
Btu
0.25200
0.039441
in of water
J
0.00027778 W hr
W
3.4121
Btu/hr
kcal/hr
3.9683
Btu/hr
2
kp/m
mm of water 0.09370 0.09370
in of water
in of water
kPa
0.24864 0.24864
POWER
2
Btu
k cal ca l
in of water
0.036063 0.03606 3
lbf/in lbf/in
ft lbf/hr
0.0012851 Btu/hr
in of water
0.0024539
atmospheres
Hp
2544.4
Btu/hr
3.2808
ft/s
Btu/hr
0.29307 0.29307
W
m/min
3.2808
ft/min
Btu/h r
0.25200
kcal/hr
mi/hr
1.4667 1.4667
ft/s
W/m2
0.31700
Btu/hr ft2
ft/s
0.30480
m/s
Kcal/hr m2
0.36867
Btu/hr ft2
ft/min ft/min
0.016667
ft/s
W/cm2
3170.0 3170. 0
Btu/hr ft 2
kg/s
7936.6
lb/hr
Cal/s cm
2
13272.0
Btu/hr ft
Btu/hr ft
2
3.1546
W/m
Btu/hr ft
2
2.7125
kcal/hr m
5.6783 5.678 3
hr ft2 F/Bt
VELOCITY m/s m/s
HEAT FLUX
FLOW
2
32.1740 32.1740
kPa
MASS
2
lbf PRESSURE Pa kPa
MASS
TO OBTAIN
SPECIFIC
2
FORCE
BY
in
cm
MASS
MULTIPLY
0.039370 0.03937 0
2
VOLUME
TO OBTAIN
mm
ft AREA
BY
Table 5
kg/hr
2,2046
lb/hr
THERMAL
2-14
m2 C/W
2
2 2
MULTIPLY FLOW
BY
TO OBTAIN
hr ft2 F/Bt
kg/hr
hr m2 C/kcal
4.8824
hr ft2 F/Bt
2118.9
ft3 /min /min
s cm2 C/kcal C/kc al
0.00013562 hr ft2 F/Bt
ft 3 /s
60.000 60.00 0
ft3 /min
Cm2 C/W
0.00056783 hr ft2 F/Bt
ft3/ min
0.00047195 m3/s
hr ft2F/Btu
0.17611
m2 C/W
kg/s m2
737.34 737.3 4
lb/hr ft2
hr ft 2F/Btu
0.20482
hr m2C/W
020482
lb/hr ft
2
HEAT
W/m C
2
0.17611
Btu/hr ft F
3600.0 3600. 0
lb/hr ft
2
TRANSFER
kj/hr m C
0.048919 0.04891 9
Btu/hr ft F
4.8824
kg/hr m2
COEFFICIENT
kcal/hr m2 C
0.20482
Btu/hr ft2F
3
0.062428
lb/ft
g/cm
3
62.428 62.42 8
lb/ft
3
1728.0 1728. 0
lb/ft
3
16.018
kg/m
lb/hr
lb/hr
0.45359 0.45359
VOLUME
m3 /s
FLOW
2
VELOCITY kg/hr m lb/s ft
2
lb/hr ft 2 DENSITY
kg/m lb/in lb/ft
RESISTANCE
2
2 2
3
cal/s cm C
2
7373.4 7373. 4
Btu/hr ft F
3
W/cm C
1761.1 1761. 1
Btu/hr ft F
3
Btu/hr ft F
2
5.6783
W/m C
2
4.8824
kcal/hr m C
2
3
TEMPERATURE 1.8c + 32 = F (f-32)/1.8 = C C + 273.15 = K F + 459.69 = R
TO OBTAIN
20.442 20.44 2
3600.0
2
BY
hr m C/kj
lb/s
MASS
MULTIPLY
Btu/hr ft METRIC EQUIVALENTS 2 2 Pa = w/m = kg/s m N = kg m/s 2 J = Ws 3 Liter = dm Kp = kgt
2-15
2 2
2
2
E.
Heat Exchangers (General)
1. Heat Exchanger Area Approx.
Area = [Heat Duty, Btu/hr]/[U.LMTD] Where: U = heat transfer coeff. (GPSA typical U’s Heat Exch.)
2. Limit temperature approach in gas to gas exchanger to 20 ° F. 3. For preliminary preliminary design for cooling water systems, systems, use a cooling cooling water temperature temperature rise of 15° to 20° F through the heat exchangers. exchangers. In most cases a process process stream temperature approach of 10° F to the cold water to the exchanger is reasonable. 4. If flow through through the exchanger is not countercurrent, countercurrent, hot fluid fluid outlet temperature temperature should be greater than cold fluid outlet temperature. 5. For exchangers of 200 ft2 and less, consider compact heat exchangers. F.
Condensers Condensers
1. Be aware when when condensing pure pure components such as propane propane that the the limiting limiting temperature occurs when the desuperheating stops and condensing starts. 2. For water cooling, cooling, try to to cool no further further than a 10° F approach to to the warm cooling water leaving the condenser. Use 4 to 8 ft/sec velocity for water through the tubes. Restrict cooling water return temperature to maximum of 125° F.
G.
3. For water cooled propane condenser condenser design, generally generally use 10° F temperature temperature rise on cooling water through the exchanger. exchanger. Reboilers and Chillers 1. Many failures failures occur because because the pressure pressure on the condensate condensate return return header is higher than the low pressure steam at the reboiler. 2. About 50° F delta T is all that can profitably profitably be used. Limit the approach temperature of the gas to the refrigerant in gas chillers to 10° F. Less than 10° F delta T requires excess exchanger surface area. 3. Usually design design reboilers reboilers for a conservative heat flux of 8,000 to 12,000 BTU/ft2 and reduce pressure of steam to prevent film boiling. 4. Submergence Submergence of Bundle Bundle – Level generally controlled controlled at top of bundle. 5. For thermosiphon and side reboiler designs for demethanizer columns, limit the vaporization of the reboiler liquid stream stream to a maximum of about 35% by volume.
2-16
H.
Attempting to vaporize more fluid may result in problems with the thermosiphon flow. Sizing Plate Heat Exchangers A method for calculating plate heat exchangers is presented in “Plate-Type Heat Exchangers” by F. J. Lowry, Chem. Eng., 66, 89-94, June 29, 1959. Generally, accurate sizing of [late heat exchangers must be done by the manufacturer.
I.
Brazed Aluminum Plate Heat Exchangers
1. For aluminum plate fin reboilers, methanol may tend to accumulate in the reboiler and eventually log off the exchanger limiting thermosiphon flow. It can usually be cleared if a drain is provided on the lower header of the exchanger. 2. Mercury occurring naturally in some natural gas steams is extremely corrosive to aluminum heat exchangers used extensively in LNG plant processes. Plan to check for mercury in feed gas up front in any project. 3. For Aluminum Plate Fin Core-in-Shell evaporator heat exchange design use a 2 to 4 degree temperature approach to shell side evaporating fluid temperature. Use a maximum evaporation of 25% of the thermosiphon circulated fluid in the evaporator when preparing preliminary core specifications. J.
Air Fin Heat Exchangers
1. All air coolers should be designed in accordance with API 661. 2. Normally design for 40° F approach to inlet air temperature if no cooling water used (20° F Minimum). 3. Forced draft units are preferred over over induced draft units. 4. For preliminary estimates assume 4 rows of tubes. Estimate power requirements at 3 HP/MMBTU/hr for fans for face velocity of air to the coil of 450 to 550 ft/min. 5. Limit tip speed of fans: < 9 feet in diameter to 12,000 ft/min (FPM) >9 feet in diameter to 11,000 ft/min. 6. Hot air recirculation can be a problem, especially in hot weather. Consider air recirculation when locating air Cooled Exchangers. Locate coolers away from taller buildings or structures, especially downwind of the cooler. Do not locate coolers downwind of other heat generating equipment: i.e. furnaces, boilers, etc. §
§
2-17
Mount coolers high enough from the ground to avoid high inlet air approach velocities. Consider mounting them on pipe lanes or provide at lease ½ a fan diameter clearance between the ground and the plenum. Locate large banks of coolers with the banks long axis perpendicular to the prevailing summer wind direction. Do not mix forced and induced draft coolers in close proximity and do not locate coolers of different heights in close proximity. Fired Heaters §
§
§
K.
1. Maximum recommended heat flux for a direct fired Triethylene Glycol regenerator in a TEG Dehydration Unit is 8000 BTU/ft2 of fire tube surface area. The recommended heat flux for maximum fire tube life is 6000 BTU/ft2 . 2. For most process heaters, assume a thermal efficiency of 75 to 80% when calculating fuel requirements. Where: % Thermal Efficiency = (Heat Transferred/Heat Released)*100. 3. Organic Heat Transfer Fluids A. Fired heaters for organic heat transfer fluids are usually designed with average radiant heat fluxes ranging from 5000 to 12,000 BTU/hr-sq ft. Actual allowable heat flux is usually limited by fluid maximum allowable film temperature. Film temperature is dependent on: a. Maximum fluid bulk temperature b. Velocity of the fluid across the heat transfer surface c. Uniformity of heat distribution in the furnace d. Heat transfer properties of the heat transfer fluid B. If too high film temperature results, too much fluid is vaporized and the heat transfer surface is blanketed with vapors. The heat transfer coefficient is rapidly reduced and dangerously high surface temperatures can develop resulting severe fluid degradation and mechanical fa ilure.
L.
C.
High surface temperatures may also cause the fluid to carbonize forming carbon scale on the heat transfer surface which may lead to over heating and tube metal failure.
D.
All other things being equal, any organic heat transfer fluid degrades in proportion to its temperature. Operation at approximately 100° F below vendors maximum recommended bulk fluid operation temperature may extend the life of the fluid by 10 times.
Cooling Towers
2-18
1. The evaporation rate on a cooling tower is dependent on the amount of water being cooled and the temperature differential. For each 10° F temperature drop across the tower, 1% of the recirculation rate is evaporated. In other words, 0.001 times the circulation rate in gpm times the temperature drop equals the evaporation rate in gpm. 2. Windage losses for cooling towers: Spray ponds 1.0 to 5.0% of circulation Atmospheric Cooling Towers 0.3 to 1.0% of circulation Forced Draft Cooling Towers 0.1 to 0.3% of circulation Evaporation Losses for Cooling Towers: Evaporation Losses are usually 0.85 to 1.25% of the tower circulation rate. An evaporation loss of 1% of tower circulation per each 10 degrees F temperature drop across the tower can be assumed for estimating purposes. 3. Cooling Water System Feasibility Design: Feasibility designs for cooling water systems may be completed by setting the water temperature rise across all exchangers, usually 15 to 20 °F rise, (or at a 10 °F approach to the process outlet temperature if the assumed rise results in a temperature cross for some exchanger), and setting the inlet water temperature to the exchangers to the site wet bulb temperature plus 8 °F. 4. Cooling Water System Fluid Flow and Piping: For preliminary sizing branch offs with different flowrates from the main header, the following rule of thumb equations may be used. D2 = Summation di2 qi/di2 = Q/D2 M.
where: Q and qi are volumetric flowrates through the header and branch i, and D and di are the diameters of the header and branch i. Round to nearest standard size.
Insulation
1.
For estimating insulation thickness: Thickness = {3 +[(T - 100)]} Truncated / 2 Thickness – inches T (Process Temperature) – ° F
2. Typical thermal conductivities for insulating materials The first table below contains recommended insulation conductivities for insulating materials. The second table contain conductivities for various materials.
2-19
Representative Conductivities of Pipe Insulation (Btu/hr·ft·F) Insulation temperature 300º F 400º F (149º C) (204º C)
100º F 38º C
200º F (93º C)
calcium silicate
0.033
0.037
0.041
cellular glass
0.039
0.047
0.055
fiberglass
0.026
0.030
0.034
magnesia, 85%
0.034
0.037
0.041
Polyurethane
0.016
0.016
0.016
2
500º F (260º C)
600º F (316º C)
0.046
0.057
0.060
0.064
0.074
0.085
0.044
2
(Multiply Btu-ft/hr-ft -º F by 12 to get Btu i n/hr-ft -º F.) (Multiply Btu-ft/hr-ft2 -º F by 1.730 7 to g et W / m·K.) (Multiply Btu-ft/hr-ft2 -º F by 4.1365 x 10-3 to get cal·cm/s·cm2 ·º C.)
Material
Asbestos-cement boards Asbestos Kaolin brick Kaolin firebrick Petroleum coke Molded pipe covering Mica Aluminum Iron Steel N.
NGL Expander Plants
1.
Thermal Conductivity Btu/hr-ft-°F 0.43 0.090-0.129 0.15-0.26 0.050-0.113 3.4 0.051 0.25 117 30 26
For thermosiphon reboiler and side reboiler designs for demethanizer columns, limit the vaporization of the reboiled liquid stream to a maximum of about 35% by volume. Attempting to vaporize more fluid may result in problems with thermosiphon flow.
2-20
2.
O.
For aluminum plate fin reboilers, methanol may tend to accumulate in the reboiler and eventually log off the exchanger limiting thermosiphon flow. It can usually be cleared if a drain is provided on the lower header of the exchanger.
Miscellaneous Plant Systems
1.
Cooling Water Systems – For Preliminary design for cooling water systems, use a cooling water temperature rise of 15 to 20° F through the heat exchangers. In most cases a process stream temperature approach of 10 ° F to the cold water to the exchanger is reasonable.
2.
For Aluminum Plate Fin Core in Shell evaporator heat exchanger design use a 2 to 4 degree temperature approach to shell side evaporating fluid temperature. Use a maximum evaporation of 25% of the thermosiphon circulated fluid in the evaporator when preparing preliminary core specifications.
3.
Wind Chill & Tw = 33-[(10.45+10 V ) (33-T)]/32 Heat Loss H = (10.45+10 V – V)(33-T) Where: Tw = Wind chill temp. °C T = actual temp. V = wind speed in meters/sec. H = heat loss in kcal/m2 -hr.
4.
Heat Transfer From Pipes
2-21
2-22
5.
Typical material emissivities for radiation heat transfer problems Material Aluminum
Iron
Steel
Brick Refractory Paint
Polished Oxidized Polished Polished cast New cast Rusted Polished Oxidized Rough plate Poor Good Black matte Black lacquer White lacquer Aluminum
2-23
Emissivity 0.040 0.11-0.19 0.14-0.38 0.21 0.435 0.685 0.52-0.56 0.79 0.94-0.97 0.93 0.65-0.75 0.80-0.90 0.91 0.80-0.95 0.80-0.95 0.27-0.67
P. Method For Feasibility Study Sizing of Gas Plant Gas/Gas shell & Tube Heat Exchanger:
1. From the process simulator output for the process, determine the required UA rate for the gas/gas exchanger. Assume U = 60 BTU/Hr Ft °F A = UA = UA U 60 Assume a 20 ft long exchanger with ¾” OD tubes on a 15 /16 ” triangular pitch. Go to a Tube Count Table and read the number of tubes required for the Area A and unit diameter and/or number of units. You now have a feasibility estimate which includes: 1. Exchanger Area (Ft2 ) 2. Number of ¾” tubes 3. Unit length, diameter, and number of units.
2-24
Start Surface A>150 ft2
Yes
Q Recovery @ T > 1000F
No No Q Removed
No
@ T> 140F
Yes
Yes
Exotic Alloy
Yes
Q Recovery Economical
Yes
Yes No Close dT Approach
T > 350F P > 200 psi
Yes
No
No
Yes
Double pipe Exchanger
RODBaffle Exchanger
Vibration Low dP
No
Plate Baffle Exchanger
Figure 1 Heat Exchanger Selection
No
2-25
PlateFrame Exchanger
Air Finned Exchanger
Heavy Duty Finned Surface
Start Fouling Service
Extent of Fouling
Low to Moderately Heavy
Very Heavy
Viscosity Low To Moderately High Very High
Pressure
Less Than Atmospheric
Kettle Reboiler (Finned Tubes) Pump-Through Reboiler
Pump-Through Reboiler (Critical Operations)
Area Required
Greater Than Atmospheric
Small To Moderate
Vertical Thermosyphon Horizontal Thermosyphon
Pump-Through Reboiler
Kettle Reboiler (Finned Tubes)
Figure 2 Reboiler Selection
Relatively Clean Service
Fluid Fouling Characteristics
2-26
Internal Reboiler Vertical Thermosyphon
Large
Kettle Reboiler Horizontal Thermosyphon
Start Corrosive High Pressure
Yes
No Mechanical Cleaning Tubeside Coolant
Yes
Change Design To Reduce Condensing dP
Yes
Very Low Allowable Condensing dP
Use Large Diameter Tubes
Yes
Large Condensing Range
No
Yes
Yes
No
No Internal/ Kettle Reboiler
No
Yes
Vertical E Shell
Yes
Large Subcooling
Yes No
Horizontal J Shell RODBaffle
Tubeside Condensers Horizontal RODBaffle J Shell
Vetical Downflow Single Pass
Figure 3 Condenser Selection
Boiling Coolant
No
Yes
Shear Control At Exit
Yes
Shellside Condensers Horizontal E Shell
No
Yes
No
No
Low Medium
Temperature Cross
(dP/P) > 0.1
Allowable Condensate dP
Very Low
Low
Large Subcooling
No
Yes
Med-High
Bioling Coolant
ShellSide Condensation Required
No
2-27
Horizontal Single Pass or U Tube
HEAT TRANSFER REFERENCES
Kern, D. Q., Process Heat Transfer, McGraw-Hill, 1950 Tabork, J. et. al, Heat Exchangers, Theory & Practice, McGraw-Hill, 1981 Phillips Eng. Std. 10.44-2, Shell & Tube Process Design Criteria Phillips Eng. Std. 10-44-3, Reboiler Characteristics & Selection Phillips Eng. Std. 15.18-4, Shell & Tube Mechanical Design Criteria Phillips Eng. Std. 15.18-5, RODbaffle Heat Exchanger Specifications Phillips Engineering Standard 15.18-2, Air Cooled Heat Exchanger Mechanical Design Criteria Phillips Engineering Standard 25.04-85, General Design, shell and Tube Heat Exchangers, Mechanical Fabrication Requirements Phillips Engineering Standard 25.04-89, Heater-Fired-Mechanical Design Specifications Perry, R. H., Chemical Engineers Handbook, Section 11, 4th Ed. 1963 HTRI Design Manuals, Vol. I & II. Premises for Design & Specification of Shell & Tube Heat Exchangers, 1992 Gas Processors Suppliers Association Engineering Data Book, Tenth Edition, 1987; Volume I, Sections 8, 9, & 10 Standards of the Tabular Exchanger Manufacturers Association, Seventh Edition, 1988 API Standard 660, Fifth Edition (to be issued in 1993), Shell-And-Tube Head Exchanger for General Refinery Services API Standard 661, Third Edition, April 1992, Air-Cooled Heat Exchanger for General Refinery Services “Quick Calculation of Cooling Tower Blowdown and Makeup”, Chemical Engineering, July 7, 1975, pg 110 “Designing a Near Optimum Cooling-Water System”, Chemical Engineering, April 21, 1980 pg 118-125) “Guidelines on Fluid Flow systems”, Hydrocarbon Processing, April, 1990. Pg 47-60
2-28
HEAT TRANSFER REFERENCES (CONTINUED)
The Randall Corporation process group used this method for preliminary estimates and reports close match to their final design “Organic Fluids for High Temperature Heat-Transfer Systems” W. F. Seifert, and L. L. Jackson, Chemical Engineering, October 30, 1972, pg 95-104
2-29
3 A.
TREATING
Dehydration
1. Dehydrate gas to 60% of the saturation water content at the conditions of lowest saturation. Sulfur Recovery Units A. Thermal zone will produce 55-65% of the sulfur and is a function of the H2 S content of the feed. Catalytic region makes the rest.
B. If the acid gas feed is less than 30% H2 S then flame stability in the reaction furnace is a potential problem. Minimum temperature for effective operation is 1700° F. C. Temperature in catalyst beds should be kept below 800° F. D. SRU steam production will be approximately 6700 lbs of steam per long ton of sulfur produced. E. Glossy carbon deposits on catalyst indicates amine carryover. F. Sulfur fog is caused by too much cooling capacity. Sulfur mist can be caused by excessive velocity in the condenser. G. Ferrules should extend at least 6” inside the tubesheet. Refractory lining is usually 2 12 - 3" thick on the tubesheet. H. Mass velocity in waste heat exchanger and sulfur condenser tubes should be 2-6 lbs/sec-ft2 . I. Space velocity through catalyst beds should be 700-1000 SCFH of gas per cubic foot of catalyst. Lean streams, lower value and rich streams, higher value. J. Sulfation of catalyst caused by SO3 . Oxygen combines with SO2 to form SO3 which is chemisorbed on alumina surface. K. Velocity in process piping should not exceed 100 ft/sec. L. Liquid sulfur solidifies at 246° F and becomes very viscous above 320-350° F M. Approximate Stack Gas Flow, scfm: SGF = (Sulfur Production, LT/D) x (100) 2. Glycol Dehydration
3-1
a. TEG – Dew point depression ranges 80-140° F. Degree of dehydration which can be obtained depends on amount of water removed from glycol in the reboiler & circulation rate. Minimum circulation rate to assure good glycol gas contact is approx. 2 gal. glycol for each pound of water to be removed. Max is ≈ 7 gal. and standard is ≈ 3 gal. b. Stripping Gas-Approx. 3-8 scf/gas of glycol circ. c. Glycol will absorb ≈ 1 scf of gas/gallon of glycol. Glycol contactor – For best scrubbing of overhead gas install “Mist Pad” on the face of “Vane Type” mist extractor. d. Estimate total reboiler duty from 2000 BTU/US gal of TEG circulation rate. Note that the use of glycol/glycol heat exchangers will reduce the total reboiler duty. e. Estimate glycol loss from 0.1 gal TEG/MMSCF. f.
Packing-Minimum of 4’ in any gas-glycol contactor.
g. Triethylene Glycol Dehydration Unit-Maximum recommended heat flux for a direct fired TEG regenerator is 8000 BTU/square foot of fire tube surface area. The recommended heat flux for maximum fire tube life is 6000 BTU/ft2 . 3. Trouble shooting
B.
Black, viscous solution indicates that heavy hydrocarbons have been carried over with the gas. Sweet, burnt sugar smell accompanied by low pH and a dark, clear solution signals that thermal degradation is occurring. Amine Treating 1. Amine Circulation: 3 cu.ft acid gas/gal amine Reboiler Steam Rate: ≈ 1.2 lbs steam/gal amine MEA gpm = 41.0 * Q*X/Z DEA gpm = 45.0 * Q*X/Z (conventional) DEA gpm = 32.0 * Q*X/Z (high load) Where Q = Gas, MMscfd X = Acid Gas, volume percent Z = Amine Concentration, wt.% 2. Max acid gas pickup not more than 0.35 mols/mol of MEA. Normal value around 0.3. 3. Amine treating processes tend to be troubled by the same problems regardless of the type amine used.
3-2
4. Typical MEA losses due to entrainment: Absorber: 1.0 #/mmscf Still: 2.5 #/mmscf 5. Flow Velocity – Rich Stream: Not to Exceed 5 fps Lean Stream: Not to Exceed 7 fps 6. Filter Beds: Recommended flow rate through carbon bed 4 gpm/ft2 (cross sectional area) ≈ 20 minutes superficial contact time 7. Loadings: .36 mols CO2/mol MEA [absorber RICH] .12 mols C02/mol MEA [still LEAN] 8. Reflux Ratio: MEA & DEA 1.5 to 3.0 [mols H20/mols acid gas leaving reflux drum] 9. Equivalent Steam Rate: MEA 0.9 to 1.2 lbs steam/gal amine DEA 0.8 to 1.1 lbs steam/gal amine 10. Lean amine can contain 0.05 to 0.08 mols total acid gas and still meet specs. 11. CO2 and H2S gases appreciably increase total water content and dehydration requirements of gas streams. 12. Recommended maximum ranges for amine strength and acid gas loadings that have proven historically to adequately address corrosion concerns are: Amine MEA DEA MDEA
Wt% 15 - 20 25 – 30 50 – 55
Rich Loading, M/M 0.30 – 0.35 0.35 – 0.40 0.45 – 0.50
13. Recommended loading in the lean circuit to minimize acid gas flashing are: Amine MEA DEA MDEA
* Total Lean Loading, M/M 0.10 – 0.15 0.05 – 0.07 0.004 – 0.010
* These loadings should be easily achieved with a 1.0-2.0 M/M stripper reflux ratio. 14. Recommended Minimum Water Quality Standards for Make- up Water for Amine Plants:
3-3
Total Dissolved Solids Total Hardness Chlorides Sodium Potassium Iron
<100 ppm <3 grains/gal <2 ppm <3 ppm <3 ppm <10 ppm
15. Liquid/Liquid Contactors: (Feasibility Sizing Data) For rough diameter sizing of liquid/liquid contactors fo r amine treating of light hydrocarbon liquids, use 12 to 15 GPM of hydrocarbon per square foot of packed tower cross section area. This should correspond to approximately 10% of flooding velocity. For hydrocarbon distributor nozzles for liquid/liquid contactors, use an orifice velocity of approximately 1 ft/sec. Higher velocities than this can lead to emulsion problems. Velocities lower than 0.5 ft/sec can result in NGL being entrained in the sour amine stream. To estimate height of packing required, assume 6 to 8 ft of packing for each theoretical separation stage. Mercaptan Removal from Gas
MEA & DEA will remove approx.: 40-55 mol% methylmercaptan 20-25 mol% ethylmercaptan 0 –10 mol% propylmercaptan Regenerative Caustic process will remove mercaptans down to <10 ppm.
C.
Activated Carbon, Calgon FCA, will remove 4- 5 wt% mercaptans. Mol Sieve Treating 1. Bed Design (Length/Diameter) Minimum Liquid L/D 3:1 Gas L/D 2:1
Maximum 5:1 4:1
2. Max. Gas Velocity: 0.33 to 0.75 ft/sec (superficial linear velocity) 3. Max. Liquid Velocity: 50 bbls/hr/ft2 (bed area) Dehy-ALCOA: 30 gpm/ft2 ( ≈ 43 bbl/hr/ft 2 )
3-4
4. Min. Velocity: Liquid: 60 sec contact time, or 0.01 psi/ft ∆ p (liquid). Gas: 3.5 sec (gas) 5. Mol Sieve – Draining bed leaves approx. 25 vol% of total bed volume on bed as sponged liquid. 6. Alumina – Draining bed leaves approx. 0.048 gals. per pound of alumina on bed as sponged liquid.
D.
7. Molecular Sieve Dehydrators – As strictly a rule of thumb based on many Phillips designs, when the pressure drop through a mol sieve bed reaches 20 psi, the bed support is nearing its maximum load capacity and action should be taken to reduce the pressure drop. Corrosion 1.
C02 Corrosion: Low Corrosion – Pco2 < 7 psia Possible High Corrosion - Pco2 ≈ 7-15 psia High Corrosion – Pco2 > 15 psia Where: Pco2 = Partial Pressure of C02
E.
2. Corrosion rate directly related to temperature. Copper Strip
F.
Copper Strip Test ASTM 5.05 D1838 No. 1A copper strip normally < 1-2 ppm H2S. H2S corrosive to copper strip ≈ 1ppm or .16 gr/100 scf. Copper Strip will not detect Mercaptan or other Organic Sulfides. Conversion Factors
G.
H2S & C02 factors 1 mol% H2S = 630 grains/100 scf 1 mol% C02 = 813 grains/100 scf Grains of H2S/100 scf x 1.591 x 10 = Mol% H2S 1 MMscf H2S = 37.6 long tons sulfur 1 grain H2S/100 scf = 17.1 ppmw = 22.8 mg/m 1 grain H2S/100 scf = 15.9 ppmv 1 grain C02/100 scf = 12.3 ppmv H2S ppmw = [gr. H2S/100 scf] x [542/(mol wt gas)] Caustic Washer Design
H.
Vertical washers sized by using a factor of 400-500 gal/hr/sq.ft. of cross-sectional area of empty tower. In the tower, 5-7 ft. of raschig rings equal 1 stage. Metallurgy Requirements For Amine Treaters
3-5
1. Vessels a. Amine contactor, flash tank, stripper, surge tank, accumulator, inlet scrubber, and outlet scrubber shall be carbon steel and stress relieved with corrosion allowances as shown in 1.C. b. Trays for the contractor and stripper should be 304 stainless. c. Corrosion Allowances for MEA and DEA Systems C02 /H2 S<20 Inches 1/8 1/8 1/16 1/8 1/8 3/16 1/8 1/8 1/4 0 1/16 1/8 1/8
Inlet Scrubber Amine Contactor Outlet Scrubber Flash Tank Cross Exchanger Amine Stripper Reflux accumulator Reboiler Reclaimer Surge Tank Piping Amine Cooler* Stripper Overhead Condenser*
C02 /H2 S>20 Inches 1/8 1/8 1/16 1/8 1/8 3/16 1/4 1/8 1/4 0 1/16 1/8 1/4
*Corrosion allowance applies to shell side exchangers with water cooling in the tubes.
2. Heat Exchangers a. Shells – carbon steel and stress relieved b. Tubes – 12 gauge minimum, carbon steel, seamless c. Reclaimer element – carbon steel, 2- inch schedule 80 tubes. Amine temperature in reclaimer should not exceed 310° F. d. Temperature of amine in reboiler should not exceed 250° F. e. U-bends of U-tube carbon steel bundles shall be stress relieved. 3. Pumps a. The amine circulation and stripper reflux pumps shall be carbon steel with 316 stainless trim. 4. Piping
3-6
a. Piping should be carbon steel. The weld and heat affected zone of piping containing H2 S and H2 0, with or without amine, shall have hardness no greater than Brinnell hardness number 235. b. Velocities in the rich and lean solution piping should be limited to 2-3 ft/sec. for MEA solution, and to 7 ft/sec. for DEA solution. c. Piping from the letdown valve to the flash tank and from the letdown valve to the stripper should be 304 stainless with the letdown valves of 316 stainless. 5. An inert gas blanket should be maintained on the fresh amine storage and amine surge to prevent oxygen contact with the amine. 6. No copper-bearing materials such as Admiralty, Monel, etc., should be used anywhere in amine units. When the C0 2/H2S ratio is greater than 20, the following exceptions to the above requirements should be used.
7. The exchanger tubes in the reboiler, stripper overhead condenser, lean-rich cross exchanger, and amine cooler should be 304 stainless. Those exchangers with water in the 304 stainless tubes should have a minimum water flowrate of 5 ft/sec. Limit amine velocity in tubes to 5 ft/sec. 8. The stripper vessel head and shell down through the top 3 trays should be 304 stainless clad or solid 304 stainless. 9. The stripper overhead piping from the stripper to the accumulator should be 304 stainless. The reflux piping from the accumulator back to the stripper can be either thin wall 304 stainless or carbon steel with 1/16” corrosion allowance.
I.
10. The vapor line from the reboiler to the stripper should be 304 stainless. H2S Gas Toxicity ppm gr./100scf 10 0.65 100 6.48 200 12.96 500 32.96
J.
Can smell. Safe for 12 hr. exp. Kills sense of smell in 2-15 min. Kills sense of smell quickly. Stings eyes & throat. Loses sense of reasoning & balance Respiratory paralysis in 30-40 min 700 45.36 Breathing will stop & death result if not rescued promptly. Immed. artificial resuscitation. 1000 64.80 Unconscious at once. Permanent brain damage or death may result unless rescued promptly. Liquefied Natural Gas (LNG) Plants Feed gas to LNG plants should be treated to provide a maximum CO2 content in feed gas to liquefaction unit of 100 ppmv.
3-7
K.
Gas Treating Iron Sponge
L.
H2S Removal: W = 1.43*GQ = .09*PQ Where: W = H2S removal lbs/day G = H2S gr/100 scf inlet P = H2S ppmv Q = MMscfd Distribution of Sulfur Compounds in NGL Product 1.
H2 S will be concentrated mainly in the C3 and lighter.
2.
COS will be concentrated mainly in the C3 streams.
3.
CH3 SH will primarily be split between the C 3 and C4 streams.
4.
CH3 CH2 SH will be concentrated mainly in the C4 and heavier streams.
A.
LNG Process – Aluminum Platefin Mercury Corrosion:
1.
Any natural gas stream used for feed to a gas liquefaction plant will likely contain measurable amounts of mercury.
2.
Elemental mercury amalgamates (forms an alloy) with the surface layer of the aluminum being corroded.
3.
It is probable that liquid water must be present for the corrosion or attack to occur. It follows that the temperature of the system must be above 32° F for the corrosion to take place.
4.
Most probable time for mercury corrosion to occur is during deriming or defrost periods (or possibly unscheduled shutdown for extended time periods) when free water is likely to be present. It follows that the most likely time for mercury corrosion induced equipment failure is during start-up immediately after a planned or extended emergency shutdown.
5.
It is not known at what levels of mercury in the inlet gas problems may occur. However, it is likely most of the mercury entering with the feed in a cryogenic LNG plant with no liquid draws will remain within the plant equipment until defrost. A number of studies associated with the Skikda failure have suggested a threshold mercury content of 0.01 micrograms/Nm3 which should not be exceeded at the inlet to the cryogenic plant.
6.
Mercury can exist in the feed gas co-currently with H2 S as evidenced from an LNG plant feed gas in Sumatra, Indonesia which contains 200 to 330 micrograms/Nm3 mercury and 60 to 70 ppm of H2 S.
3-8
B.
7.
Mercury is more soluble in heavier liquid hydrocarbons than the lighter hydrocarbons. (For instance solubility of mercury in propane is about ½ the solubility in hexane). Mercury in an LNG plant should concentrate in the heavier liquid phases such as a hydrocarbon draws for heavies removal or fractionation.
8.
There is more risk of mercury induced failure on the cycle gas (refrigerant) side of the exchangers, and more risk in general with a mixed refrigerant plant than a cascade cycle since a cascade system uses no refrigerant heavier than propane. (This assumes refrigerants are derived from the mercury contaminated feed gas).
9.
If stress is present in the parent aluminum (such as at weldments), cracking will also occur accelerating the corrosion.
10.
The pipeline system feeding the LNG plant is an excellent trap as mercury tends to be absorbed by the pipe metal. The initial mercury front progress towards the plant slowly until pipeline saturation is reached. This is a good reason to test the plant feed gas at least annually.
Treating Systems: Mercury Removal with sulfur Impregnated Carbon: Following are Vendor provided design rules of thumb for mercury removal beds: 1. Design for minimum gas/carbon contact time of 10 seconds. 2. Design for a maximum superficial velocity of 50 ft/min. 3. Maximum temperature is 150° F. Above this the sulfur impregnate is boiled off. 4. C02 and other sulfur compounds do not effect the beds performance. 5. Water vapor or free water present in the gas stream being treated will reduce the mercury removal capacity of the beds as much as 75%. 6. Liquid hydrocarbons will rapidly deplete the bed mercury removal capacity as the liquids dissolve the sulfur impregnated on the carbon. 7. L/D (actual carbon bed height to vessel diameter) ratios of 1.0 to 1.5 are preferred for charcoal beds to prevent channeling. Minimum acceptable L/D is about 0.6. 8. Mercury removal efficiency is better for dry gas than for water saturated gas. 9. Mercury removal efficiency is better for lower feed gas temperatures.
3-9
10. Maximum attainable mercury removal efficiency is not effected by inlet mercury content or inlet gas pressure. 11. Can be designed to remove feed gas mercury down to less than 1 nanogram/Nm3 (0.001 microgram/Nm3 ). 12. Vendors report that bed capacity is not affected if the bed becomes saturated with gas phase heavy end hydrocarbons (as long as no free liquid is allowed to contact the bed). In fact vendors often presaturate the beds with hydrocarbons for ethylene service to prevent high temperatures resulting from initial adsorption in beds at start-up. 13. Removal mechanism is as follows: elemental mercury physically adsorbs onto the activated carbon surface inside the pores of the carbon. Then the mercury chemically reacts with the sulfur impregnate to form mercuric sulfide. For organic mercury compounds; Vendor 1: the systems function similarly except the mercury compounds much break down to elemental mercury and organic compounds first. Vendor 2: the organic mercury compounds are physically absorbed on the carbon. This organic mercury criteria sets the minimum 10 second contact time.
3-10
TREATING REFERENCES
GPSA Engineering Data Book – Tenth Edition 1987 Gas Conditioning and Processing, Vol. 4, Gas and Liquid Sweetening, Campbell Petroleum Series, 1982 Chemical Engineer’s Handbook, Perry & Chilton, Fifth Edition Gas Purification, Kohl & Riesenfeld, Fourth Edition 1985 Gulf Publishing Company Acid and Sour Gas Treating Processes, 1985, Gulf Publishing Gas Conditioning Conference Proceedings, 1954-1992 Linde Molecular Sieves, Union Carbide Gas Treating, Gas/SPEC Dow Corporate Engineering Process, Gas Section Reference, Files, GR1020, GR1180, GR1559, GR1800, GR1802, GR1820, GR1870, GR7047 Sulfur Recovery, Paskall & Sames, 1988, Western Research Oil & Gas Journal, Dec 1, 1980, pg 135-138 “Liquid – Liquid contractors need car eful attention” “Considerations For Mercury in LNG Operations” by W. W. Bodie, A. Attari – Institute of Gas Technology R. Serauskas – Gas Developments Corporation “Mercury-LNG’s Problems” – by J. E. Leeper – Hyrocarbon Processing – November, 1980 “Causes and Remedies of The Corrosion of Cryogenic Exchangers By Mercury” by Tewfik Hasni – February, 1978 Phillips Metallurgist – Hisham Hashim Calgon – HGR Sulfur Impregnated Activated Carbon Pinion (Alcoa/NUCON) – Mersorb Activated Carbon Adsorbent
3-11
4 A.
FLUID FLOW
Misc.
1. Absolute pressure of atmosphere at height ‘H’ above sea level: P = P1 (1-0.00000687H)n n = 5.256 Where: P1 = pressure at sea level, psia Density: W = W1(1-0.00000687H)n n = 4.256 Where: W1 = density of air at sea level P = e[2.6876 – 0.0000368 (H)] psia Where H is height, ft above sea level Borger- 13.2 psia: Woods Cross- 12.4 psia 2. Acoustic Velocity Va = 80.53 √ (P/ρ) where: P = psia ρ = density lb/ft3 Va = ft/sec
For perfect gas: Va = √ (gc*k*R*T/M) Where: gc = 32/17 k = Cp/Cv R = 1546 T = °R M = mol wt.
3.
Vortex Breaker Vortex Breaker is needed if flow is greater than 1.9 ft/sec.
B.
NGL Expander Plants
1. Flowrate through an expander can generally be controlled from 0 to 150% of the design value. Expanders adiabatic efficiency will generally be within 4 percentage points of the design between 75 and 125% of design flowrate. (#/hr) 2. Preliminary Piping Sizing For preliminary liquid piping sizing, the following table may be used as a guide:
Liquid Service
Pressure Drop
4-1
Velocity
(Psi/100 FT) 0.4 1.5 – 3.0 --0.4 ---
(FT/Sec) 2.0 – 4.0 7.0 - 10.0 3.0 – 5.0 --2.5 – 9.0 **
Pump Suction Pump Discharge To Reboiler Gravity Flow Water High Viscosity to 200 CP Pump Suction 0.5 – 1.0 0.25 – 0.5 Pump Discharge 10.0 – 1.0 1.0 – 1.5 ** Water with high CO2 , seawater, etc requires lower maximum velocities, linings, or special material. For preliminary vapor piping sizing, the following table may be used as a guide. Vapor Pressure Drop Velocity Total Allowable Service (PSI/100 FT) (FT/Sec) Pressure Drop (PSI) Tower 0.5 --1.0 Overhead On Plot 0.5 ----Gas Comp. 0.3 5-150 1.5 Suct Comp. 0.5 1y00 4.0 Disc Steam --100 4.0 C.
Piping
1. Initial maximum fluid velocities for line sizing: Most liquids: 10 ft/s All vapors: 50 ft/s Raw sea water: 11.5 ft/s (CuNi piping) Gravity drains 1.5 ft/s Steam condensate 1.0 ft/s 2. Fluid velocity for vapor and two-phase flows should not exceed the erosional velocity. Estimate erosional velocity from Ve = 100/√ρ where Ve = erosional velocity in ft/s and ρ = fluid density in lb/ft 3 . Limiting Velocities – Liquids: (Another source)
Normal limiting velocities (highest normal design velocities) in process lines are given by the following formula: 100 V = ρ
Where: V is limiting design velocity in ft/sec
4-2
ρ is density in lb/ft3 at system T&P Erosion velocity (velocity at which erosion of process line is expected) for clear liquids is given by the following formula: V e=
150 ρ
Where: Ve is the erosion velocity in ft/sec ρ is density in lb/ft3 at system T&P 3. Compressible gases (i.e. HC, air, steam) can be treated as incompressible when the pressure loss for the segment in question is less than 10% of the inlet pressure. Maximum Operable Velocities:
When rating existing liquid and/or gas piping systems, it is sometimes desirable to determine the limiting fluid velocity for the system. This may be estimated as follows: a. For liquids: Limiting velocity u m =
48 ρ
Where; um = limiting velocity, ft/sec ρ = density, lb/ft3 at system T&P For erosive or corrosive liquids = 0.5 x um b. For gases: Turbulent flow average limiting velocity u m = 148.7
kZT m
(2 /3 sonic velocity)
Where; um = Average limiting velocity, ft/sec K = specific heat ratio Z = Compressibility factor T = temperature, °R m = Molecular Weight lb/lb mol For erosive or corrosive gases = 0.5 x um Note:
Economic factors such as pressure loss usually necessitate operating at lower than limiting velocity. 4. Friction Factor – Project Life: Piping friction factor increases with operating time. As a result, the piping head loss nearly doubles in: 4-3
a. 25 to 30 years for clean gases and light hydrocarbons. b. 15 to 25 years for most middle distillates. b. 10 to 15 years for residues. While sizing pipelines and pumping facilities, this aspect should be duly considered. 5. Piping Noise:
D.
Liquid velocities above 20 to 30 ft/sec can cause noise. As a rule, a velocity head less than 1.3 psi avoids excessive noise. Physical Fan Laws 1. The following relations are characteristic of fans operating in a given system with constant air density: a. With constant fan size and varying fan speed: (1) Volume (CFM) varies directly as the fan speed. CFM 2 CFM 1
=
RPM 2 RPM 1
(2) Static Pressure Varies directly as the square of the RPM. SP 2 SP 1
RPM 2 = RPM 1
2
3) Horsepower absorbed by the fan varies directly as the cube of the fan speed. HP 2 HP 1
RPM 2 = RPM 1
3
b. With varying fan sizes at the same speed: (1) Volume (CFM) varies directly as the cube of the fan speed.
4-4
CFM 2 CFM 1
DIA 2 = DIA 1
3
(2) Static Pressure Varies directly as the square of the fan speed
DIA 2 = DIA1
SP 2 SP 1
2
(3) Horsepower absorbed by the fan varies directly as the fifth power of size. HP 2 HP 1
DIA = 2 DIA1
5
2. The following relations are characteristic of a fan of a given size delivering a constant mass of air of varying density. (Density varies directly as absolute temperature and inversely as the atmospheric pressure.): a. Volume, fan speed, and total pressure vary inversely as the density. b. Horsepower absorbed by the fan varies inversely as the square of the density. 3. Fan Horsepower varies directly as the product of the volume (ACFM) and the total pressure (inches W.G.) divided by the constant 6370 times the total aerodynamic efficiency. Actual Fan Horsepower =
E.
ACFM x Total Pressure 6370 X Total Aerodynamic Efficiency
Control Valves
1. Pumped Circuit Allocating Pressure Drops to Control Valves : In a pumped circuit, the pressure drop allocated to the control valve should be 33% of all other friction losses in the system at pump rated flow (exclusive of the valve pressure drop itself) or 15 psi whichever is greater. Valid for < 750 GPM & < 150 psi pump delta p Valid for > 300 GPM & 150 to 275 psi pump delta p If outside these ranges, pressure drop allocated may be 25% of system dynamic losses at pump rated head, or 15 psi whichever is greater. In both cases above, use no more than 90% of valve's Cv .
4-5
2. Compressor discharge and suction lines The pressure drop allocated to a control valve in the suction or discharge line of a centrifugal compressor should be 5% of the suction absolute pressure, or 50% of the system dynamic losses (exclusive of the control valve) at the compressor rated point, whichever is larger. Also, no more than 90% of the valve Cv should be used. 3. Pressure motivated systems In a system where tank pressure moves liquid from one vessel to another, the pressure drop should be 10% of the lower terminal vessel pressure, or 50% of the system dynamic losses, whichever is greater. The above rule also applies to vapor, but in addition, the assigned Delta P should not exceed 42% of the upstream pressure to avoid critical flow problems through the valve. 4. Steam and flashing water Valves in steam lines to turbines, reboilers, and process vessels, should be allocated 10% of the design absolute pressure of the system or 5 psi, whichever is greater. The valve should be sized for twice the normal flow rate since steam usage rates can vary widely, especially during start-up.
F.
For valves handling a flashing mixture, the allocated pressure drop should be equal to 0.9 times the difference in between the absolute inlet pressure and the absolute saturation pressure if flowing temperature is more than 5° F below the saturation temperature. If less than 5° F below the saturation temperature, the pressure drop should not be greater than 0.06 times the absolute inlet pressure. Two Phase Flow: For the seven types of two phase flow patterns in pipes, some guidelines on liquid and vapor superficial velocities (LSV and GSV respectively) which can be used to make initial predictions on the type of flow pattern are given below. 1. Horizontal Pipes: a. In DISPERSED FLOW PATTERN, nearly all the liquid is entrained as spray by the gas. This occurs at GSV > 200 ft/sec. b. In ANNULAR FLOW PATTERN , liquid forms a film around the inside wall of pipe and gas flows at a high velocity as a central core. This occurs at GSV > 20 ft.sec. c. In BUBBLE FLOW PATTERN , bubbles of gas move along at about the same velocity as the liquid. This occurs at LSV of 5 to 15 ft/sec, and GSV of 1 to 10 ft/sec.
4-6
d. In STRATIFIED FLOW PATTERN , liquid flows along the bottom of the pipe and gas flows over the smooth liquid gas interface. This normally occurs for LSV< 0.5 ft/sec, and GSV of 2 to 10 ft/sec. e. In WAVE FLOW PATTERN, the interface is disturbed by waves moving in the direction of flow; otherwise it is similar to stratified flow pattern. This occurs for LSV < 1 ft/sec and GSV of about 15 ft/sec. f. In SLUG FLOW PATTERN, waves are picked up periodically in the gas stream and form a slug which moves at much greater velocity than average liquid velocity. Slugs can cause severe vibration due to impact on fittings such as return bends. g. In PLUG FLOW PATTERN , alternate plugs of liquid and gas move along the pipe. This occurs at LSV < 2 ft/sec and GSV < 3 ft/sec. 2. Upflow Vertical Pipes: a. b. c. d. e. f. g.
For dispersed flow, GSV > 70 ft/sec. For annular flow, LSV < 2 ft/sec. and GSV > 30 ft/sec. For bubble flow, GSV < 2 ft/sec. Stratified flow does not occur. For wave flow, the SV’s are unpredictable. For slug flow, GSV = 2 to 30 ft/sec. For plug flow, GSV = 2 to 30 ft/sec.
4-7
FLUID FLOW REFERENCES
Flow of Fluids Through Values, Fittings, and Pipe, Crane, Technical Paper No. 410 E&P General Design Specifications Section VI. B. Piping Hartzell Fan Engineering Data, Bulletin A-108-E, Hartzell Fan Inc. Pigua, OH 45356 “Allocating Pressure Drops to Control Valves” Instrumentation Technology, pg 102105, 10/1970 “Preliminary Pipeline Sizing,” Chemical Engineering, September 25, 1978, pg 119120 “Guidelines on Fluid Flow systems,” Hydrocarbon Processing, April, 1990, pg 47-60 Bechtel Process Engineering methods for limiting and erosion liquid velocity
4-8
5 A.
FRACTIONATION
Minor Components Non Ideal In Hydrocarbons
1. CO2
Behaves with volatility between methane and ethane. If 1% in gas, be sure to check for solidification in NGL expander plant process.
2. H2 O
Behaves with volatility similar to ethane but causes hydrates with light hydrocarbons and CO2 .
3. H2 S
Behaves with volatility similar to propane. Can be removed from liquids by partial depropanization.
4. Methanol Behaves with a volatility similar to propane. Used to inhibit hydrates in wet gas. Excess dissolves in water. In propane refrigeration systems, collects as a separate phase in bottom of the chiller and must be drained out.
B.
5. Ethyl Behaves with a volatility less than propane. Used as stench for odorizing Mercaptan commercial propane. In propane refrigeration systems, collects with the oil in the propane chiller and is removed with the oil. Columns 1. Rules of Thumb on Sequence a. Get rid of troublesome components first. N2 , H2, H2S, CO2 , H2O. b. Take successive cuts, lightest to heaviest. c. Do easy separations first, i.e., where K Lt key/K Hv key is greatest. d. Try to get ½ distillate. (Mole basis) e. Take a finished product as distillate. 2. All light hydrocarbon distillation columns should be controlled with floating pressure. 3. Heuristic for distillation: In a multicomponent mixture, it is usually best to remo ve components one at a time beginning with the lightest. 4. Heuristic for separations: The next separation in a sequence is the one that can be done the cheapest. 5. In large diameter (greater than 2 feet) packed distillation columns, liquid distribution is the key element to separation efficiency. 6. Do not design a process based on an UNCONVERGED simulation! 7. For optimum economy, in terms of minimum column diameter and maximum tray efficiency, distillation columns should be designed with 5 to 10 percent entrainment.
5-1
8. Practical limits for packed columns a. 30 gpm/ft2 liquid rate maximum b. 3 gpm/ft2 liquid rate minimum (These can be extended by proper distributor design) 9. A quick estimate of the number of trays for a given separation – double the minimum number of stages obtained from the Fenske equation:
N min
D l B l Log D h B h = Log α
Where D1 = Mole fraction of light key in distillate Dh = Mole fraction of heavy key in distillate B1 = Mole fraction of light key in bottoms Bh = Mole fraction heavy key in bottoms α = Relative volatility, light key/heavy key The minimum reflux ratio can be calculated by the following equation, if there are no divided keys, i.e., no components whose volatilities lie between the light and heavy keys. Multiply the minimum reflux ratio by 1.3 to correspond to double the minimum stages. R min =
D 1 α (1 − D1 ) − (1 − F ) α − 1 F 1 1 1
Where F1 = Mole fraction light key in Feed If divided keys are present, it is probably easier to run a rigorous simulation than to solve the short cut equations by hand. 10. Sizing reflux accumulators: Generally accumulators are sized based on liquid holdup time from normal liquid level to low liquid level, where NLL is approximately the center of the vessel. Holdup time of 8 to 10 minutes is generally acceptable, based on total overhead condensed. Low liquid level should be 8 to 12 inches above the bottom of the vessel or the height of the vortex breaker, if one is present. Generally reflux accumulators should be 2 feet or greater in diameter.
5-2
For a column with a partial condenser, sufficient vapor space must be left above high level for vapor/liquid separation. This can be roughly estimated for light hydrocarbon systems as: Area vapor, req’d (ft 2 ) = (Vapor flow rate, ft3 /sec) /0.7 This vapor cross sectional area should be added to the total cross sectional area of the vessel to obtain the diameter. Even if the condenser is a total condenser, the high liquid level should be 6 to 8 inches below the top of the vessel. 11. Estimating column diameter (low pressure columns only). The active area of a trayed column can be fairly accurately estimated. The downcomer areas can also be reasonable estimated. (Downcomer area should never be less than 5% of the total area of the column.) Therefore a rough estimate of the diameter can be readily obtained. Area active =
W
1.6
ρv
DCarea (sq.ft.) = Liquid rate (gpm) / 175 Tower area = AreaActive + DCarea Where W = Vapor rate, lb/sec ρv = Vapor density, lb/ft3 12. Fractionation: Reflux to Feed Ratio 40 Tray – 0.55 mol/mol 30 Tray – 0.7 mol/mol 13. Estimating Tray Efficiency The graph below can be used for rough estimates of tray efficiency. Use with Caution.
5-3
5-4
FRACTIONATION REFERENCES FRI Fractionation Tray Design Handbook, Vol 1, 2, & 5 PPCo Data Input Guide for Program 6180 H. L. Walker: Factors to be considered in the Design of Fractionation Equipment. PPCo. 1980. O’Connell, H. E.: Trans. Am. Inst. Chem. Engrs., 42, 741 (1946)
5-5
6
A.
COMBUSTION
Flare 1. Flare tip velocities
Open pipe flare: flare: 0.5 MACH emergency emergency flaring 0.2 MACH continuous flaring Proprietary flare tips: 0.5 – 0.7 MACH (max) 2. Limit flare piping velocities to 0.7 MACH in laterals and to 0.5 MACH in headers. 3. Flare stacks stacks and headers headers must be continuously purged. 4. Flare tips shall be provided provided with gas seals to reduce purge purge gas requirements. requirements. 5. Sound level from flare tip shall not exceed 100db for 15 minutes. 6. Limit flare flare radiation radiation to adjacent flares and other equipment to 1,500 BTU/Hr/Sq.Ft. BTU/Hr/Sq.Ft. 7. Flare pilots shall be monitored. 8. At all times headers headers shall maintain a positive pressure. pressure. 9. No relief/depressurizi relief/depressurizing ng streams containing oxygen shall be tied tied into flare flare header. 10. Laterals and headers shall not have low places to collect liquids & corrosion products. 11. Laterals to enter headers at top and at 45 degree angle in direction of header flow. 12. Slope laterals to drain to headers. 13. PSV bypass shall have a minimum opening equal to the PSV orifice, no less than 2”. 14. Normally limit PSV back pressure to: 10% Conventional Relief Valves 30% Bellow Relief Valves 15. Estimate final temperature of gas remaining inside vessel following depressurization to flare from mid-point between resulting temperature calculated by isenthalphic (“JT”) and isentropic (“turboexpander”) expansion.
B.
16. Flare, reference API RP 521. Fired Heaters Heaters For most process heaters, assume a thermal efficiency of 75 to 80% when calculating fuel requirements. Where: % Thermal Efficiency = Heat Transferred x 100 Heat Released
6-1
C.
Fuel Requirements For determining fuel requirements for process equipment, always use the net heating value of the fuel rather than the gross heating value.
6-2
COMBUSTION REFERENCES
Furnace operations by Robert D. Reed, Gulf Publishing Co. Copyright 1973 Steam/Its Generation, Babcock & Wilcox 161 East 42nd Street, New York, NY 10017, Copyright-1972, thirty-eighth edition Consult API RP t521 “Guide For Pressure-Relieving and Depressuring System” For Flare Sizing Calculations
6-3
7 A.
PHYSICAL PROPERTIES
Standard Conditions
Standard conditions for the U.S. refining industry are 1 atmosphere and 60F for both gases and liquids. Other standards apply at different international sites. At standard conditions, 1 lb-mole of gas occupies 379 cubic feet B.
Characterization of Liquid Refinery Feeds and Products
Refinery feeds and products can be described by o API , Characterization Factor, K, and a distillation curve such as D-86 or D-2887. 1.
o
API gravity is related to specific gravity by the following equation: API =
O
141.5 sp. gr .
− 131.5
The API gravity of water is 10. 2. K or characterization factor is a measure how aromatic, naphthenic or paraffinic the liquid is. K is related to boiling point and specific gravity by: K =
3
T B
sp. gr .
Paraffinic stocks Intermediate stocks Naphthenic stocks Aromatic Residues
where TB = average boiling point at 1 atm, o Rankine K 12.2 – 12.5 11.5 – 12.1 11.0 – 11.4 9.8 - 10.9
Figures are available that relate K to molecular weight, specific heat, heat of vaporization, viscosity, and critical properties. 3.
Distillation Curves Distillation curves are obtained by distilling a standardized volume of sample a defined rated. The distillation curve is a plot of the c umulative volume of the condensed distillate vs. the vapor temperature. GC methods are also available the produce distillation curves. ASTM methods include: D-86, D-1160, D-2887, D-3710. The D-86 method is the simplest and most common. A D-86 curve for distillate is below. The D-2887 and D-3710 methods are GC methods and commonly referred to as simulated distillations or a SimDist.
7-1
D-86 Curve for Distillate 600 F , 550 e r u t a r 500 e p m e T 450 r o p a 400 V
350 0
10
20
30
40
50
60
70
80
90
100
Volume Distilled, %
C.
Physical Properties of Selected Liquids
Compound
Properties of Selected Liquids Specific Heat, Heat of Normal Boiling Btu/lb-F @ 60F Vaporization, Point, F Btu/lb @ NBP
Density, lb/gal @ 60F
Viscosity, cP @60F
IsoButane
0.57
158
11
4.70
0.2
n-Butane IsoPentane
0.57 0.54
166 148
31 82
4.88 5.22
0.2 0.2
n-Pentane
0.54
145
97
5.27
0.2
Octane Decane
0.53 0.52
131 120
258 345
5.90 6.20
0.6 1.0
Benzene Toluene
0.41 0.40
191 156
176 231
7.36 7.28
0.7 0.6
Gasoline Diesel
0.50 0.47
125 105
100-440 378-670
6.32 7.18
0.6 4.0
Arabian Extra Light Arabian Light
0.45 0.44
-------
570 (50% pt.) 611 (50%pt.)
7.00 7.18
4.6 9.2
Arabian Heavy
0.44
----
733 (50% pt.)
7.43
38.4
Methanol
0.59
472
148
6.66
0.6
Ethanol
0.57
364
173
6.63
1.3
Water
1.00
974
212
8.32
1.7
1. Specific Heats of Hydrocarbon Liquids
7-2
2. Latent Heat of Vaporization at Atmospheric Pressure
7-3
3. Specific Gravity
7-4
4. Thermal Conductivity
7-5
5. Viscosity Kinematic viscosity, ν , is measured in centistokes, and is equal to viscosity in cP divided by specific gravity or ν
= µ sp. gr
7-6
6. Vapor Pressure
7-7
D.
Physical Properties of Selected Gases/Vapors Specific Heats for Selected Gases/Vapors
Compound
Specific Specific Density, Thermal Viscosity, cP Heat, Btu/lb- Heat, Btu/lb- lb/cuft @ Conductivity @ 212F, F @ 60F F @ 200F 1 atm, 60F , Btu/hr-ft-f 1 atm @ 212F, 1 atm
Hydrogen
3.404
3.451
0.005
0.129
0.0104
Methane
0.526
0.578
0.042
0.0262
0.0135
Ethane
0.407
0.487
0.079
0.0175
0.0115
Propane
0.281
0.305
0.116
0.0151
0.0102
Butane
0.293
0.316
0.153
0.0135
0.0093
Hydrogen Sulfide
0.239
0.246
0.090
0.0110
0.0157
Air
0.239
0.239
0.076
0.0183
0.0218
Nitrogen
0.248
0.248
0.074
0.018
0.0210
Oxygen
0.217
0.217
0.084
0.0185
0.0246
Carbon Monoxide
0.249
0.250
0.074
0.0174
0.0209
Carbon Dioxide Water Vapor
0.200 0.445
0.219 0.451
0.116 0.047
0.0133 0.0137
0.0184 0.0125
Helium
1.647
1.647
0.011
0.0104
0.0232
7-8
1. Specific Heat of Gases and Vapors
7-9
2. Thermal Conductivity
7-10
7-11
3. Viscosity
7-12
E.
Physical Properties Recommendations
Experience has shown that the recommended choice of physical property models depends upon the simulator being used, process being modeled and the process modeling objectives. Each situation has unique conditions that may or may not require additional analysis. For crude still and refinery applications, the Chao-Seader method with Grayson Streed modifications is a good first choice. This method will adequately model the typical light gas processes in the refinery as well. In Aspen, this method is known as “GRAYSON” and should be modified to have the free water with solubility option for unsaturated systems selected (method 2 on the Properties specification form). This method is not accurate for streams containing ammonia and/or hydrogen sulfide. For gas plant or LNG processes either the Soave-Redlich-Kwong or the PengRobinson equation of state should be used. When using Hysys, the Peng-Robinson method is recommended since the Hyprotech property expertise has been focused on this method. When using Aspen, you should select the RKS-BM method, (RedlichKwong equation of state with Boston-Mathias modifications), that is included in the PPCO NGL Processing template. The template contains modifications to the property route that improve the liquid densities and other transport properties as well as make the method work better for hydrogen containing streams. These methods should not be used for streams containing Helium and or high concentrations of hydrogen without obtaining special modifications from a Phillips property specialist (e.g. Howard Wilson). For sour gas or sour water treatment there are special physical property packages available. The BR&E Tsweet amine property method works well for most gas treatment modeling. The Amsim package in Hysys can also be used for gas treatment. These amine treatment packages have a limited range of application, but are generally fairly easy to use and are fairly accurate.
A special electrolyte modeling package such as the sour water options in Hysys, the OLI Prochem software or the Aspen electrolytes package should be used to model sour water treatment. These packages are complicated and should be used with guidance from a physical property specialist. Difficult or tight separations require close scrutiny by a physical property specialist. These separations are generally characterized by cases where either very high purity is required when the relative volatility of the primary keys is between 0.8 and 1.2. Standard methods have to be specially adjusted or “tuned” to data near the pinch points in order to get a reasonable fit and a good match with equipment. Hydrates and solids formation can be modeled with some other special purpose software. Gas hydrates are generally best estimated using the hydrate program within PVTSIM and the Multiflash program from InfoChem. Bill Parrish is a world-class expert in gas hydrates and should be consulted anytime they are of concern. Different
7-13
F.
programs are to be used for different types of solids prediction. For cryogenic hydrocarbon systems, the multiphase flash program from DTH is currently the best option, the GPA Kohn-Luks program is also usually a good option for these systems. Dale Embry is Phillips expert on these programs. The OLI software is the best available for prediction of chemical salt and hydrate formation. Contact Dale Embry for more information about this program. Separations involving other chemicals or other thermodynamic properties should be reviewed by a physical property and/or separations specialist before any designs are finalized. Simulation Techniques for Characterization of Oils Most of the time, data about oils are provided as some sort of curve. However, most simulation programs are designed to work with pure components that have one value for a property. So we take portions of the curve and treat it as though it were a pure component. These fractions are commonly referred to as "pseudocomponents". The number of pseudocomponents needed for accurate simulations is as much a function of the process being simulated as the type of oil being characterized. The number of pseudocomponents needed for an accurate simulation increases as the tightness of separation increases and the number of unique products increases. Narrow cuts, cuts having a boiling range of 25 to 50 F (10C to 25C) are needed at the boundaries of the unique products, larger cuts can be used further away from the product boundaries. For most field gas/oil separation calculations there are between one and three relatively broad separations being performed in the butane to octane boiling range. These simulations are well characterized by using standard components through hexane and then using between 4 to 10 pseudocomponents with increasingly broad boiling ranges. Often 5 cuts with the ranges of C7, C8, C9- C10, C11-C14, C15+ will adequately model the separator performance and the oil properties. For refinery simulations, as many as 40- 50 pseudocomponents may be needed to characterize a whole crude as it is being split into its various products. One often uses 25F cuts for the 200-800 F boiling range materials, 50F cuts for the 800-1000F boiling range materials and 100F cuts for the 1000+ materials. If only single phase properties are needed, it is very likely that a single pseudocomponent with a 100 to 200F boiling range can be used for a stream.
7-14
PHYSICAL PROPERTIES REFERENCE
“Engineering Data Book”, GPSA, Tulsa, Vol II, 1987 J. M. Campbell, “Gas Conditioning and Processing”, Norman, Vol 2, 1981 Phillips Engineering Standards 7.00 – 7.16 “Technical Data Book – Petroleum Refining”, API, Vol 1-3 (Chapter 2, Vol 1 contains procedures for pseudo compound characterization) K. B. Maxwell, “Data Book on Hydrocarbons: Applications to Process Engineering”, Van Nostrand New York, 1950 W. L. Nelson, “Petroleum Refinery Engineering”, McGraw-Hill New York, 1949
7-15
8 A.
COMPRESSORS, EXPANDERS & PUMPS
Reciprocating Compressors
1. Limit compression ratio to get a maximum temperature of 300 ° F discharge temperature. This is generally less than a compression ration of 3.7 per stage. 2. Horsepower required is usually a maximum at a compression ratio of 2.0. 3. For reciprocating propane compressor calculations, add 10% to the final Horsepower calculated by conventional means and 10 °F to the final discharge temperature for preliminary design. 4. To avoid excessive vibration, the mass of the foundation must be approximately 5 times the mass of the unit. B.
Compressor Quickies
1. 1 lb-mole (ideal) gas occupies 379 SCF. Thus mass flow in lb/min = MMSCFD*MW/(1440*379). 2. Least flow through a centrifugal compressor is 175 ACFM (300 m3/hr) discharge volume. ACFM = MMSCFD*106 *14.7*T*Z/ (1400*P*520) (T-temperature in °R) To achieve a reasonable compressor efficiency with a ce ntrifugal compressor, the suction acfm needs to be above 1500 to 2000. Avoid applications at lower acfm. 3. Centrifugal compressor head −1 Z ∗ 1545 ∗T S P d γ 1 H a= ∗ ∗ − γ − 1 P s MW γ
γ
Maximum head per impeller is 10,000 ft. 4. Discharge temperature γ
P Reciprocating T d =T s∗ d P s Centrifugal
P T d =T s∗ d P s
−1
γ
−1 ∗
γ
γ η p
η p =
Polytropic efficiency
8-1
γ =
C p /Cv
5. Maximum allowable discharge temperature for associated gas (i.e. gas from crude oil wellhead separation) compression is 300 °F (150 °C). Recommended maximum discharge temperature for centrifugal compressor is 350 °F. Absolute maximum discharge temperature for centrifugal compressor is 400 °F. High temperature seals are absolutely necessary to operate at the absolute maximum temperature. 6. “Head” for a centrifugal compressor is really energy imported to the gas. “Feet” head is in fact ft-lb force/ft-lb mass. SI expresses it as kiloJoule/kilogram. Conversion: 1000 ft- lbf /ft-lbm = 2.989 kJ/kg Other expressions such as “meters” head (converting feet into meters) are meaningless as they do not take the gravitational constant into account.
C.
7. For centrifugal compressors, a 20% surge margin from the design operating point is recommended, 30% is preferred. Absolute minimum acceptable is 10%. Liquefied Natural Gas (LNG) Plants
D.
For centrifugal compressor calculations, assume a polytropic efficiency of 79 to 82 % for LNG plant preliminary design. Energy Conservation Natural Gas Engines
E.
If compression ratio is increased from 8:1 to 10:1 a 5% reduction in fuel rate and a 9% increase in brake horsepower results (assumes 1000 Btu fuel & minimum octane of 115). Octane Number: C1-120 iC4-97.6 nC5-80.2 C2-100.7 nC4-89.1 C6-26.0 C3-98.1 iC5-61.9 C7-0.0 Fuel consumption
F.
Reciprocating Engines – 1000 hp ≈ 7200 Btu/hr/bhp turbines – Solar T4500 ≈ 9800 Btu/hr/bhp NGL Expander Plants 1.
For expander-compressor preliminary designs, limit the adiabatic efficiency to 72 % for the expander end and 65 % for the compressor end. Vendors will claim higher efficiencies, but actual tests indicate the above efficiencies to be more representative in service. Use a 95 % mechanical efficiency for transferring horsepower from the expander to the compressor.
2.
For expanders, limit liquefication of feed stream through expander to 15 to 18 weight % maximum. In excess of this, mechanical problems may result with expander.
8-2
G.
Gas Processing – Simulation guidelines
1. Turboexpander maximum size approx. 8,000 hp (6 MW). 2. Turboexpander maximum pressure ratio 2.5:1. 3. Temperature reduction on expansion (° F/psi):
Lean gas Rich gas H.
Turboexpander 0.06 0.1
JT-valve 0.03 0.05
Pump sizing
Example – Size The Lean Oil Pump Suction Conditions – 145 psig @ 90° F Discharge = 485 psig Capacity = 700,000 GPD @ 60° F SP. GR. @ Pump Temp. = 0.805 700,000 GPM @ Pump Temp. = 1440 min/Day = 486
Head
=
0.815 0.805
.815 .805
∆ P x 2.31
= (485-145)2.31 = 976 Ft. of Liquid SP. GR. = 0.805
B.H.P. = (GPM) (Ft of Hd.) SP. GR.) (3960) x Eff. B.H.P. = (492) (976) (0.805) = 139.5 H.P. (3960) (.70) Use 150 H.P. electric Motor Drive. Check Curve of Pump Purchased for End of Curve H.P. Requirement. Specify Pump for 4.92 GPM @ 976 Ft. Head. Case to be good for 485# Disch. and shut in head of pump. NPSH Not Critical. H.P. x 33,000 = Ft. - LB Min. LB. H2 O @ 62° F x 0.12 = GAL. H.P. x 33,000 x 0.12 = 3960 x H.P. = FT. - GAL. Min.
8-3
I.
Pumps
1. Suction Specific Speed: When selecting centrifugal pumps, the suction specific speed (NSS) for the pump should be le ss than 11,000. Experience has indicated that pumps operating with suction specific speeds above 11,000 have a much higher failure frequency. N SS =
RPMx GPM
( NPSHR )0.75
Where: N SS = Suction specific Speed GPM = Flow Rate in gallons per minute NPSHR = Net positive suction head required for the pump in feet of fluid. If the pump is double suctioned, divide flow rate by two. 2. Net Positive suction Head: Be sure to check the size and configuration of the pump suction and piping before increasing the pump speed as it affects the pump required Net Positive Suction Head (NPSHR) considerably. See the relation: NPSHR 2 NPSHR 1
n = 2 n1
2
Where: n is pump speed in rpm and subscripts 1 & 2 indicate initial and final conditions. 3. Axial Compressors 1. Axial compressors cannot be used with side loads. 2. Maximum discharge pressure is around 300 psia.
3. The minimum suction actual cubic feet per minute is 60,000. 4. Axial compressors cannot operate over as large a range of flow as centrifugal compressors.
5. The maximum recommended discharge temperature is similar to a centrifugal or about 350 °F.
8-4
J.
General: 1. Maintenance costs for gas engine driven reciprocating compressor units is about 6 times those of gas turbine driven centrifugal compressor units.
For 1993: Reciprocating = $30/HP Yr Centrifugal
= $ 5/HP Yr
2. Gas Turbine Wasteheat Available About 1/3 of fuel gas BTU requirements can be considered recoverable as wasteheat from a gas turbine for quickly preliminary estimates.
8-5
COMPRESSORS, EXPANDERS & PUMPS REFERENCE
GPSA Engineering Data Book, Section 13 “Centrifugal Pump Seminar” by C. C. Fletcher 10.92 “Guidelines on Fluid Flow Systems”, Hydrocarbon Processing, April, 1990, pg 47- 60 C. C. Fletcher ROT advice for process engineers – 6/93 & Nils Nilsen – Pignone contact July 1993 C. C. Fletcher wasteheat rule of thumb
8-6
9 A.
REFRIGERATION
Condensers
For water cooling with propane condensers, try to maintain no greater than a 10 °F approach to the warm cooling water leaving the condenser. Use a velocity of 4 to 8 ft/sec for the water through the tubes. Restrict cooling water return temperatures to a maximum of 125 °F. B.
Propane Refrigeration Systems
1. For design of propane refrigeration systems, suggest using a composition of 3.0 mol % C2, 95 mol % C3 and 2 Mol % C4. Experience has indicated it is difficult and expensive due to propane losses to maintain 99 % + propane content. The utilization of continuous purge systems will typically result in high losses. 2. Air, which is pulled in around piston rod packing with low suction pressure of 1 to 2 psig and high valve losses) tends to accumulate in the propane accumulator vessel vapor space after the condensers. This leads to high compressor discharge pressure and a potentially hazardous situation. A manual purge of the accumulator vapor space weekly will keep the air concentration down towards acceptable levels. 3. For water cooled propane condenser design, generally use 10 °F temperature rise on cooling water through exchanger and a 10 °F approach of condensed propane out to the warm water from the exchanger.
C.
4. For preliminary design of reciprocating propane compressor calculations, add 10% to the final Horsepower calculated by conventional means and 10 °F to the final stage discharge temperature. Gas Processing – Simulation Guidelines
D.
1. Refrigerant chiller temperature approach to gas is normally 5 to 10 °F. Condensing Temperature Effects: For gas turbine driven propane refrigeration systems, there will be approximately ½ % to 1% increase in gas load horsepower for every degree Fahrenheit increase in refrigerant condensing temperature.
9-1
REFRIGERATION REFERENCE
GPSA Engineering Data Book, Section 14 Bechtel Process Engineers rule of thumb and CE Process checks
9-2
10
A.
MISCELLANEOUS
Large Production & Processing Platforms
1. Specific Process Considerations The following process considerations are relevant to the platform layout: a. the layout should reflect a logical progression of the process. This should then minimize major piping requirements. Other equipment with small diameter piping, can be located where necessary e.g. glycol regeneration package. b. Related equipment should be modularized to allow maximum hook up and precommissioning to be achieved, e.g. compressor with associated knock out drum and aftercooler. c. Requirements for pump NPSH should be established e.g. export pumps. d. Pump suction lines should have no vapor pockets and be of minimum length. e. Systems based on gravity drainage should be identified. f. Requirements for future equipment should be identified. to allow incorporation into initial layout. g. Allowances for straight piping runs should be made where appropriate e.g. compressor suction lines, metering runs. h. Compressor suction lines should have no liquid pockets. i. Two phase lines should be of minimum length and with no pockets to minimize potential slugging. j. Two phase lines from coolers should free drain to the knock out drum. k. Vapor lines in wet Carbon Dioxide or corrosive service should have no liquid pockets. l. Filter separator/KO drums located upstream of TEG contacting/acid gas treating units should be located close to prevent hydrocarbons or millscale from entering the contactor. m. Flare lines should slope to the KO drum with no liquid pockets. n. Locate control valves in bubble point liquid service so that there is no possibility of flashing at the valve inlet. 2. Preliminary Space Estimate On the assumption that a modularized concept is used, the following factors can be used to assess deck area requirements for preliminary layout studies. 3. Seawater System: A. Carbon Steel Service Although Carbon Steel is fine for service in low velocity seawater at ambient temperatures, the corrosion rate increases rapidly with temperature and agitation.
10-1
At 122 ° F (50 ° C) and above, the corrosion rate in agitated seawater is greater than 50 mils/year. At velocities over 15 ft/sec, turbulence may greatly accelerate the corrosion rate by eroding away the protective film. This occurs frequently at heat exchanger tube inlets, U bends, and p iping elbows. At more extreme velocities, erosion – corrosion will occur. At 39 ft/sec, the corrosion rate reaches over 200 mils, year. B. For Seawater Waterflood systems: Typical specifications to prevent formation plugging with solids, downhole corrosion, and bacteria growth are as follows. 1. Solids – Remove 97% of all solids greater than 5 microns size. 2. Oxygen – Remove to a maximum level of 10 ppb(wt) 3. Sterilization – Chlorination and/or UV sterilization required. 4. Portable Water systems: a. Based on Ekofisk experience, provide the following potable water quantities for preliminary design: 1.
200 to 250 liters (53 to 66 gallons) per day per person for personnel who live on the platform.
2.
100 to 150 liters (26 to 40 gallons) per day per person for day workers and visitors who do not live on the platform.
b. Based on Ekofisk experience, provide the following potable water storage preliminary design:
1.
For platforms with potable water makers, provide a minimum of 2 days storage.
2.
For platforms which depend solely on hauled potable water, provide a minimum of 3 days storage.
10-2
Module type Wellhead
B.
Area Ratio* Set by well pattern/drilling requirments Separation 0.20 – 0.40 Gas Compression 0.15 – 0.20 Water Injection 0.20 – 0.25 PowerGeneration 0.20 – 0.30 Utilities 0.15 – 0.20 *Ratio of Major Equipment footprint area to total module area. Water and Steam Systems 1. Approximate break point for steam pressure at which silica becomes a problem with vaporization and deposition on turbine blades is at 500 psig. 2. The evaporation rate on a cooling tower is dependent on the amount of water being cooled and temperature differential. For each 10 ° F temperature drop across the tower, 1 % of the recirculation rate is evaporated. In other words, 0.001 times the circulation rate in gpm times the temperature drop equals the evaporation rate is gpm.
C.
Economics
1. Capex Ratio Exponents For Processing Plants and Ancillaries Same No. of Units COST 2=(SIZE OR CAPACITY 2/SIZE OR CAPACITY 1)0.5 (COST1) Unit number change required for new capacity. USE 0.6 exponent Infrastructure (Camps, Warehouses, Maintenance facilities) USE 0.3 exponent 2. Capex Factors From Major Equipment Cost Installed Cost Onshore 2.5 x (Major equipment cost) Offshore 5.0 x (Major equipment cost) (excludes deck and jacket costs) 3. Annual Operating Costs [excludes fuel and depreciation] Onshore; 3% of Capital cost Offshore; 5% of Capital Cost 4.
Capex (Total) Remote Area LNG Plants
10-3
D.
[One Train} $ 2x109 per 2x106 MTY [1990 BASIS] Note: Cost Reductions via technology offset regulatory increases. Use 5%/Yr esc. in general costs. MTY = Metric Tonnes/Year Hydrates 1. Hydrates generally form at 50 to 60 ° F. 2. Expect 1 ° F depression for each 1% methanol in liquid drainage. (Methanol content may be estimated with a hydrometer) 3. Hydrate Control Add a margin of 50% to calculate hydrate inhibitor injection rates. 4. Typical hydrate inhibitor concentrations: MEG: 70 – 80 wt% at inlet and 60 wt% in solution outlet MeOH: 98 wt% at inlet and 90 wt% in solution outlet 5. Glycol inhibitor loss estimate is 1 lb/MMSCF plus 200 ppm (v) in liquid hydrocarbon.
E.
6. MeOH will melt hydrates already formed. MEG will not. NGL Expander Plants If the CO2 in the feed gas to the cryogenic plant is in excess of 0.25 mol %, be sure to check for CO2 solidification in both the liquid and vapor phases immediately downstream of the expander and in the top four stages of the demethanizer.
F.
Miscellaneous Plant Systems
G.
Instrument air – As a preliminary estimate for instrument air requirements for feasibility study design, use 0.5 to 0.75 scfm per control instrument. Liquified Natural Gas (LNG) Plants 1. For preliminary estimates of LNG plant design inlet volume for a premised LNG delivery to ships for transport, use a 93 plant availability factor.
H.
2. Mercury occurring naturally in some natural gas streams is extremely corrosive to aluminum heat exchangers used extensively LNG plants processes. Plan to check for mercury in feed gas up front in any project. Gas Processing – Simulation Guidelines
I.
Include a 5 ° F margin from the sales contract hydrocarbon dewpoint temperature to accommodate uncertainties in process simulator to predict dewpoint. Offshore Pipeline Gas Specifications
10-4
J.
1.
Hydrocarbon dewpoint temperature sho uld be 10 – 20 ° F below minimum operating temperature in pipeline for the operating pressure range to prevent liquid drop out.
2.
Water dewpoint temperature should be 10 – 20 ° F below minimum operating temperature in pipeline for the operating pressure range to prevent free water drop out.
3.
Maximum allowable platform discharge temperature is typically 90 – 125 ° F.
Offshore Crude Oil Specifications
1. Max RVP 10 psia for export to tanker/storage. 2. Salt content 70 to 200 mg/1 (sales spec to pipeline) 3. Water max 2 wt% (sales spec to pipeline)
K.
4. BS&W 0.5 vol% maximum ; 0.1 vol% average (sales spec to pipeline) Wind Loadings
L.
Flat Surfaces P=0.004 x V2 Cylindrical Members P = 0.6 x 0.004 x V2 Where: P = Pressure in PSF V = Velocity of wind MPH Steam Leaks @ 100 psi
M.
Orfice Steam Size Wasted/mo. 1 /2 ” 835000 3 /8 ” 470000 1 /4 ” 210000 1 /8 ” 52500 1 /16” 13200 1 /32” 3400 Composition of Air
N.
Air Composition ~ N2 – 78.07 mol% 02 – 20.99 mol% Ar – 0.94 mol% Platform Deflection
Energy Loss / Gas (scf) #2 878347 494737 221053 55263 13895 3579
Month Fuel oil 7383 4155 1857 464 117 30
Maximum deflection of platform due to wind and sea is: deflection = Span / 400 Where: span = distance from cellar deck to mud line, Ft.
10-5
O.
Kinetics
P.
For a second order reaction in a constant volume reactor, if 95% of the reactants react in one time, it will require 20 more time units to react 95% of the remaining reactants. Storage, Vessel Capacity
Q.
A. Vessel Capacity: Capacity (gallons)=(Diameter, ft)2 +2) x Length, inches. Pipeline Volume:
R.
(Diameter of pipe, inches)2 = Barrels/1000 feet (Results are approximately 3% high) Pressure Vessels In general, the maximum operating pressure (MOP) of a process pressure vessel is established from the maximum internal or external pressure at which the vessel operates while fulfilling it’s normal function. In general, for vessels subject to internal pressure only, the maximum allowable working pressure (MAWP) of a process vessel can be arrived at by adding the greater of 10% or 10 to 25 psi to the maximum operating pressure.
S.
T.
Economic L/D ratios for pressure vessels generally fall in the 2 to 5 range where L = shell seam length and D = inside diameter, both in feet. NACE Requirements “Material shall be selected to be resistant to Sulfide Stress Cracking (SSC) or the environment should be controlled if the gas being handled is at a total pressure of 65 ps ia or greater and if the partial pressure of H2 S in the gas is greater than 0.05 psia." Pressure Waves (e.g. water hammer) magnitude of PW in lbs f/in2 psi PW = a x d x vd/(144 x g) a = velocity of sound in the fluid fps d = density of fluid, lb mass/ft 3 vd = velocity decrease, i.e. velocity before change less the velocity after change, FPS g. = conversion factor = standard acceleration due to gravity, 32.174 fps2
10-6
U.
Insulation Types Thermal Conductivity (Btu/ (( hr ft 2 °F)/in)
Surface burning characteristics per ASTM E-84 Temperature
Insulating Material
-100
Temperature 200 500 Other
range, °F
Flame-spread Index
Fibrous Glass
0.20
0.26
0.50
-
-20 to 850
<25
Mineral Wool
-
0.35
0.52
-
60 to 1,900
Noncombustible 0 Noncombustible <5
Cellular Glass
0.27
0.42+
Polyisocyanur ate
-
-
Polystyrene
0.18
-
Polyurethane
0.14
Granular Calcium Silicate
-
Perlit e
-
<50
Water Absorbed (submersion), % by Volume 65 (high)
70 (high) 0 0
-
-290 to 1,200
0.14 at 75 °F
-290 to 300
May be a fire risk <20
-
0.23 at 75 °F
- 40 to 275
Combustible, although sometimes self-extinguishing
- (low)
0.25
-
-
-100 to 220
May be a fire risk
1.6 (low)
0.42+
0.55
-
60 to 1,500
Noncombustible 0
0
Noncombustible 0
0
0.47+
0.65
Smokedeveloped Index
0.58
-
60 to 1,500
0
0.2 (considered none)
0.7 (low)
75 (high)
16 (medium)
° +Manufactures’ data
V.
Absolute Pressure of Atmosphere at Height ‘H’ feet above Sea Level
P = P1 (1 – 0.00000687H)5.256 P1 = pressure at sea level – pisa Density: W = W1 (1-0.00000687H)4.256 W1 = density of air at sea level
10-7
Comments Must have vapor barrier in cold service, since it is water absorbent. General material. Good workability Deteriorates in a alkaline solution (e.g. 10% NaOH). Poor abrasion resistance. Significant deterioration in acids and in various organics. Generally, is replacement for Excellent workability at low temperatures (below freezing point of water) Significant deterioration in acids and various organics, but resists water and vapors. Excellent workability. High water absorbency, but will dry out. Good workability Good workability
MISCELLANEOUS REFERENCES
Process Design Handbook, Section 1.6.6, 1.6.7, Table 1.6.1 Pages 4 of 7, 5 of 7, 6 of 7. GPSA Engineering Data Book, Section 13 J.M. Campbell, “Gas Conditioning and Processing”, Norman “Pressure Vessel Design For Process Engineers” Hydrocarbon Processing, May 1979, pg 181-191 “How to Select Materials”, Chemical Engineering, Nov 3, 1980, pg 86 to 131 PPCoN Process Personnel NACE MR0175 Std Ma’l. Reg.
10-8
